Unit Operations Solutions

Unit Operations Solutions - Exercise 1.1 (i) Known:...

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Unformatted text preview: Exercise 1.1 (i) Known: Fluorocarbons can be produced from the reaction of carbon tetrachloride and hydrogen fluoride followed by a number of separation steps. Given: Flow diagram and a brief description of a fluorocarbons production process in "Chemical Process Industries" , 4th edition, by Shreve and Brink and also in "Shreve's Chemical Process Industries", 5th edition, by G. T. Austin, pages 353-355 (Fig. 20.4). Find: Draw a process flow diagram and describe the process. Description of Process: Two main reactions occur: CCl4 + HF CCl3F +HF CCl3F + HCl CCl2F2 + HCl Excess carbon tetrachloride is reacted in R1 with HF in the presence of antimony pentoxide catalyst and a small amount of chlorine to maintain catalyst activity. The HF contains a small amount of water as an impurity. The effluent from the R1 is HCl, CCl3F, CCl2F2, unreacted CCl4, and small amounts of water and chlorine. The normal boiling points in oC of these components in the order of decreasing volatility are: HCl Cl2 CCl2F2 CCl3F CCl4 H2O -84.8 -33.8 -29.8 23.7 76.7 100 Exercise 1.1 (i) (continued) The reactor effluent is distilled in D1 to remove the CCl4 as bottoms, which is recycled to R1. The distillate enters absorber A1, where HCl is absorbed by water to produce a byproduct of aqueous HCl. The gas from A1 contains residual HCl, which is neutralized, and chlorine, which is absorbed, by aqueous NaOH, in A2. The effluent liquid from A2 is waste. Moisture is removed from the gas leaving A2 by absorption with H2SO4 in A3. The exit liquid from A3 is also waste. The gas leaving A3 is distilled in D2 to obtain CCl2F2 as a distillate, which is then dried in S1 by adsorption with activated alumina. Bottoms from D2 is distilled in D3 to recover a distillate of CCl3F, which is dried with activated alumina in S2. Bottoms from D3, containing residual CCl4, is recycled to reactor R1. Exercise 1.2 Subject: Mixing is spontaneous, but separation is not. Given: Thermodynamic principles. Find: Explanation for why mixing and separation are different. Analysis: Mixing is a natural, spontaneous process. It may take time, but concentrations of components in a single fluid phase will tend to become uniform, with an increase in entropy. By the second law of thermodynamics, a natural process tends to randomness. The separation of a mixture does not occur naturally or spontaneously. Energy is required to separate the different molecular species. Exercise 1.3 Subject: Separation of a mixture requires a transfer of energy to it or the degradation of its energy. Given: The first and second laws of thermodynamics. Find: Explain why the separation of a mixture requires energy. Analysis: As an example, consider the isothermal minimum (reversible) work of separation of an ideal binary gas mixture. Therefore, the change in enthalpy is zero. However, there is a change in entropy, determined as follows. From a chemical engineering thermodynamics textbook or Table 2.1, Eq. (4): Wmin = out n(h - T0 s) - in n(h - T0 s) yi ,k ln yi , k - in = RT0 out nk nj yi , j ln yi , j It can be shown that regardless of values of y between 0 and 1, that Wmin is always positive. This minimum work is independent of the process. Exercise 1.4 Subject : Use of an ESA or an MSA to make a separations. Given: Differences between an ESA and an MSA. Find: State the advantages and disadvantages of ESA and MSA. Analysis: With an MSA, an additional separator is needed to recover the MSA. Also, some MSA will be lost, necessitating the need for MSA makeup. If the MSA is incompletely recovered, a small amount of contamination may result. The use of an MSA can make possible a separation that can not be carried out with an ESA. An ESA separation is easier to design. Exercise 1.5(b) Subject: Producing ethers from olefins and alcohols. Given: Process flow diagram and for production of methyl tert-butyl ether (MTBE). Find: List the separation operations. Analysis: The reactor effluent contains 1-butene, isobutane, n-butane, methanol, and MTBE. The separation steps are as follows: Separation Step Distillation L/L Extraction with water Distillation Products Distillate: all C4s, methanol Bottoms: 99 wt% MTBE Extract: methanol, water Raffinate: C4s Distillate: methanol to recycle Bottoms: water for recycle Exercise 1.6(c) Subject: Conversion of propylene to butene-2s. Given: Process flow diagram and for production of butene-2s. Find: List the separation operations. Analysis: The reactor effluent contains ethylene, propylene, propane, butene-2s, and C5+. The separation steps are as follows: Separation Step Distillation Products Distillate: ethylene Sidestream: propylene, propane Bottoms: butene-2s, C5+ Distillate: propylene Bottoms: propane Distillate: butene-2s Bottoms: C5+ Distillation Distillation Exercise 1.7 Subject: Use of osmosis for separating a chemical mixture. Given: The definition of osmosis. Find: Explain why osmosis can not be used for separating a mixture. Analysis: Osmosis is the transfer of a solvent through a membrane into a mixture of solvent and solute. Thus, it is a mixing process, not a separation process. Exercise 1.8 Subject: Osmotic pressure for the separation of water from sea water by reverse osmosis with a membrane. Given: Sea water containing 0.035 g of salt/cm3 of sea water on one side of a membrane Molecular weight of the salt = 31.5 Temperature = 298 K Pure water on the other side of a membrane Find: Minimum required pressure difference in kPa across the membrane Analysis: The minimum pressure difference across the membrane is equal to the osmotic pressure of the sea water, since the osmotic pressure of pure water on the other side is zero. The equation given for osmotic pressure is =RTc/M. R = 8.314 kPa-m3/kmol-K T = 298 K c = 0.035 g/cm3 = 35 kg/m3 M = 31.5 kg/kmol Minimum pressure difference across a membrane = ( 8.314 )( 298)( 35 ) = 2, 750 kPa = 31.5 Exercise 1.9 Subject: Use of a liquid membrane to separate the components of a gas mixture Given: A liquid membrane of ethylenediaminetetraacetic acid, maintained between two sets of microporous, hydrophobic hollow fibers, packed into a cell, for removing sulfur dioxide and nitrogen oxides from flue gas. Required: A sketch of the membrane device. Analysis: A sweep fluid is generally required. In some cases, a vacuum could be pulled on the permeate side. The membrane device is shown below. Exercise 1.10 Wanted: The differences, if any, between adsorption and gas-solid chromatography. Analysis: Adsorption can be conducted by many techniques including fixed bed, moving bed, slurry, and chromatography. In chromatography, unlike the other adsorption techniques, an eluant is used to carry the mixture through the tube containing the sorbent. Multiple pure products are obtained because of differences in the extent and rate of adsorption, resulting in different residence times in the tube. The tube is made long enough that the residences do not overlap. Exercise 1.11 Wanted: Is it essential in gas-liquid chromatography that the gas flows through the packed tube in plug flow? Analysis: Plug flow is not essential, but it can provide sharper fronts and, therefore, the chromatographic columns can be shorter. Exercise 1.12 Wanted: The reason why most small particles have a negative charge. Analysis: Small particles can pick up a negative charge from collisions in glass ware. In an aqueous solution, inorganic and polar organic particles develop a charge that depends on the pH of the solution. The charge will be negative at high pH values. Exercise 1.13 Wanted: Can a turbulent-flow field be used in field-flow fractionation? Analysis: Field-flow fractionation requires a residence-time distribution of the molecules flowing down the tube. This is provided best by laminar flow. The residencetime distribution with turbulent flow is not nearly as favorable. Turbulent flow would not be practical. Exercise 1.14 Subject: Sequence of three distillation columns in Fig. 1.9 for separating light hydrocarbons. Given: Feed to column C3 is stream 5 in Table 1.5. Alter the separation to produce a distillate containing 95 mol% iC4 at a recovery of 96%. Find: (a) Component flow rates in the distillate and bottoms from column C3. (b) Percent purity of nC4 in the bottoms. (c) Percent recovery of iC4, for 95 mol% iC4 in the distillate, that will maximize the percent purity of nC4 in the bottoms. Assumptions: Because of the relatively sharp separation in column C3 between iC4 and nC4, assume that all propane in the feed appears in the distillate and all C5s appear in the bottoms. Analysis: (a) Isobutane to the distillate = (0.96)(171.1) = 164.3 lbmol/h Total distillate rate = 164.3/0.95 = 172.9 lbmol/h Normal butane to the distillate = 172.9 - 2.2 - 164.3 = 6.4 Material balance around column C3, in lbmol/h: Component Feed Distillate Propane Isobutane Normal butane Isopentane Normal pentane Total 2.2 171.1 226.6 28.1 17.5 445.5 2.2 164.3 6.4 0.0 0.0 172.9 Bottoms 0.0 6.8 220.2 28.1 17.5 272.6 (b) % Purity of nC4 in bottoms = (220.2/272.6) x 100% = 80.8% (c) Let x = lbmol/h of nC4 in the distillate y = lbmol/h of iC4 in the distillate P = mole fraction purity of nC4 in the bottoms 226.6 - x 226.6 - x = (1) (1711 - y ) + (226.6 - x ) + 281 + 17.5 443.3 - y - x . . y Fractional purity of iC 4 in the distillate = = 0.95 (2) 2.2 + y + x Combining (1) and (2) to eliminate x, and optimization of P with respect to y gives: P = 0.828 or 82.8 mol% nC4 in the bottoms, x = 6.8 lbmol/h, y = 171.1 lbmol/h Therefore, 100% recovery of iC4 in the distillate maximizes the purity of nC4 in the bottoms. P= Exercise 1.15 Subject: Sequence of two distillation columns, C1- C2, for the separation of alcohols. Given: 500 kmol/h feed of 40% methanol (M), 35% ethanol (E), 15% isopropanol (IP), and 10% normal propanol (NP), all in mol%. Distillate from column C1 is 98 mol% M, with a 96% recovery. Distillate from column C2 is 92 mol% E, with a recovery of 95% based on the feed to column C1. Find: (a) Component flow rates in the feed, distillates and bottoms. (b) Mol% purity of combined IP and NP in the bottoms from column C2. (c) Maximum achievable purity of E in the distillate from column C2 for 95% recovery of E from the feed to column C1. (d) Maximum recovery of E from the feed to column C1 for a 92 mol% purity of E in the distillate from column C2. Assumptions: Because of the sharp separation in column C1, neglect the presence of propanols in the distillate from column C1. Neglect the presence of M in the bottoms from column C2. The distillate from C2 does not contain normal propanol. (a) M in distillate from C1 = (0.96)(500)(0.40) = 192 kmol/h Total distillate from C1 = 192/0.98 = 195.92 kmol/h E in distillate from C1 = 195.92 - 192 = 3.92 kmol/h E in feed to C2 = (500)(0.35) - 3.92 = 171.08 kmol/h M in feed to C2 = M in distillate from C2 = (500)(0.40) - 192 = 8 kmol/h E in distillate from C2 = (500)(0.35)(0.95) = 166.25 kmol/h Total distillate from C2 = 166.25/0.92 = 180.71 kmol/h IP in distillate from C2 = 180.71 - 166.25 - 8 = 6.46 kmol/h Block flow diagram: Analysis: Exercise 1.15 (continued) Analysis: (a) continued Material balance table (all flow rates in kmol/h): Component Stream 1 2 3 M 200 192.00 8.00 E 175 3.92 171.08 IP 75 0.00 75.00 0.00 50.00 NP 50 Total 500 195.92 304.08 4 8.00 166.25 6.46 0.00 180.71 5 0.00 4.83 68.54 50.00 123.37 The assumption of negligible NP in stream 4 is questionable and should be corrected when designing the column. (b) Mol% purity of (IP + NP) in bottoms of C2 = (68.54 + 50.00)/123.37 or 96.08% (c) If the overall recovery of E in the distillate from C2 is fixed at 95%, the maximum purity of E in that distillate occurs when no propanols appear in that distillate. Then, mol% purity of E = 100% x 166.25/(166.25 + 8.0) = 95.41% (d) The maximum recovery of E in the distillate from C2 occurs when E does not appear in the bottoms from C2. Thus, that maximum is 100% x (171.08/175) = 97.76% Exercise 1.16 Subject: Pervaporation for the partial separation of ethanol and benzene Given: 8,000 kg/h of 23 wt% ethanol and 77 wt% benzene. Polymer membrane is selective for ethanol. Permeate is 60 wt% ethanol. Retentate is 90 wt% benzene. Find: (a) and (b) Component flow rates in feed, permeate, and retentate on a diagram. (c) Method to further separate the permeate. Analysis: (a) and (b) Let: P = permeate flow rate R = retentate flow rate Total material balance: 8,000 = P + R Ethanol material balance: 8,000(0.23) = (0.60) P + (0.10) R Solving (1) and (2) simultaneously, P = 2,080 kg/h and R = 5,920 kg/h (1) (2) The resulting material balance and flow diagram is: (c) Gas adsorption, gas permeation, or distillation to obtain ethanol and the azeotrope, which can be recycled. Exercise 1.17 Subject: Separation of hydrogen from light gases by gas permeation with hollow fibers. Given: Feed gas of 42.4 kmol/h of H2, 7.0 kmol/h of CH4, and 0.5 kmol/h of N2 at 40oC and 16.7 MPa. Retentate exits at 16.2 kPa and permeate exits at 4.56 kPa. Gas heat capacity ratio = = 1.4. Assumptions: Membrane is not permeable to nitrogen. Reversible gas expansion with no heat transfer between the retentate and permeate. Separation index is based on mole fractions. Find: (a) Component flows in the retentate and permeate if the separation index, SP, for hydrogen relative to methane is 34.13, and the split fraction (recovery), SF, for hydrogen from the feed to the permeate is 0.6038. (b) Percent purity of hydrogen in the permeate. (c) Exit temperatures of the retentate and permeate. (d) Process flow diagram with complete material balance Analysis: (a) and (d) Hydrogen in permeate = (0.6038)(42.4) = 25.6 kmol/h Hydrogen in retentate = 42.4 - 25.6 = 16.8 kmol/h Let: x = kmol/h of methane in permeate Then, 7.0 - x = kmol/h of methane in retentate From Eq. (1-4), 25.6 / 16.8 SP = 34.13 = (1) x / (7 - x ) Solving (1), x = 0.3 kmol/h of methane in the permeate Methane in the retentate = 7.0 - 0.3 = 6.7 kmol/h The resulting material balance and flow diagram is: Exercise 1.17 (continued) Analysis: (continued) (b) Percent purity of hydrogen in permeate = 100% x 25.6/25.9 = 98.8% (c) For reversible adiabatic (isentropic) expansion, assuming an ideal gas, the final temperature is given from thermodynamics by: -1 P Tout = T1 out (2) P1 where subscript 1 refers to upstream side, subscript out refers to downstream side, both temperature and pressure are absolute, and is the gas heat capacity ratio. For both the retentate and the permeate, T1 = 40oC = 313 K and P1 =16.7 MPa. For the retentate, Pout = P3 = 16.2 MPa. From (2), 1.4 -1 1.4 Tout 16.2 = T3 = 313 16.7 = 310 K = 37 o C For the permeate, P2 = 4.56 MPa. From (2), 4.56 T2 = 313 16.7 1.4 -1 1.4 = 216 K= - 57 o C Exercise 1.18 Subject: Natural gas is produced when injecting nitrogen into oil wells. The nitrogen is then recovered from the gas for recycle. Given: 170,000 SCFH (60oF and 14.7 psia) of gas containing, in mol%, 18% N2, 75% CH4, and 7% C2H6 at 100oF and 800 psia. Recover the N2 by gas permeation followed by adsorption. The membrane is selective for nitrogen. The adsorbent is selective for methane. The adsorber operates at 100oF, and 275 psia during adsorption and 15 psia during regeneration. Permeate exits the membrane unit at 20oF and a low pressure. Two stages of compression with cooling are needed to deliver the permeate gas to the adsorber. The regenerated gas from the adsorber is compressed in three stages with cooling, and is combined with the retentate to give the natural gas product. Assumptions: The membrane is not permeable to ethane. The separation index, SP, defined by Eq. (1-4), is applied to the exiting retentate and permeate. Find: (a) Draw a labeled process flow diagram. (b) Compute the component material balance, based on the following data: Separation index, SP, for the membrane = 16 for nitrogen relative to methane. The adsorption step gives 97 mol% methane in the adsorbate with an 85% recovery based on the feed to the adsorber. The pressure drop across the membrane is 760 psi. The retentate exits at 800 psia. The combined natural gas product contains 3 mol% nitrogen. Place the results of the material balance in a table. Analysis: (b) Refer to the process flow diagram on next page for stream numbers. Let: ai = molar flow rate of N2 in lbmol/h in stream i. bi = molar flow rate of CH4 in lbmol/h in stream i ci = molar flow rate of ethane in lbmol/h in stream i Feed flow rate = 170,000 SCFM / 379 SCF/lbmole at SC = 448.5 lbmol/h a1 = 0.18(448.5) = 80.7, b1 = 0.75(448.5) = 336.4, c1 = 0.07(448.5) = 31.4 Because ethane does not permeate through the membrane, c3 = c6 = 31.4 and c2 = c4 = c5 = 0 Solve for a2, a3, a4, a5, a6, and b2, b3, b4, b5, b6 from 10 equations in 10 unknowns. Membrane selectivity: a2 / a3 SP = 16 = (1) b2 / b3 Component balances around the membrane unit: a2 + a3 = 80.7 b2 + b3 = 336.4 Component balances around the adsorber: a2 = a4 + a5 b2 = b4 + b5 Exercise 1.18 (continued) Component balances around the line mixer that mixes retentate with adsorbate gas: a6 = a3 + a5 b6 = b3 + b5 Methane purity in the adsorbate: b5 = 0.97(b5 + a5) Find: (b) (continued) Methane recovery: Mol% nitrogen in the final natural gas: b5 = 0.85 b2 a6 = 0.03(a6 + b6 + 31.4) All equations are linear except (1). Solving these 10 equations with a nonlinear equation solver, such as in the Polymath program, results in the following material balance table: Component Nitrogen Methane Ethane Total Stream 1 80.7 336.4 31.4 448.5 2 73.3 128.7 0.0 202.0 Flow rate, 3 7.4 207.7 31.4 246.5 lbmol/h 4 69.9 19.3 0.0 89.2 5 3.4 109.4 0.0 112.8 6 10.8 317.1 31.4 359.3 (a) Labeled process flow diagram Exercise 1.19 Subject: Separation of a mixture of ethylbenzene, o-xylene, m-xylene, and p-xylene Find: (a) Reason why distillation is not favorable for the separation of m-xylene from p-xylene. (b) Properties of m-xylene and p-xylene for determining a means of separation. (c) Why melt crystallization and adsorption can be used to separate m-xylene from p-xylene. Analysis: (a) In the order of increasing normal boiling point: Component Ethylbenzene Paraxylene Metaxylene Orthoxylene nbp, oR 737.3 741.2 742.7 751.1 Relative volatility 1.08 1.02 1.16 From the values of relative volatility, the separation of p-xylene from m-xylene by distillation is not practical. The other two separations are practical by distillation, but require large numbers of stages. (b) From Reference 10, the following properties are obtained: Property Molecular weight van der Waals volume, m3/kmol van der Waals area, m2/kmol x 10-8 Acentric factor Dipole moment, debye Radius of gyration, m x 1010 Normal melting point, K Normal boiling point, K Critical temperature, K Critical pressure, MPa m-xylene 106.167 0.07066 8.84 0.3265 0.30 3.937 225.3 412.3 617 3.541 p-xylene 106.167 0.07066 8.84 0.3218 0.00 3.831 286.4 411.5 616.2 3.511 From the table, the difference of 61.1 K in melting points is very significant and can be exploited in melt crystallization. The difference in dipole moments of 0.30, while not large, makes possible the use of adsorption or distillation with a solvent (c) Explanations are cited in Part (b). Exercise 1.20 Subject: Separation of a near-azeotropic mixture of ethyl alcohol and water. Given: A list of possible methods to break the azeotrope. Find: Reasons why the following separation operations might be used: (a) Extractive distillation (b) Azeotropic distillation (c) Liquid-liquid extreaction (d) Crystallization (e) Pervaporation membrane (f) Adsorption Analysis: Pertinent Properties: Property Melting point, oC Dipole moment, debye Ethanol -112 1.69 Water 0 1.85 (a) Extractive distillation is a proven process using ethylene glycol. (b) Heterogeneous azeotropic distillation is possible with benzene, carbon tetrachloride, trichloroethylene, and ethyl acetate. (c) Liquid-liquid extraction is possible with n-butanol. (d) Crystallization is possible because of the large difference in the melting points. (e) Pervaporation is possible using a polyvinylalcohol membrane. (f) Adsorption is possible using silica gel or activated alumina to adsorb the water. Exercise 1.21 Subject: Removal of ammonia from water. Given: 7,000 kmol/h of water containing 3,000 ppm by weight of ammonia at 350 K and 1 bar. Find: A method to remove the ammonia. Analysis: From Perry's Handbook, 7th edition, page 2-87, the volatility of ammonia is much higher than that of water. Therefore, could use distillation or air stripping. Also, could adsorb the ammonia on a carbon molecular sieve or use a liquid organic membrane containing an acidic complexing agent to form an ion-pair with the ammonia ion, NH4+. Exercise 1.22 Subject: Separation of a mixture of distillation by a sequence of distillation columns. Given: Feed stream containing in kmol/h: 45.4 C3, 136.1 iC4, 226.8 nC4, 181.4 iC5, and 317.4 nC5. Three columns in series, C1, C2, and C3. Distillate from C1 is C3-rich with a 98% recovery. Distillate from C2 is iC4-rich with a 98% recovery. Distillate from C3 is nC4-rich with a 98% recovery. Bottoms from C3 is C5s-rich with a 98% recovery. Find: (a) Process-flow diagram like Figure 1.9. (b) Material-balance table like Table 1.5. (c) Mole % purities in a table like Table 1.7 Assumptions: Reasonable values for splits of iC4 to streams where they are not the main component. Analysis: (a) Process-flow diagram: 2 4 6 1 C 1 C 2 C 3 3 5 7 Exercise 1.22 (continued) (b) Using given recoveries: Propane in Stream 2 = 0.98(45.4) = 44.5 kmol/h i-Butane in Stream 4 = 0.98(136.1) = 133.4 kmol/h n-Butane in Stream 6 = 0.98(226.8) = 222.3 kmol/h C5s in Stream 7 = 0.98(181.4 + 317.5) = 488.9 kmol/h Material-balance table in kmol/h: Comp C3 iC4 nC4 iC5 nC5 Total 1 45.4 136.1 226.8 181.4 317.5 907.2 2 44.5 1.3 0.5 0.0 0.0 46.3 0.9 134.8 226.3 181.4 317.5 860.9 3 0.9 133.4 2.5 0.0 0.0 136.8 4 0.0 1.4 223.8 181.4 317.5 724.1 5 0.0 1.4 222.3 7.0 3.0 233.7 6 0.0 0.0 1.5 174.4 314.5 490.4 7 (c) Mol% purity of propane-rich product = 45.4/46.3 = 0.961 = 96.1% Mol% purity of i-butane-rich product = 133.4/136.8 = 0.975 = 97.5% Mol% purity of n-butane-rich product = 222.3/233.7 = 0.951 = 95.1% Mol% purity of C5s-rich product = (174.4 + 314.5)/490.4 = 0.997 = 99.7% Exercise 1.23 Subject: Methods for removing organic pollutants from wastewater. Given: Available industrial processes: (1) adsorption (2) distillation (3) liquid-liquid extraction (4) membrane separation (5) stripping with air (6) stripping with steam Find: Advantages and disadvantages of each process. Analysis: Some advantages and disadvantages are given in the following table: Method Adsorption Distillation L-L extraction Membrane Air stripping Steam stripping Advantages Adsorbents are available. May be practical if pollutant is more volatile. Solvent are available. May be practical if a membrane can be found that is highly selective for pollutant. May be practical if pollutant is more volatile. May be practical if pollutant is more volatile. Disadvantages Difficult to recover pollutant. Best to incinerate it. Impractical is water is more volatile. Water will be contaminated with solvent. May need a large membrane area if water is the permeate. Danger of producing a flammable gas mixture. Must be able to selectively condense pollutant from overhead. With adsorption, can incinerate pollutant, but with a loss of adsorbent. With distillation, may be able to obtain a pollutant product. With L-L extraction, will have to separate pollutant from solvent. With a membrane, may be able to obtain a pollutant product. With air stripping, may be able to incinerate pollutant. With steam stripping may be able to obtain a pollutant product. Exercise 1.24 Subject: Removal of VOCs from a waste gas stream. Given: Waste gas containing VOCs that must be removed by any of the following methods: (1) absorption (2) adsorption (3) condensation (4) freezing (5) membrane separation Find: Advantages and disadvantages of each method. Analysis: Some advantages and disadvantages are given in the following table: Method Absorption Adsorption Condensation Freezing Membrane Advantages Good absorbents probably exist. Good adsorbents probably exist. May be able to recover the VOC as a product. May be able to recover the VOC as a product. May be able to recover the VOC as a product. Disadvantages Absorbent may stripped into the waste gas. May have to incinerate the spent adsorbent. May require high pressure and/or low temperature. May require a low temperature. May be difficult to obtain high selectivity. May require a very high pressure. With absorption, may be able to distil the VOC from the absorbent. With adsorption, may be able to incinerate the VOC or recover it. With condensation, can recover the VOC as a product. With freezing, can recover the VOC as a product. With a membrane, can recover the VOC as a product. The process shown on the following page shows a process for recovering acetone from air. In the first step, the acetone is absorbed with water. Although water is far from being the most ideal solvent because of the high volatility of acetone in water, the air will not be contaminated with an organic solvent. The acetone-water mixture is then easily separated by distillation, with recycle of the water. Exercise 1.24 (continued) One possible process-flow diagram: Clean gas Makeup Water Acetone Absorber Distillation Gas Feed Recycle Water Exercise 1.25 Subject: Separation of air into nitrogen and oxygen. Given: Air to be separated. Find: Three methods for achieving the separation. Analysis: Three methods are used commercially for separating air into oxygen and nitrogen: 1. Gas permeation mainly for low capacities 2. Pressure-swing gas adsorption for moderate capacities. 3. Low-temperature distillation for high capacities. Exercise 1.26 Subject: Separation of an azeotropic mixture of water and an organic compound. Given: An azeotropic mixture of water and an organic compound such as ethanol. Find: Suitable separation methods. Analysis: Several suitable methods are: Pervaporation Heterogeneous azeotropic distillation Liquid-Liquid extraction. Exercise 1.27 Subject: Production of magnesium sulfate from an aqueous stream. Given: An aqueous stream containing 5 wt% magnesium sulfate. Find: Suitable process for producing nearly pure magnesium sulfate. Analysis: Use evaporation to a near-saturated solution, followed by crystallization to produce a slurry of magnesium sulfate heptahydrate crystals. Use a centrifuge or filter to remove most of the mother liquor from the crystals, followed by drying. A detailed flow sheet of the process is shown and discussed near the beginning of Chapter 17. Exercise 1.28 Subject: Separation of a mixture of acetic acid and water. Given: A 10 wt% stream of acetic acid in water Find: A process that may be more economical than distillation. Analysis: The solution contains mostly water. Because water is more volatile than acetic acid, distillation will involve the evaporation of large amounts of water with its very high enthalpy of vaporization. Therefore, it is important to consider an alternative method such as liquid-liquid extraction. Such a process is shown and discussed near the beginning of Chapter 8. Exercise 2.1 Subject: Minimum work for separating a hydrocarbon stream. Given: Component flow rates, ni , of feed and product 1, in kmol/h. Phase condition; temperature in K; enthalpy, h, in kJ/kmol; and entropy, s, in kJ/kmol-K for feed, product 1, and product 2. Infinite heat sink temperature = T0 = 298.15 K. Find: Minimum work of separation, Wmin , in kJ/h Analysis: From Eq. (4), Table 2.1, Wmin = out nb - in nb From Eq. 2-1, b = h - T0 s For the feed stream (in), n = 30 + 200 + 370 + 350 + 50 = 1,000 kmol/h b = 19,480 - (298.15)(36.64) = 8,556 kJ/kmol For product 1 (out), n = 30 + 192 + 4 + 0 + 0 = 226 kmol/h b = 25,040 - (298.15)(33.13) = 15,162 kJ/kmol For product 2 (out), n = nfeed - nproduct 1 = 1,000 - 226 = 774 kmol/h b = 25,640 - (298.15)(54.84) = 9,289 kJ/kmol From Eq. (4), Table 2.1, Wmin = 226(15,162) + 774(9,289) - 1,000(8,556) = 2,060 kJ/h Exercise 2.2 Subject: Minimum work for separating a mixture of ethylbenzene and xylene isomers. Given: Component flow rates, ni , of feed ,in lbmol/h. Component split fractions for three products, Phase condition; temperature in oF; enthalpy, h, in Btu/lbmol; and entropy, s, in Btu/lbmol-oR for feed and three products. Infinite heat sink temperature = T0 = 560oR. Find: Minimum work of separation, Wmin , in kJ/h Analysis: From Eq. (4), Table 2.1, Wmin = out nb - in nb From Eq. 2-1, b = h - T0 s For the feed stream (in), n = 150 +190 + 430 + 230 = 1,000 lbmol/h b = 29,290 - (560)(15.32) = 20,710 Btu/lbmol For product 1 (out), using Eq. (1-2), n = 150(0.96) + 190 (0.005) + 430(0.004) = 146.7 lbmol/h b = 29,750 - (560)(12.47) = 22,767 Btu/lbmol For product 2 (out), using Eq. (1-2), n = 150(0.04) + 190(0.99) + 430(0.99) + 230(0.015) = 623.3 lbmol/h b = 29,550 - (560)(13.60) = 21,934 Btu/lbmol For product 3 (out), by total material balance, n = 1,000 - 146.7 - 623.3 = 230 lbmol/h b = 28,320 - (560)(14.68) = 20,099 Btu/lbmol From Eq. (4), Table 2.1, Wmin = 146.7(22,767) + 623.3(21,934) + 230(20,099) - 1,000(20,710) = 924,200 Btu/h Exercise 2.3 Subject: Second-law analysis of a distillation column Given: Component flow rates, ni , from Table 1.5 for feed, distillate, and bottoms in kmol/h for column C3 in Figure 1.9. Condenser duty, QC ,= 27,300,00 kJ/h. Phase condition; temperature in K; enthalpy, h, in kJ/kmol; and entropy, s, in kJ/kmol-K for feed, distillate and bottoms. Infinite heat sink temperature = T0 = 298.15 K. Condenser cooling water at 25oC = 298.15 K and reboiler steam at 100oC = 373.15 K. Assumptions: Neglect shaft work associated with column reflux pump. Find: (a) (b) (c) (d) (e) Analysis: Reboiler duty, QR , kJ/h Production of entropy, Sirr , kJ/h-K Lost work, LW, kJ/h Minimum work of separation, Wmin , kJ/h Second-law efficiency, (a) From Eq. (1), the energy balance for column C3, nh + in out QR = QC - nh = 27,300,000 - 445.5(17,000) + 175.5(13,420) + 270(15,840) = 26,360,000 kJ/h (b) From Eq. (2), the entropy balance for column C3, Sirr = out ns + Q - Ts ns + in Q Ts = 175.5(5.87) + 270(21.22) + = 16,520 kJ/h-K (c) From Eq. (2-2), LW = T0Sirr = 298.15(16,520) = 4,925,000 kJ/h (d) Combining Eqs. (3) and (4) of Table 2.1, Wmin = QR 1 - T0 T - QC 1 - 0 - LW Tsteam Tcw 27,300,000 26, 360, 000 - 445.5(25.05) - 298.15 373.15 = 26,360, 000 1 - 298.15 298.15 - 27, 300, 000 1 - - 4, 925, 000 373.15 298.15 = 373, 000 kJ/h (e) From Eq. (5), Table 2.1, = 373,000/(4,925,000 + 373,000) = 0.0704 = 7.04% Exercise 2.4 Subject: Second-law analysis for a membrane separation of a gas mixture Given: Component flow rates in lbmol/h for the feed. Permeate of 95 mol% H2 and 5 mol% CH4. Separation factor, SP, for H2 relative to CH4, of 47. Phase condition; temperature in oF; enthalpy, h, in Btu/lbmol; and entropy, s, in Btu/lbmol-oR for the feed, permeate, and retentate. Infinite heat sink temperature = T0 = 539.7oR. Assumptions: Neglect heat transfer to or from the membrane. Find: (a) Production of entropy, Sirr ,Btu/h-oR (b) Lost work, LW, Btu/h (c) Minimum work of separation, Wmin , Btu/h (d) Second-law efficiency, Suggest other separation methods. Analysis: First compute the material balance to obtain the flow rates of the retentate and permeate. From Eq. (1-4), for the separation factor, SP, using the subscript P for permeate and R for retentate, nH 2 / nH 2 P R 47 = (1) nCH 4 / nCH 4 P R For 95 mol% H2 and 5 mol% CH4 in the permeate, nH 2 P 0.95 = nH 2 + nCH 4 P P (2) By component material balances for H2 and CH4 around the membrane separator. nH 2 = 3,000 = nH 2 + nH 2 F P R (3) and (4) nCH 4 = 884 = nCH 4 + nCH 4 F P R Solving Eqs. (1), (2), (3), and (4) for the 4 unknowns, nH 2 P = 2,699.9 lbmol / h, nH 2 R = 3001 lbmol / h, nCH 4 . P = 142.1 lbmol / h, nCH 4 R = 741.9 lbmol / h Therefore, nP = 2,699.9 + 142.1 + 0 = 2,842 lbmol/h nR = 300.1 + 741.9 + 120 = 1,162 lbmol/h Also, nF = 2,842 +1,162 = 4,004 lbmol/h Exercise 2.4 (continued) Analysis: (continued) (a) From Eq. (2), the entropy balance for the membrane Sirr = out ( ns ) - in ( ns ) = 1,162(2.742) + 2,842(4.222) - 4, 004(1.520) = 9,099 Btu/h- o R (c) From Eq. (2-2), LW = T0Sirr = 536.7(9,099) = 4,883,000 Btu/h (d) Combining Eqs. (3) and (4) of Table 2.1, Wmin = -LW = -4,883, 000 Btu/h Note the negative value. The energy of separation is supplied from the high pressure of the feed. (e) Because the minimum work of separation is equal to the negative of the lost work, Eq. (5), Table 2.1 does not apply. The efficiency can not be computed unless heat transfer is taken into account to give a permeate temperature of 80oF. Gas adsorption could also be used to make the separation. Exercise 2.5 Subject: Expressions for computing K-values. Given: Seven common thermodynamic expressions for estimating K-values. Find: Expression assumptions if not rigorous Analysis: iL is a rigorous expression. iV (b) K i = iL is not rigorous. It assumes ideal solutions so that iL and iV are 1.0. iV (c) K i = iL is not rigorous. It assumes ideal solutions and an ideal gas, such that iV , iL , and iV are 1.0. (d) K i = iL iL is a rigorous expression. iV (a) K i = (e) K i = Pi s is not rigorous. It assumes ideal solutions and an ideal gas, P such that iV , iL , and iV are 1.0. It also assumes a low pressure for the liquid phase, such that iL =1.0. (f) K i = (g) K i = iLiL is a rigorous expression iV iV iL Pi s P is not rigorous. It assumes low pressure for the liquid and the ideal gas law, such that iL = Pi s and iV =1.0. P Exercise 2.6 Subject: Comparison of experimental K-values to Raoult's law predictions. Given: For the propane-isopentane system at 167oF and 147 psia, propane mole fractions of 0.2900 in the liquid phase and 0.6650 in the vapor phase. Vapor pressures at 167oF of 409.6 psia for propane and 58.6 psia for isopentane. Find: (a) Experimental K-values. (b) Raoult's law K-values. Analysis: (a) Mole fraction of isopentane in the liquid phase = 1 - 0.2900 = 0.7100 Mole fraction of isopentane in the vapor phase = 1 - 0.6650 = 0.3350 From Eq. (2-9), Ki = yi /xi = 0.6650/0.2900 = 2.293 for propane = 0.3350/0.7100 = 0.472 for isopentane (b) From Eq. (3), Table 2.3, K i = Pi s / P = 409.6/147 = 2.786 for propane = 58.6/147 = 0.399 for isopentane This is rather poor agreement. Note that the experimental values give a relative volatility of 2.293/0.472 = 4.86, while Raoult's law predicts 2.786/0.399 = 6.98. A modified Raoult's law should be used to incorporate a liquid-phase activity coefficient. Also a fugacity correction for the gas phase might improve the agreement. Exercise 2.7 Subject: Liquid-liquid phase equilibrium data. Given: Experimental solubility data at 25oC for the isooctane (1)-furfural (2) system. In the furfural-rich liquid phase, I, x1 = 0.0431 In the isooctane-rich liquid phase, II, x1 = 0.9461 Assumptions: The furfural activity coefficient in phase I = 1.0 The isooctane activity coefficient in phase II = 1.0 Find: (a) The distribution coefficients for isooctane and furfural. (b) The relative selectivity for isooctane relative to furfural. (c) The activity coefficients. Analysis: From summation of mole fractions in each liquid phase, x2 in phase I = 1 - 0.0431 = 0.9569 x2 in phase II = 1 - 0.9461 = 0.0.0539 (a) From Eq. (2-20), K Di = xi( I ) / xi( II ) K D1 = 0.0431 0.9461 = 0.456 and K D2 = = 17.75 0.9461 0.0539 ij = K Di / K D j (b) From Eq. (2-22), 12 = they K D1 K D2 = 0.0456 = 0.00257 17.75 Note that the I and II designations for the two liquid phases are arbitrary. If were interchanged, the relative selectivity would be 1/0.00257 = 389. (c) From rearrangements of Eq. (2-30), 1( I) = 1(II) (II) 2 x1(II) 0.9461 = 1.0 = 21.95 (I) x1 0.0431 (I) x2 0.9569 = 1.0 = 17.75 (II) x2 0.0539 = (I) 2 Exercise 2.8 Subject: Activity coefficients of solids dissolved in solvents. Given: Solubility at 25oC of naphthalene in 5 solvents. Vapor pressure equations for solid and liquid naphthalene Find: Activity coefficient of naphthalene in each solvent. Analysis: From a rearrangement of Eq. (2-34) for solid-liquid phase equilibrium of naphthalene, L = s 1 Psolid s x L Pliquid (1) From the given naphthalene vapor pressure equations at 25oC = 298.15 K, s Psolid = exp 26.708 - 8,712 = 0.080 torr 298.15 3992.01 = 0.234 torr 29815 - 71.29 . s Pliquid = exp 16.1426 - From Eq. (1), L = 1 0.080 0.342 = x L 0.234 xL (2) Using the solubility data with Eq. (2), Solvent Benzene Cyclohexane Carbon tetrachloride n-Hexane Water Solubility, xL 0.2946 0.1487 0.2591 0.1168 0.18 x 10-5 L 1.16 2.30 1.32 2.93 190,000 Exercise 2.9 Subject: Minimum isothermal work of separation for a binary gas mixture. Given: A feed gas mixture, F, of A and B to be separated at infinite surroundings temperature, T0, into two products, P1 and P2. Assumptions: Ideal gas law and ideal gas solution at temperature T0. Isobaric at P0. Find: Derive an equation for the Wmin in terms of the mole fractions of the feed and products. Plot Wmin/RT0nF versus the mole fraction of A in the feed for: (a) A perfect separation. (b) A separation with SFA = 0.98 and SFB = 0.02. (c) A separation with SRA = 9.0 and SRB = 1/9. (d) A separation with SFA = 0.95 and SRB = 361 Determine the sensitivity of Wmin to product purities. Does Wmin depend on the separation method? Prove that the largest value of Wmin occurs for an equimolar feed. Analysis: From Eq. (4), Table 2.1, Wmin = out nb - in nb (1) From Eq. 2-1, b = h - T0 s (2) Combining Eqs. (1) and (2) for one feed, F, in and two products, P1 and P2, out: Wmin = nP1 hP1 + nP2 hP2 - T0 nP1 sP1 + nP2 hP2 - nF hF + T0nF sF (3) However, for isothermal separation of an ideal gas mixture, the change in enthalpy = 0. Therefore, from Eq. (3), Wmin = T0nF sF - T0 nP1 sP1 + nP2 hP2 (4) From Eq. (3), Table 2.4, for an ideal gas mixture at T0 and P0 , s = -R i yi ln yi (5) Exercise 2.9 (continued) Analysis: (continued) Substituting Eq. (5) into Eq. (4), Wmin = - nF yA F ln yA F + 1 - yA F ln 1 - yA F RT0 + nP1 yA P ln yA P + 1 - yA P ln 1 - yA P 1 1 1 2 2 2 1 (6) 2 + nP2 yA P ln yA P + 1 - yA P ln 1 - yA P } By combining a component material balance for A with a total material balance, nP1 = nF yA F - yA P 1 2 yA P - yA P (7) 2 nP2 = nF yA F - yA P 2 1 yA P - yA P (8) 1 Equations (6), (7), and (8) give a relationship for Wmin/RT0 in terms of the molar feed rate and the mole fractions of A in the feed and two products. (8), (a) Let product P1 be pure A and product P2 be pure B. Then, from Eqs. (7) and nP1 = yA F nF and nP2 = 1 - yA F nF (9) and (10) Combining Eqs. (6) through (10), noting that 1x ln(1) = 0 and 0x ln(0) = 0 Wmin = - yAF ln yA F + (1 - yAF ) ln (1 - yAF ) nF RT0 (b) From Eq. (1-2), letting 1 be P1, (11) yA P nP1 = 0.98 yA F nF or yA P = 0.98 yA F 1 1 nF nP1 (12) Exercise 2.9 (continued) Analysis: (b) (continued) and, yBP = 0.02 1 - yA F 1 nF nP1 (13) Combining (12 and (13) to give yA P + yBP = 1, we obtain, 1 1 nP1 = nF 0.96 yA F + 0.02 By total material balance, (14) nP2 = nF - nP1 or nP2 = nF 0.98 - 0.96 yA F (15) Also, from the SFA = 0.98 for the split fraction of A to P1 , we can write for the split fraction of A to P2 , yA P nP2 = 0.02 yA F nF or yA P = 0.02 yA F 2 2 nF nP2 (16) The final equations are (6) combined with (12) and (14) through (16). (c) From Eq. (1-3), yA P nP1 1 yA P nP2 2 =9 and, (17) yBP nP1 1 yBP nP2 2 = 1 9 (18) Combining (17) and (18), with component and total material balances around the separator gives the following equations that can be used with Eq. (6): yA P = 0.9 yA F 1 nF nP1 (19) Exercise 2.9 (continued) Analysis: (c) (continued) yA P = 01yA F . 2 nF nP2 (20) nP1 = nF 0.8 yA F + 01 . nP2 = nF 0.9 - 0.8 yA F (d) From Eq. (1-5), (21) (22) SPA,B = SFA / SFB = 361 1 - SFA / 1 - SFB (23) SFA = 0.95 is given. Combining this with Eq. (23), gives SFB = 0.05. This part then proceeds as in part (b) to give: nP1 = nF 0.90 yA F + 0.05 (24) nP2 = nF - nP1 or nP2 = nF 0.95 - 0.90 yA F (25) yA P = 0.95 yA F 1 nF nP1 (27) (26) yA P = 0.05 yA F 2 nF nP2 Equations (24) through (27) are combined with Eq. (6). A spreadsheet can be used to compute the dimensionless minimum work, Wmin RT0nF with the following results for Parts (a) through (d): Exercise 2.9 (continued) From the plot on the previous page, it is seen that the dimensionless minimum work is very sensitive to the feed mole fraction and to the product purities. From the derivation of the minimum work equations, it is seen that they are independent of the separation method and only depend on thermodynamics. To prove that the largest value of Wmin occurs for a feed with equimolar quantities of A and B, consider the case of a perfect separation, part (a), as given by Eq. (11). Let W = Wmin/nFRT0 and y = yA F . Then, the derivative of W with respect to y is, dW y -1 = - ln y + - ln(1 - y ) + (1 - y ) dy y 1- y = - [ ln y - ln(1 - y ) ] For min/max, set the derivative to zero and solve for y. Therefore, dW = 0 = - [ ln y - ln(1 - y ) ] dy Solving, y = 1 - y or y = 0.5. This is an equimolar feed. Furthermore, it gives a maximum value of W . Exercise 2.10 Subject: Relative volatility of the isopentane-normal pentane system Given: Experimental data for relative volatility 125-250oF. Vapor pressure constants. Find: Relative volatilities from Raoult's law over the same temperature range and compare them to the experimental values. Analysis: Combining Eq. (5), Table 2.3, for Raoult's law, with Eq. (2-21) for . s PiC iC5 ,nC5 = s 5 (1) PnC5 Only the first three constants of the extended Antoine equation are given. (T=K, P= kPa) 2345.09 s PiC5 = exp 13.6106 - T - 40.2128 (2), (3) 2554.60 s PnC5 = exp 13.9778 - T - 36.2529 Using a spreadsheet to calculate the relative volatility from Eqs. (1), (2), and (3): s T, F T, K Raoult's law Expt. PiC5 , kPa PnsC5 , kPa 125 1.26 325 214 167 1.28 150 1.23 339 314 251 1.25 352 446 364 1.22 175 1.21 200 1.18 366 614 512 1.20 225 1.16 380 823 700 1.18 250 1.14 394 1079 834 1.16 The Raoult's law values are within 2% of the experimental values. Exercise 2.11 Subject: Condenser duty of a vacuum distillation column separating ethyl benzene (EB) and styrene (S). Given: Phase condition, temperature, pressure, flow rate, and compositions for streams entering and leaving a condenser, which produces subcooled reflux and distillate. Property constants in Example 2.3. Assumptions: Ideal gas and ideal gas and liquid solutions. Find: Condenser duty in kJ/h. Analysis: For the thermodynamic path, cool the overhead from 331 K to 325 K. Then condense at 325 K. Because this temperature change is small, compute vapor specific heats at the two temperatures and take the arithmetic average. From Eq. (2-35) and the constants in Example 2.3, for vapor heat capacity in J/kmol-K and temperature in K, o CPEB = -43098.9 + 707.151T - 0.481063T 2 + 130084 10 -4 T 3 . o CPS = -28248.3 + 615.878T - 0.40231T 2 + 0.993528 10 -4 T 3 Solving, Comp. EB S kg/h 77,500 2,500 MW 106.17 104.15 kmol/h 729.975 24.003 Vapor heat capacity, J/kmol-K 331 K 325 K Avg. 142,980 140,380 141,680 132,130 132,830 133,980 The sensible vapor enthalpy change, Qsensible from 331 K to 325 K is, Qsensible = i o ni CPi T = i o ni CPi (331 - 325) = 729.975(141,680)(331 - 325) + 24.003(133,980)(331 - 325) = 620,560,000 J / h = 620,560 kJ / h For the latent heat of condensation, use Eq. (2-41) to estimate the molar heats of vaporization of at 325 K for the two components, using the vapor pressure constants given in Example 2.3. Exercise 2.11 (continued) -7440.61 325 2 vap HEB = 8314 325 2 - + 0.00623121 + -9.87052 5 + 6(4.13065 10-18 ) 325 325 = 40,740,000 J / kmol = 40,740 kJ / kmol HSvap = 8314 325 2 - -914107 . 325 2 + 0.0143369 + -17.0918 5 + 6(18375 10-18 ) 325 . 325 = 42,440,000 J / kmol = 42,440 kJ / kmol The latent heat of condensation, Qlatent , in kJ/kmol is, Qlatent = i ni Hivap = 729.975(40,740) + 24.003(42,440) = 30,760,000 kJ / h The total condenser duty = Qsensible + Q latent = 620,000 + 30,760,000 = 31,380,000 kJ/h Exercise 2.12 Subject: Thermodynamic properties of a benzene (B) -toluene (T) feed to a distillation column. Given: Temperature, pressure, and component flow rates for the column feed. Property constants, critical temperature Assumptions: Phase condition is liquid (needs to be verified). Ideal gas and liquid solution. Find: Molar volume, density, enthalpy, and entropy of the liquid feed. Analysis: The feed is at 100oF and 20 psia. From Fig. 2.4, since the vapor pressures of benzene and toluene are 3.3 and 0.95 psia, respectively, the feed is a subcooled liquid. T = 311 K. From Eqs. (4) , Table 2.4 and (2-38), L = i L Mi i = Mi AB 78.11 - 1- T Tc 2/ 7 L = B L = T 304.1(0.269) 92.14 290.6(0.265) 311 - 1- 562 2 /7 = 0.0905 m3 / kmol = 01083 m3 / kmol . 311 - 1- 593.1 2/ 7 Total flow rate = 415 + 131 = 546 kmol/h Benzene mole fraction = 415/546 = 0.760 Toluene mole fraction = 1-0.76 = 0.240 From Eq. (4), Table 2.4 for a mixture (additive volumes), L = 0.76(0.0905) + 0.24(0.1083) = 0.0948 m3/kmol = 1.52 ft3/lbmol Mixture molecular weight = M = 0.76(78.11) + 0.24(92.14) = 81.48 From Eq. (4), Table 2.4 for conversion to density, L = M/L = 81.48/1.52 = 53.6 lb/ft3 Exercise 2.13 Subject: Liquid density of the bottoms from a distillation column. Given: Temperature, pressure, component flow rates Assumptions: Ideal liquid solution so that volume of mixing = 0 Find: Liquid density in various units using Fig. 2.3 for pure component densities. Analysis: From Eq. (4), Table 2.4, L = i xi iL = M M L iL = iL The calculations are summarized as follows using Fig. 2.3 for pure component densities, Comp. C3 iC4 nC4 iC5 nC5 MW 44.09 58.12 58.12 72.15 72.15 Flow rate, lbmol/h 2.2 171.1 226.6 28.1 17.5 Mole fraction, xi 0.0049 0.3841 0.5086 0.0631 0.0393 Density, Fig. 2.3, g/cm3 0.2 0.40 0.43 0.515 0.525 Molar volume, cm3/mol 220 145 135 140 130 xi 1.1 55.7 68.7 8.8 5.1 L = i xi iL = 139.4 cm3/mol MW 59.5 = = 0.427 g/cm 3 L 139.4 MW = 59.5 L = = 427 kg/m3 = 26.6 lb/ft3 = 3.56 lb/gal = 149 lb/bbl where 1 bbl = 42 gal Exercise 2.14 Subject: Condenser duty for a distillation column, where the overhead vapor condenses into two liquid phases. Given: Temperature, pressure, and component flow rates of overhead vapor and the two liquid phases. Assumptions: Ideal gas and ideal liquid solution for each liquid phase. Find: Condenser duty in Btu/h and kJ/h Analysis: Take a thermodynamic path of vapor from 76oC to 40oC and condensation at 40oC. For water, use the steam tables. Change in enthalpy from vapor at 76oC to 40oC and 1.4 bar = 1133.8 - 71.96 = 1,062 Btu/lb = 2,467,000 J/kg = 2,467 kJ/kg. For benzene, using data on p. 2-221 from Perry's 7th edition, change in enthalpy from vapor at 76oC to 40oC and 1.4 bar = 874 - 411 = 463 kJ/kg For isopropanol, using data on p. 2-179 from Perry's 7th edition, Average CP over the temperature range = 1.569 kJ/kg-K From data of p. 2-157 of Perry's 7th edition, Heat of vaporization at 40OC = 313 K = 770.4 kJ/kg Therefore, the enthalpy change of isopropanol = 1.569(76-40) +770.4 = 827 kJ/kg Condenser duty, QC = 2,350(2,467) + 24,600(463) + 6,800(827) = 22,810,000 kJ/h = 21,640,000 Btu/h Exercise 2.15 Subject: K-values and vapor tendency of light gases and hydrocarbons Given: Temperature of 250oF and pressure of 500 psia. Find: K-values in Fig. 2.8 and vapor tendency. Analysis: If the K-value is < 1.0, tendency is for liquid phase. If the K-value > 1.0, tendency is for vapor phase. Using Fig. 2.8, Component N2 H2S CO2 C1 C2 C3 iC4 nC4 iC5 nC5 K-value 17 3.1 5.5 8 3 1.5 0.71 0.35 0.38 0.10 Tendency vapor vapor vapor vapor vapor vapor liquid liquid liquid liquid Exercise 2.16 Subject: Recovery of acetone from air by absorption in water. Given: Temperature, pressure, phase condition, and component flow rates of feeds to and products from the absorber, except for exiting liquid temperature. Assumptions: Ideal gas and zero heat of mixing. Find: Temperature of exiting liquid phase. Potential for explosion hazard. Analysis: From the given component flow rates, water evaporates at the rate of 22 lbmol/h, and 14.9 lbmol/h of acetone is condensed. Take a thermodynamic path that evaporates water at 90oF and condenses acetone at 78oF. Energy to heat air from 78oF to 80oF = nCPT = 687(7)(80-78)=9,620 Btu/h Energy to heat unabsorbed acetone from 78oF to 80oF is negligible. Energy to vaporize water at 90oF = nHvap = 22(1,043)(18) = 413,000 Btu/h Total required energy = 9,620 + 413,000 = 423,000 Btu/h Energy available from condensation of acetone with Hvap = 237 Btu/lb and a molecular weight of 58.08 = 14.9(58.08)(237) = 205,000 Btu/h Energy available from the cooling of evaporated water from 90oF to 80oF = 22(18)(0.44)(90-80) = 2,000 Btu/h Total available energy = 205,000 + 2,000 = 207,000 Btu/h Energy required - Energy available = 423,000 - 207,000 = 216,000 Btu/h This energy must come from cooling of the water absorbent from 90oF to T , and condensed acetone from 78oF to T. Therefore, using a CP of 0.53 for liquid acetone, 216,000 = 14.9(58.08)(78-T) + 1,722(18)(1.0)(90-T) Solving, T = temperature of exiting liquid = 83oF. The mol% acetone in the entering gas = 15/702 x 100% = 2.14 %. This is outside of the explosive limits range of 2.5 to 13 mol%. Exercise 2.17 Subject: Volumetric flow rates of entering and exiting streams of an adsorber for removing nitrogen from subquality natural gas. Given: Temperature and pressure of feed gas and two product gases. Composition of the feed gas. Specification of 90% removal of nitrogen and a 97% methane natural gas product. Assumptions: Applicability of the Redlich-Kwong equation of state. Find: Volumetric flow rates of the entering and exiting gas streams in actual ft3/h. Analysis: Removal of nitrogen = 0.9(176) = 158.4 lbmol/h Nitrogen left in natural gas = 176 - 158.4 = 17.6 lbmol/h Methane in natural gas product = 17.6(97/3) = 569.1 lbmol/h Material balance summary: Component Nitrogen Methane Totals Temperature, oF Pressure, psia lbmol/h: Feed gas 176 704 880 70 800 Waste gas 158.4 134.9 293.3 70 280 Natural gas 17.6 569.1 586.7 100 790 Using the ChemCAD simulation program, the following volumetric flow rates are computed using the Redlich-Kwong equation of state: Stream Feed gas Waste gas Natural gas Actual volumetric flow rate, ft3/h 5,844 5,876 4,162 Exercise 2.18 Subject: Estimation of partial fugacity coefficients of propane and benzene using the R-K equation of state. Given: From Example 2.5, a vapor mixture of 39.49 mol% propane and 60.51 mol% benzene at 400oF and 410.3 psia. Assumptions: Applicability of the Redlich-Kwong equation of state. Find: Partial fugacity coefficients Analysis: From Example 2.5, the following conditions and constants apply, where the Redlich-Kwong constants, A and B, for each component are computed from Eqs. (2-47) and (2-48) respectively. T = 477.6 K P = 2829 kPa ZV = 0.7314 R = 8.314 kPa-m3/kmol-K A = 0.2724 B = 0.05326 From Eqs. (2-47) and (2-48), aP ai (2819) a Ai = 2i 2 = = i 2 2 RT (8.314) (477.6) 5593 bP bi (2819) Bi = i = = 0.7099bi RT (8.314)(477.6) Component, i Propane Benzene a, kPa-m6/kmol2 836.7 2,072 b, m3/kmol 0.06268 0.08263 Ai 0.1496 0.3705 Bi 0.04450 0.05866 The Redlich-Kwong equation for partial vapor fugacity coefficients is given by Eq. (256). Applying it to propane (P) and benzene (B), using the above values for constants and conditions, PV = exp (0.7314 - 1) = 0.934 0.05866 0.2724 0.3705 0.05866 0.05326 - ln ( 0.7314 - 0.05326 ) - 2 - ln 1 + 0.05326 0.05326 0.2724 0.05326 0.7314 0.04450 0.2724 0.1496 0.04450 0.05326 - ln ( 0.7314 - 0.05326 ) - 2 - ln 1 + 0.05326 0.05326 0.2724 0.05326 0.7314 BV = exp (0.7314 - 1) = 0.705 Exercise 2.19 Subject: Estimation of K-values by the P-R and S-R-K equations of state for a butanesbutenes stream. Given: Experimental K-values for an equimolar mixture of isobutane, isobutene, nbutane, 1-butene, trans-2-butene, and cis-2-butene at 220oF and 276.5 psia. Find: K-values by the P-R and S-R-K equations of state using a process simulator. Analysis: Using the ChemCAD process simulation program, the following values are obtained and compared to the experimental values: Component Isobutane Isobutene n-butane 1-butene Trans-2-butene Cis-2-butene Experimental K-value 1.067 1.024 0.922 1.024 0.952 0.876 P-R K-value 1.088 1.029 0.923 1.015 0.909 0.882 S-R-K K-value 1.095 1.036 0.929 1.022 0.916 0.889 The experimental and estimated K-values agree to within 3% for all components except trans-2-butene. For that component, the P-R and S-R-K values are in close agreement, but deviate from the experimental value by from 4 to 5%. Exercise 2.20 Subject: Cooling and partial condensation of the reactor effluent in a toluene disproportionation process. Given: Reactor effluent component flow rates, and temperature and pressure before and after a cooling-water heat exchanger. Find: Using the S-R-K and P-R equations of state with a process simulation program, compute the component K-values, and flow rates in the vapor and liquid streams leaving the cooling-water heat exchanger, and the rate of heat transfer in the cooling-water heat exchanger. Analysis: Using the ChemCAD simulation program, the following phase equilibrium results are obtained for 100oF and 485 psia. S-R-K Equation of State: lbmol/h: Component K-value Reactor effluent Equilib. Vapor Equilib. liquid Hydrogen 85.2 1900 1873.79 26.21 Methane 10.12 215 192.35 22.65 Ethane 1.715 17 10.03 6.97 Benzene 0.00827 577 3.98 573.02 1349 2.98 1346.02 Toluene 0.00264 Paraxylene 0.000881 508 0.37 507.63 Total 4566 2083.51 2482.49 P-R Equation of State: Component Hydrogen Methane Ethane Benzene Toluene Paraxylene Total lbmol/h: K-value 34.4 11.27 1.890 0.0110 0.00359 0.001008 Reactor effluent 1900 215 17 577 1349 508 4566 Equilib. Vapor 1834.46 193.85 10.30 4.95 3.93 0.42 2047.91 Equilib. liquid 65.54 21.15 6.70 572.05 1345.07 507.58 2518.09 Except for hydrogen, the results are in good agreement. The rate of heat transfer is computed by an energy balance, using an exchanger inlet condition of 235oF and 490 psia. Stream enthalpies are obtained from the ChemCAD program. For S-R-K, QC = 25,452,000 - (-6,018,000) - 15,059,000 = 16,411,000 Btu/h For P-R, QC = 26,001,000 - (-6,057,000) - 15,928,000 = 16,130,000 Btu/h Exercise 2.21 Subject: Minimum work for the separation of a nonideal liquid mixture. Given: A 35 mol% acetone (1) and 65 mol% water (2) liquid mixture at 298 K and 101.3 kPa, to be separated into 99 mol% acetone and 98 mol% water. Van Laar constants for the system. Find: Minimum work for the separation in kJ/kmol of feed. Analysis: Material balance for 1 kmol of feed. Let the two products be R and S, where the former is acetone-rich. Total mole balance: 1 = nR + nS Acetone balance: 0.35 = 0.99 nR + 0.02 nS Solving, nR = 0.3469 kmol and nS = 0.6531 kmol From the problem statement, Wmin = nR RT0 xi ln i xi iR + nS iS xi ln i xi - nF iS xi ln i xi (1) The activity coefficients are given by the van Laar equations, (3) in Table 2.9, which with the given constants, A12 = 2.0 and A21 = 1.7, become, ln 1 = 2.0 2.0 x1 1+ 1.7 x2 2 (2) ln 2 = . 17 1.7 x2 1+ 2.0 x1 2 (3) Applying Eqs. (2) and (3) to the given mole-fraction compositions, Component Acetone (1) Water (2) x 0.35 0.65 Feed 2.116 1.291 x 0.99 0.01 Product R 1.000 5.318 Product S_____ x 0.02 6.735 0.98 1.000 Substituting the above values into Eq. (1), Exercise 2.21 (continued) Analysis: (continued) Wmin = 0.3469 {0.99 ln [ (0.99)(1.00)] + 0.01ln [ (0.01)(5.318) ]} RT0 +0.6531{0.02 ln [ (0.02)(6.735) ] + 0.98ln [ (0.98)(1.00) ]} = 0.1664 kmol Wmin = 0.1664 RT0 = 0.1664(8.314)(298) = 412 kJ/kmol feed For ideal liquid solutions, all values of are equal to 1.0. Therefore, Wmin = 0.3469 {0.99 ln [ (0.99)(1.00)] + 0.01ln [ (0.01)(1.00]} RT0 - 1.0000 {0.35 ln [ (0.35)(2.116)] + 0.65 ln [ (0.65)(1.291)]} +0.6531{0.02 ln [ (0.02)(1.00)] + 0.98 ln [ (0.98)(1.00) ]} = 0.5639 kmol - 1.0000 {0.35 ln [ (0.35)(1.00)] + 0.65 ln [ (0.65)(1.00)]} Wmin = 0.5639 RT0 = 0.5639(8.314)(298) = 1,397 kJ/kmol feed Thus, the minimum work of separation for the ideal solution is 3.4 times that of the nonideal solution. Exercise 2.22 Subject: Relative volatility and activity coefficients of the benzene (B) - cyclohexane (CH) azeotropic system at 1 atm Given: Experimental vapor-liquid equilibrium data, including liquid-phase activity coefficients, and Antoine vapor pressure constants. Assumptions: Ideal gas and gas solutions Find: (a) Relative volatility of benzene with respect to cyclohexane as a function of benzene mole fraction in the liquid phase. (b) Van Laar constants from the azeotropic point and comparison of van Laar predictions with experimental data. Analysis: (a) From Eqs. (2-21), (2-19), and (3) in Table 2.3, B,CH = KB y /x yB / xB Ps = B B = = B Bs KCH yCH / xCH CH PCH 1 - yB / 1 - xB (1) Using the y-x data for benzene, the following values of relative volatility are computed from (1): Temperature, oC 79.7 79.1 78.5 78.0 77.7 77.6 77.6 77.6 77.8 78.0 78.3 78.9 79.5 xB 0.088 0.156 0.231 0.308 0.400 0.470 0.545 0.625 0.701 0.757 0.822 0.891 0.953 yB 0.113 0.190 0.268 0.343 0.422 0.482 0.544 0.612 0.678 0.727 0.791 0.863 0.938 B,CH 1.317 1.269 1.218 1.173 1.095 1.049 0.996 0.946 0.898 0.854 0.821 0.771 0.746 Exercise 2.22 (continued) Analysis: (a) (continued) (b) From the data, take the azeotrope at xB = yB = 0.545 and xCH = yCH = 1 - 0.545 = 0.455, and with = 1.079 and CH = 1.102. To determine the van Laar constants, use Eqs. (2-73) and (2-74) with 1= benzene and 2 = cyclohexane: 2 0.455ln1.102 AB,CH = ln1.079 1 - = 0.3247 0.545ln1.079 ACH,B 0.545 ln1.079 = ln1.102 1 - 0.455 ln1.102 2 = 0.3247 Compute with a spreadsheet values of activity coefficients, using these values for the binary interaction parameters with the van Laar equations, (3), Table 2.9: 1 ln B = 2 1 + 0.3247 xB / 0.3647 xCH ln CH = 1 1 + 0.3647 xCH / 0.3247 xB 2 Note that because the activity coefficients are provided, the vapor pressure data are not needed. Exercise 2.22 Analysis: (b) (continued) Temp., oC 79.7 79.1 78.5 78.0 77.7 77.6 77.6 77.6 77.8 78.0 78.3 78.9 79.5 xB 0.088 0.156 0.231 0.308 0.400 0.470 0.545 0.625 0.701 0.757 0.822 0.891 0.953 Experimental 1.300 1.256 1.219 1.189 1.136 1.108 1.079 1.058 1.039 1.025 1.018 1.005 1.003 CH 1.003 1.008 1.019 1.032 1.056 1.075 1.102 1.138 1.178 1.221 1.263 1.328 1.369 van Laar______ B 1.317 1.271 1.224 1.181 1.136 1.107 1.079 1.054 1.035 1.023 1.013 1.005 1.001 CH 1.002 1.007 1.016 1.030 1.052 1.074 1.102 1.139 1.181 1.218 1.266 1.326 1.387 It is seen that the van Laar equation fits the experimental data quite well. Exercise 2.23 Subject: Activity coefficients from the Wilson equation for the ethanol-benzene system at 45oC. Given: Wilson constants and experimental activity coefficient data. Find: Wilson activity coefficients and comparison with experimental data. Analysis: Let: 1 = ethanol and 2 = benzene The Wilson constants are 12 = 0.124 and 21 = 0.523 From Eqs. (4), Table 2.9, ln 1 = - ln ( x1 + 0.124 x2 ) + x2 0.124 0.523 - x1 + 0.124 x2 x2 + 0.523 x1 0.124 0.523 - x1 + 0.124 x2 x2 + 0.523 x1 ln 2 = - ln ( x2 + 0.523 x1 ) - x1 Using a spreadsheet and noting that = exp(ln ), the following values are obtained, x1 0.0374 0.0972 0.3141 0.5199 0.7087 0.9193 0.9591 1 8.142 5.029 2.032 1.368 1.140 1.000 0.992 Experimental 2 1.022 1.053 1.297 1.715 2.374 3.735 4.055 1 8.182 4.977 2.033 1.370 1.120 1.009 1.002 Wilson_______ 2 1.008 1.044 1.294 1.708 2.350 3.709 4.108 It is seen that the Wilson equation fits the data very well. Exercise 2.23 (continued) Exercise 2.24 Subject: Activity coefficients for the ethanol (1) - isooctane (2) system at 50oC. Given: Infinite-dilution activity coefficients for the liquid phase. Find: (a) (b) (c) (d) (e) splitting van Laar constants Wilson constants Activity coefficients from van Laar and Wilson equations Comparison to azeotropic point y-x curve from van Laar equation to show erroneous prediction of phase Analysis: (a) From van Laar Eqs. (2-72) for infinite dilution, A12 = ln 1 = ln(21.17) = 3.053 A21 = ln = ln(9.84) = 2.286 2 (b) From Wilson Eqs. (2-80) and (2-81) for infinite dilution, ln 1 = 3.053 = 1 - ln 12 - 21 (1) (2) ln = 2.286 = 1 - ln 21 - 12 2 Solving simultaneous, nonlinear Eqs. (1) and (2) using Newton's method, 12 =0.1004 and 21 = 0.2493 (c) Activity coefficients can be calculated with the above constants, using Eqs. (3), Table 2.9 for the van Laar equations and Eqs. (4), Table 2.9 for the Wilson equations. Results from a spreadsheet are as follows: x1 0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.0 1 21.17 10.13 5.56 3.44 2.35 1.75 1.40 1.198 1.079 1.018 1.000 van Laar 2 1.000 1.039 1.154 1.354 1.661 2.11 2.77 3.71 5.06 7.02 9.84 1 21.17 6.63 3.76 2.61 2.00 1.631 1.387 1.219 1.103 1.029 1.000 Wilson______ 2 1.000 1.054 1.162 1.310 1.510 1.784 2.174 2.77 3.74 5.58 9.84 Exercise 2.24 (continued) Analysis: (c) (continued) Note that the Wilson activity coefficients vary more steeply at the infinitedilution ends. (d) At the azeotropic point, x1 = 0.5941 and x2 = 0.4059. Using the van Laar constants from part (a) with Eqs. (3), Table 2.9, van Laar gives 1 = 1.419, compared to 1.44 experimental van Laar gives 2 = 2.72, compared to 2.18 experimental Using the Wilson constants from part (b) with Eqs. (4), Table 2.9, Wilson gives 1 = 1.400, compared to 1.44 experimental Wilson gives 2 = 2.147, compared to 2.18 experimental The Wilson equation is acceptable for both components. The van Laar equation gives poor agreement for isooctane. (e) At 50oC, the vapor pressures are 221 torr for ethanol and 146 torr for isooctane. Thus, system pressure will be low over the entire range of composition. Therefore, the modified Raoult's law K-value expression, given by Eq. (4), Table 2.3 applies. When combined with Eq. (2-19), we obtain the following expression for predicting the y - x curve: x1 1 P1s y1 = (3) P Exercise 2.24 (continued) Analysis: (e) (continued) By Raoult's law, partial pressure is given by pi = xiPis Therefore, the modified Raoult's law gives pi = xi i Pis By Dalton's law, the sum of the partial pressures equals total pressure. Thus, P= i pi = i xi i Pi s (4) Using a spreadsheet with Eqs. (3) and (4) and the van Laar activity coefficients from the table above in part (c), values of y1 are computed for values of x1: x1 P, torr y1 0.0 146 0.000 0.1 361 0.620 0.2 381 0.645 0.3 367 0.622 0.4 356 0.587 0.5 348 0.556 0.6 349 0.553 0.7 348 0.532 0.8 339 0.563 0.9 305 0.663 1.0 221 1.000 The y-x plot exhibits the same characteristics as the system in Fig. 2.20. Therefore, the van Laar equation erroneously predicts phase splitting. Exercise 3.1 Subject: Evaporation of a mixture of ethanol (AL) and ethyl acetate (AC) from a beaker into still air within the beaker. Given: Initial equimolar mixture of AL and AC, evaporating into still air at 0oC and 1 atm. Vapor pressures and diffusivities in air of AL and AC at 0oC. Assumptions: Well-mixed liquid and Raoult's law. Negligible bulk flow effect. Air sweeps across the top of the beaker at a rate such the mole fractions of AL and AC in the air at the top of the beaker are zero. Find: Composition of the remaining liquid when 50% of the initial AL has evaporated. Analysis: All of the mass-transfer resistance is in the still air layer in the beaker, which increases in height, z, as evaporation takes place. Apply Fick's law to both AL and AC with negligible bulk flow effect. Thus, from Eq. (3-16), the molar flux for ethanol through the gas layer in the beaker is as follows, where Di is the diffusivity of component i in air. dc dy N AL = - DAL AL = - DAL c AL dz dz Rearranging, N AL dz = - DAL c Similarly, N AC dz = - DAC c 0 yAL 0 yAC dyAL dyAC (1) (2) Dividing Eq. (1) by (2), N AL DAL yAL 6.45 10 -6 yAL = = (3) N AC DAC yAC 9.29 10 -6 yAC where yAL and yAC are mole fractions in the vapor at the vapor-liquid interface. By material balance, the molar flux of component i is equal to the rate of decrease in moles, ni ,, of component i in the well-mixed liquid in the beaker per unit of mass-transfer area. Thus, dn N AL = - AL (4) Adt dn N AC = - AC (5) Adt By Raoult's law, at the gas-liquid interface, using Eqs. (2-19) and (3) in Table 2.3, Ps Ps nAL 1.62 nAL yAL = AL xAL = AL = (6) P P nAL + nAC 101 nAL + nAC yAC s s PAC PAC nAC 3.23 nAC = xAC = = P P nAL + nAC 101 nAL + nAC (7) Exercise 3.1 (continued) Analysis: (continued) Substituting Eqs. (4) to (7) into (3), and rearranging, dnAL 162 . = dnAC 3.23 9.29 6.45 nAL nAC (8) Integrating Eq. (8) from the start of the evaporation, letting ni0 be the initial values, ln nAL n = 0.722 ln AC nAL0 nAC 0 (9) As a basis, assume 100 moles of original mixture. Thus, nAL0 = 50 mol and nAC 0 = 50 mol When nAL0 = 25 mol, i.e. half of the original, Eq. (9) gives nAC = 19.2 mol Therefore, the mole fractions in the well-mixed liquid when 50% of the AL has evaporated are, xAL = xAC 25 = 0.566 25 + 19.2 19.2 = 0.434 = 25 + 19.2 Exercise 3.2 Subject: Evaporation of benzene (B) at 25oC and 1 atm from an open tank through a stagnant air layer of constant thickness. Given: Tank diameter = 10 ft, with a stagnant gas layer above liquid benzene of 0.2-in. thickness. For benzene, liquid density = 0.877 g/cm3, MW = 78.11, vapor pressure = 100 torr, and the diffusivity in air = 0.08 cm2/s. Assumptions: All mass-transfer resistance is in the thin gas layer of constant thickness. Steadystate with a benzene mole fraction in the air adjacent to the liquid given by Raoult's law, and the benzene mole fraction in the air at the other side of the gas layer equal to zero, assuming the evaporated benzene is continuously swept away as in Example 3.2. Ideal gas. Find: Loss of benzene by evaporation in pounds per day. Analysis: From Eq. (3-35) for unimolecular diffusion, taking into account bulk flow, NB = cDB yBI - 0 z 1 - yB LM (1) where, the total gas concentration, c, is obtained from the ideal gas law, c = P/RT = 1/(82.06)(298) = 4.09 x 10-5 mol/cm3 From Eq. (2-44) for Raoult's law, noting that for pure benzene liquid, xB = 1, yB at the gas-liquid interface = Ps/P = 100/760 = 0.132 The gas film thickness = z = 0.2 in. = 0.508 cm 1 - yB LM = (1 - 0) - 1 - yBI ln 1- 0 1 - y BI = (1 - 0) - 1 - 0.132 = 0.930 1- 0 ln 1 - 0132 . Substituting these values into Eq. (1), NB = 4.09 10 -5 0.08 0132 - 0 . 0.508 0.930 = 9.14 10-7 mol benzene / cm2 - s Cross-sectional area of the tank = mass-transfer area = D2/4 = (3.14)(10)2/4 = 78.5 ft2 = 72,930 cm3 Benzene loss rate = 9.14 x 10-7 (78.11)(72,930)(3600)(24)/454 = 991 lb/day Exercise 3.3 Subject: Countercurrent diffusion of toluene (T) and benzene (B) across vapor film at 170oF (630oR) and 1 atm. Given: Stagnant vapor film of 0.1-inch (0.00833-ft) thickness, containing 30 mol% toluene and 70 mol% benzene, in contact with liquid reflux containing 40 mol% toluene and 60 mol% benzene. Diffusivity of toluene in benzene = 0.2 ft2/h. Vapor pressure of toluene = 400 torr. Assumptions: Equal molar heats of vaporization for benzene and toluene, such that diffusion is equimolar, countercurrent. Ideal gas law and Raoult's law apply. All mass transfer resistance is in the vapor phase, i.e. liquid is assumed to be uniform in composition. Given vapor composition is for bulk conditions. Phase equilibrium at the vapor-liquid interface. Find: Mass transfer rate of toluene in lbmol/h-ft2. Analysis: At the vapor-liquid interface, use Raoult's law, Eq. (2-44). Then, the mole fraction of toluene at the interface is, Ps 400 yTI = xT T = 0.4 = 0.211 P 760 From ideal gas law, total gas concentration, c, is P/RT = 1/(0.7302)(630) = 0.00217 lbmol/ft3 From the finite-difference form of Fick's law, Eq. (3-16), for the diffusion of toluene from the bulk vapor to the vapor-liquid interface, NT = cDT,B ( yT - yTI ) = ( 0.00217 )( 0.2 )( 0.300 - 0.211) 0.00833 z = 0.00464 lbmol/h-ft 2 Benzene diffuses at the same rate in the opposite direction. Exercise 3.4 Subject: Drop in level of water (W) contained in a vertical tube when evaporating into air at 25oC (537oR). Given: Tube with an inside diameter of 0.83 inch. Initial liquid level of water in tube = 0.5 inch from the top. Air above the tube has a dew point of 0oC. Diffusivity of water vapor in air = 0.256 cm2/s or 0.992 ft2/h. Assumptions: Pressure = 1 atm. Ideal gas. Phase equilibrium at the gas-liquid interface with Raoult's law for mole fraction of water in the vapor adjacent to liquid water. Find: (a) Time for the liquid level to drop from 0.5 inch to 3.5 inches. (b) Plot of liquid level as a function of time. Analysis: (a) The mole fraction of water in the air adjacent to the gas-liquid interface is obtained from Raoult's law, Eq. (2-44), with xW = 1 for pure liquid water, using a vapor pressure of 0.45 psia for water at 25oC. P s 0.45 yWI = W = = 0.0306 P 14.7 In the bulk air, the mole fraction of water is obtained from the dew-point condition. Thus, the partial pressure of water vapor = vapor pressure of water at 0oC = 0.085 psia. Therefore, p P s 0.085 yW = W = W = = 0.00578 P P 14.7 The equation for the time, t, for the water level to drop from level z1 = 0.5 inch (0.0417 ft) to level z2 equal to as large as 3.5 inches (0.2917 ft) is derived in Example 3.2, where the result is given by Eq. (6). Applying that equation here, with = 62.4 lb/ft3 for liquid water, total gas concentration, c, by the ideal gas law to give c=P/RT = 1/(0.7302)(537) = 0.00255 lbmol/ft3, and a bulk flow factor = (1 - xW)LM = to a good approximation to the arithmetic average = [(1 - 0.0306) + (1 - 0.00578)]/2 = 0.982. t= L 1 - xW LM 2 z2 - z12 = M W cDW yWI - yW 62.4 0.982 2 z2 - z12 (18.02)(0.00255)(0.992)(0.0306 - 0.00578) (1) 2 2 = 54,200 z2 - z12 = 54,200 z2 - 0.0417 2 When z2 = 3.5 inches = 0.2917 ft, Eq. (1) gives t = 4,520 h Exercise 3.4 (continued) Analysis: (continued) (b) For other values of z2 , the following results are obtained using a spreadsheet with Eq. (1) above: z2 , inches 0.5 1.0 1.5 2.0 2.5 3.0 3.5 z2 , feet 0.0417 0.0833 0.1250 0.1667 0.2083 0.2500 0.2917 time, hours 0 282 753 1,412 2,259 3,294 4,520 Exercise 3.5 Subject: Mixing of argon (A) and xenon (X) by molecular diffusion. Given: Bulb 1 containing argon. Bulb 2 containing xenon. Bulbs connected by a 0.002 m (0.2 cm) inside diameter by 0.2 m (20 cm) long tube. 105oC (378 K) and 1 atm are maintained. At time, t, equal 0, diffusion is allowed to occur through the connecting tube. Diffusivity = 0.180 cm2/s. Assumptions: Gas in each bulb is perfectly mixed. The only mass transfer resistance is in the connecting tube. Ideal gas law. Equimolar, countercurrent diffusion. Find: At a time when the argon mole fraction = 0.75 in one bulb and 0.20 in the other bulb, determine, (a) Rates and directions of mass transfer for A and X. (b) Transport velocities of A and X. (c) Molar average velocity of the gas mixture. Analysis: This exercise is similar to Example 3.1. Area normal to diffusion = A = 3.14(0.2)2/4 = 0.0314 cm2 From the ideal gas law, total concentration of gas = c = P/RT = 1/(82.06)(378) = 0.0000322 mol/cm3 (a) From form of Eq. (3-18), nA = cDA,X A ( yA1 - yA 2 ) -9 = 5 10 z mol/s from bulb 1 to 2 = ( 0.0000322 )( 0.180 )( 0.0314 )( 0.75 - 0.20 ) 20 nX = 5 10-9 mol/s from bulb 2 to 1 (b) Because of equimolar, countercurrent diffusion, species velocities relative to stationary coordinates are equal to diffusion velocities. From Eq. (3-9), for argon (A), A = A X = D JA nA 5 10-9 0.00495 = = = = cA AcyA (0.314)(0.0000322)( yA ) yA (2) (1) 0.00495 yX Thus, the transport velocities depend on the mole fractions. Exercise 3.5 (continued) Analysis: (b) (continued) Using Eqs. (1) and (2) over the range of mole fractions, noting that mole fractions are linear with distance because of equimolar, countercurrent diffusion, Distance from Bulb 1, cm 0 (Bulb 1) 5 10 15 20 (Bulb 2) yA 0.7500 0.6125 0.4750 0.3375 0.2000 yX 0.2500 0.3875 0.5250 0.6625 0.8000 , cm/s 0.0066 0.0081 0.0104 0.0147 0.0248 , cm/s 0.7500 0.6111 0.4760 0.3367 0.1996 (c) Because we have equimolar, countercurrent diffusion, the molar average velocity of the mixture is zero. Exercise 3.6 Subject: Measurement of diffusivity of toluene (T) in air (A), and comparison with prediction. Given: Vertical tube, 3 mm in diameter and open at the top, containing toluene at 39.4oC (312.6 K). Initially, the toluene level, z1, is 1.9 cm below the top. It takes 960,000 s for the level to drop to z2 equal to 7.9 cm below the top. The toluene vaporizes into 1-atm (760 torr) air, which is stagnant inside the tube. The density of toluene is 0.852 g/cm2. Its vapor pressure is 57.3 torr. Assumptions: Isothermal vaporization. Mass transfer resistance only in the air in the tube. Neglect counterdiffusion of air. Molecular diffusion of toluene through the air in the tube. Toluene mole fraction of zero in the air at the top of the tube. Phase equilibrium at the gas-liquid interface, given by Raoult's law. Ideal gas law. Find: Experimental and predicted values of diffusivity of toluene in air at 39.4oC and 1 atm. Analysis: Applying Raoult's law, Eq. (2-44), to the gas-liquid interface, yT = xT PTs 57.3 = 10 . = 0.0754 P 760 Therefore, the driving force for molecular diffusion, yT = 0.0754 - 0.0 = 0.0754 Total gas concentration from ideal gas law, c = P/RT = 1/(82.06)(312.6) = 3.9 x 10-5 mol/cm3 Molecular weight of toluene = ML = 92.1 The correction for bulk flow = (1-xA)LM = approximately [(1-0) + (1-0.0754)]/2 = 0.962 Integrating and solving Eq. (6) in Example 3.2 for the experimental gas diffusivity, DT,A = 2 L (1 - xA ) LM z2 - z12 tM L c ( yT ) 2 = (0.852)(0.962) 7.92 - 1.92 = 0.0927 cm 2 /s -5 (960, 000)(92.1)(3.9 10 )(0.0754) 2 Compute the predicted value from Eq. (3-36) of Fuller, Schettler, and Giddings, with: M T,A = 2 = 44.1 1 1 + 92.1 29 V T From Table 3.1, V A = 19.7, = 7(15.9) + 8(2.31) - 18.3 = 1115 . DT,A = 0.00143(312.6)1.75 = 0.0887 cm 2 /s (1)(44.1)1/ 2 [(19.7)1/ 3 + (111.5)1/ 3 ]2 The predicted value is 4.3% less than the measured value. This is good agreement. Exercise 3.7 Subject: Countercurrent molecular diffusion of hydrogen (H) and nitrogen (N) in a tube. Given: Tube of 1 mm (0.1 cm) inside diameter and 6 inches (15.24 cm) long. At one end of the tube (1) is pure hydrogen (H) blowing past. At the other end of the tube (2) is pure nitrogen blowing past. The temperature is 75oC (348 K) and pressure is 1 atm. Assumptions: Ideal gas law. Find: (a) Estimate the diffusivity and the rate of diffusion of nitrogen in mol/s for equimolar, countercurrent diffusion. (c) Plot of hydrogen mole fraction with distance. Analysis: (a) Estimate diffusivity from (3-36), with T =348 K, P = 1 atm, M H,N = 2 1 1 + 2.02 28 V N = 3.765 From Table 3.1, V H = 612, . = 18.5 DH,N = 0.00143(348)1.75 = 1.033 cm 2 /s 1/ 2 1/ 3 1/ 3 2 (1)(3.765) [(6.12) + (18.5) ] From Eq. (3-18), Fick's law for equimolar, countercurrent diffusion, with, c = P/RT = 1/(82.06)(348) = 0.000035 mol/cm3 yN = mole fraction driving force for nitrogen = 1 - 0 = 1 z = distance for diffusion = 15.24 cm A = cross-sectional area for diffusion = (3.14)(0.1)2/4 = 0.00785 cm2 Nitrogen flow rate, nN = J N A = cDH,N (yN ) z A= (0.000035)(1.033)(1.0) (0.00785) = 1.86 10-8 mol/s 15.24 (c) The mole fraction of hydrogen varies linearly through the tube because of equimolar, countercurrent diffusion. Thus, Exercise 3.7 (continued) Distance from End 1, inches Hydrogen mole fraction 0 1 1 0.833 2 0.667 3 0.500 4 0.333 5 0.167 6 0 1 0.9 0.8 Hydrogen mole fraction 0.7 0.6 0.5 0.4 0.3 0.2 0.1 0 0 0.5 1 1.5 2 2.5 3 3.5 4 4.5 5 Distance from End 1, inches Exercise 3.8 Subject: Molecular diffusion of HCl (H) across an air (A) film at 20oC. Given: Air film of 0.1-inch (0.254 cm) thickness. HCl partial pressure of 0.08 atm on side 1 of the film and 0 on the other side 2. Diffusivity of HCl in air at 20oC (293 K) and 1 atm = 0.145 cm2/s. Assumptions: Ideal gas law. Unimolecular diffusion of HCl. Find: Diffusion flux of HCl in mol/s-cm2 for the following total pressures: (a) 10 atm (b) 1 atm (c) 0.1 atm Analysis: Fick's law applies. Use the form of Eq. (3-33). Note that for an ideal gas, product cDH,A is independent of total pressure because c is directly proportional to P, while from Eq. (3-36), cDH,A is inversely proportional to P. At P = 1 atm, c = P/RT = 1/(82.06)(293) = 4.16 x 10-5 mol/cm3. Therefore, cDH,A = (4.16 x 10-5)(0.145) = 6.03 x 10-6 (a) P = 10 atm. By Dalton's law, at (yH)1 = 0.08/10 = 0.008 From Eq. (3-33), NH = cDH,A z ln 1 - ( yH ) 2 (6.03 10 -6 ) 1- 0 = ln = 1.91 10-7 mol HCl/s-cm 2 1 - ( yH )1 1 - 0.008 ( 0.254) ) (b) For P = 1 atm, (yH)1 = 0.08/1 = 0.08, which gives, NH = 1.98 x 10-6 mol/s-cm2 (c) For P = 0.1 atm, (yH)1 = 0.08/0.1 = 0.8, which gives, NH =3.82 x 10-5 mol/s-cm2 Exercise 3.9 Subject: Estimation of the binary gas diffusivity for nitrogen (A) - toluene (B) at 25o C (298 K) and 3 atm Assumptions: No need to correct diffusivity for high pressure with Takahashi method. Find: Binary gas diffusivity using the method of Fuller, Shettler, and Giddings. Analysis: Use Eq. (3-36), with, M A,B = From Table 3.1, 2 1 1 + 28 92 = 42.9 VA = 18.5, DA,B = VB = 7(15.9) + 8(2.31) - 18.3 = 1115 . 0.00143(298)1.75 = 0.028 cm 2 / s (3)(42.9)1/ 2 [18.51/ 3 + 111.51/ 3 ]2 Exercise 3.10 Subject: Correction of gas binary diffusivity for high pressure. Given: Results of Example 3.3 for oxygen-benzene system at 38oC (311 K) and 2 atm, which give, DAB = 0.0494 cm2/s Find: Diffusivity at 100 atm. Analysis: If Eq. (3-36) is applied, DAB = 0.0494 (3/100) = 0.00148 cm2/s Apply the Takahashi correlation of Fig. 3.3, based on reduced T and P. For equimolar mixture, Tr=T/Tc and Pr=P/Pc where, Tc = 0.5(154 + 563) = 359 K and Pc = 0.5(48.6+49.7) = 49.1 atm Therefore, Tr = 311/359 = 0.866, and Pr = 100/49.1 = 2.04 We are outside the range of the Takahashi correlation, but it appears that the correction would greatly decrease the diffusivity, by a factor of 10 or more. Exercise 3.11 Subject: Estimation of infinite-dilution liquid diffusivity for carbon tetrachloride at 25oC (298 K) in four different solvents. Given: Experimental values diffusivity for solvents of (a) methanol, (b) ethanol, (c) benzene, and (d) n-hexane. Find: Diffusivities by the methods of Wilke and Chang (W-C), and of Hayduk and Minhas (H-M). Compare predicted values to given experimental values. Analysis: Let:A = the solute, CCl4; and B = solvent. The Wilke-Chang equation, Eq. (3-39), is DA,B 7.4 10 -8 B M B = 0 B A.6 1/ 2 T (1) From Table 3.3, A = 14.8 + 4(21.6) =101.2 cm3/mol Using Eq. (1) with the following parameters, values of diffusivity are computed. Solvent, B Methanol Ethanol Benzene n-Hexane MB 32 46 78 86 1.9 1.5 1.0 1.0 B , cP 0.57 1.17 0.60 0.32 DA,B (W-C) 1.89 x 10-5 0.98 x 10-5 2.03 x 10-5 4.00 x 10-5 DA,B (Expt.) 1.69 x 10-5 1.50 x 10-5 1.92 x 10-5 3.70 x 10-5 Except for ethanol, the Wilke-Chang equation makes good predictions. The applicable Hayduk-Minhas equation, Eq. (3-42), is, (2) 0.92 0.23 B B where, for the nonpolar solute, CCl4 , with methanol and ethanol solvents, both P B and B must be multipled by 8B in cP. Parachors for methanol and benzene are obtained from Table 3.5. Parachors for ethanol, n-Hexane, and carbon tetrachloride are obtained by structural contributions from Table 3.6. Using Eq. (2) with P A= 229.8 and the following parameters, values of diffusivity are computed. Solvent, B PB B B , cP DA,B (H-M) DA,B (Expt.) Methanol 88.8 37.0 0.57 2.55 x 10-5 1.69 x 10-5 Ethanol 125.3 59.2 1.17 1.70 x 10-5 1.50 x 10-5 -5 Benzene 205.3 96.0 0.60 1.96 x 10 1.92 x 10-5 n-Hexane 271.0 140.6 0.32 3.05 x 10-5 3.70 x 10-5 Except for methanol and n-hexane, the predictions by the Hayduk-Minhas equation are good. DA,B = 155 10 -8 . T 1.29 PB0.5 / PA0.42 Exercise 3.12 Subject: Estimation of the diffusivity of benzene (A) at infinite dilution in formic acid (B) at 25oC and comparison to the experimental value for B at infinite dilution in A. Given: Experimental value of 2.28 x 10-5 cm2/s for the diffusivity of B at infinite dilution in A. Find: Diffusivity by the Hayduk-Minhas equation, (3-42). Analysis: The applicable H-M equation is, DA,B = 155 10 . -8 T 1.29 PB0.5 / PA0.42 0.92 0.23 B B (1) From Table 3.5, parachors for A and B are 205.3 and 93.7, respectively. From Table 3.3, The molar volume of the solvent, B, is 2(3.7)+14.8+7.4+12.0 = 41.3 cm3/mol. From Perry's Handbook, the viscosity of B at 25oC is 1.7 cP. Using Eq. (1), DA,B = 1.55 10 -8 2981.29 ( 93.7 0.50 / 205.30.42 ) (1.7) 0.92 (41.3) 0.23 = 6.5 10-6 cm 2 /s This is significantly less than the experimental value of 2.28 x 10-5 and the predicted value of 2.15 x 10-5 cm2/s for the diffusivity of formic acid at infinite dilution in benzene. Exercise 3.13 Subject: Estimation of the liquid diffusivity of acetic acid at 25oC (298 K) in dilute solutions of benzene, acetone, ethyl acetate, and water, and comparison with experimental values. Given: Experimental values. Find: Infinite dilution diffusivities for acetic acid (A) in the four solvents using the appropriate equation. Analysis: For benzene, acetone, and ethyl acetate, Eq. (3-42) of Hayduk-Minhas applies: DA,B = 155 10 . -8 T 1.29 PB0.5 / PA0.42 0.92 0.23 B B (1) From Table 3.5, the parachor for acetic acid = P A = 131.2. For all three solvents, this value is multiplied by 2 to give 262.4. Parachors for benzene and acetone are obtained from Table 3.5, and are listed below. The parachors for ethyl acetate, obtained by structural contributions from Table 3.6, is 2(55.5) + 40.0 + 63.8 = 214.8. Molecular volumes are obtained from Table 3.3, and are listed below. Solvent viscosities are obtained from Perry's Handbook. The results of applying Eq. (1) are as follows Solvent, B Benzene Acetone Ethyl acetate PB 205.3 161.5 214.4 B 96 74 106.1 B , cP 0.60 0.32 0.45 DA,B (H-M) 1.5 x 10-5 2.5 x 10-5 1.93 x 10-5 DA,B (Expt.) 2.09 x 10-5 2.92 x 10-5 2.18 x 10-5 The predictions by the Hayduk-Minhas equation are low, but quite good for ethyl acetate. For water, the Wilke-Chang equation, Eq. (3-39), is most applicable, DA,B 7.4 10 -8 B M B = 0 B A.6 1/ 2 T (2) From Table 3.3, A = (2)14.8 + 4(3.7) + 7.4 + 12 = 63.8 cm3/mol. Solvent viscosity = 0.95 cP, MB = 18, and B = 2.6. Substitution into Eq. (1) gives, DA,B 7.4 10-8 ( 2.6 18 ) = 0.95(63.8)0.6 1/ 2 298 = 1.3 10 -5 cm 2 /s This is close to the experimental value of 1.19 x 10-5 cm2/s. Exercise 3.14 Subject: Vapor diffusion through an effective film thickness. Given: Water (A) at 11oC (284 K) evaporating into dry air (B) at 25oC (298 K) at the rate of 0.04 g/h-cm2. Assumptions: 1 atm pressure. Ideal gas law. Raoult's law for water vapor mole fraction. Find: Effective stagnant air film thickness, assuming molecular diffusion of the water vapor through the air film. Analysis: Apply Eq. (3-31) for Fick's law with the bulk flow effect. Solving for film thickness, cDA,B ln z = 1 - yA b (1) 1 - yA I NA From the ideal gas law, using T = (284 + 298)/2 = 291 K, c = P/RT = 1/(82.06)(284) = 4.29 x 10-5 mol/cm3. Molar flux of water vapor = NA = 0.04/MA = 0.04/(18)(3600) = 6.17 x 10-7 mol/s-cm2. Mole fraction of water vapor in the bulk dry air = yAb = 0 Mole fraction of water vapor at the interface, yA I = Ps at 11oC/P = 0.191/14.7 = 0.013. Estimate diffusivity of water vapor in air at 291 K and 1 atm from Eq. (3-36) of Fuller, Schettler, and Giddings, with, 2 M A,B = = 22.2 1 1 + 18 29 From Table 3.1, VA = 131, . VB = 19.7 DA,B = Substitution into Eq. (1), z = 0.00143(291)1.75 = 0.24 cm2 / s (1)(22.2)1/ 2 [1311/ 3 + 19.71/ 3 ]2 . 1- 0 1 - 0.013 4.29 10-5 ( 0.24 ) ln 6.17 10-7 = 0.218 cm Note that the bulk flow has little effect here. Exercise 3.15 Subject: Mass transfer of isopropyl alcohol (A) by molecular diffusion through liquid water and gaseous nitrogen at 35oC (308 K) and 2 atm Given: Critical conditions for nitrogen and isopropyl alcohol and molar liquid volume of isopropyl alcohol. Assumptions: Ideal gas law. Find: (a) (b) (c) (d) (e) (f) (g) Diffusivity of A in liquid water by Wilke-Chang equation. Diffusivity of A in gaseous nitrogen by the Fuller, Schettler, Giddings equation. The product, DAB = DAB c for part (a). The product, DAB = DAB c for part (b). Comparison of diffusivities in parts (a) and (b). Comparison of results from parts (c) and (d). Conclusions about diffusion in the liquid versus the vapor phase. Analysis: (a) For the Wilke-Chang equation (3-39), use B = 2.6 and MB =18. From Perry's Handbook, B = 0.78 cP. From Table 3.3, = 3(14.8) + 8(3.7) + 7.4 = 81.4 From Eq. (3-39), DA,B 7.4 10-8 ( 2.6 18 ) = 0.78(81.4)0.6 1/ 2 308 = 1.43 10-5 cm 2 /s (b) For the Fuller-Schettler-Giddings Eq. (3-36), use T = 308K and P = 2 atm, with, 2 M A,B = = 38.3 1 1 + 60.4 28 From Table 3.1, VA = 3(15.9) + 8(2.31) + 611 = 72.3 . VB = 18.5 DA,B = 0.00143(308)1.75 = 0.056 cm 2 / s 1/ 2 1/ 3 1/ 3 2 (2)(38.3) [72.3 + 18.5 ] Exercise 3.15 (continued) Analysis: (continued) (c) The molar density of liquid water = /M = 1/18 = 0.056 mol/cm3. Therefore, DAB = 1.43 x 10-5 (0.056) = 8.01 x 10-7 mol/s-cm (d) From the ideal gas law, the molar density of gaseous nitrogen = P/RT = 2/(82.06)(308) = 7.91 x 10-5 mol/cm3. Therefore, DAB = 0.056 (7.91 x 10-5) = 4.43 x 10-6 mol/s-cm (e) The diffusivity in the gas is about 3 orders of magnitude more than in the liquid. (f) The product of diffusivity and molar density in the gas is less than one order of magnitude more than in the liquid. (g) For equal mole fraction gradients, diffusion through the liquid phase is comparable to that in the gas phase. Exercise 3.16 Subject: Liquid diffusivities for the ethanol (A) -benzene (B) system at 45oC (318 K) over the entire composition range. Given: Experimental activity coefficients in Exercise 2.23. Find: Effect of composition on diffusivities of both ethanol and benzene. Analysis: Use Eq. (3-42) of Hayduk-Minhas to estimate the infinite-dilution liquid diffusivities, with liquid viscosities from Perry's Handbook. From Table 3.5, the parachor of benzene is given. The parachor for ethanol is estimated from Table 3.6. Table 3.3 is used to estimate molecular volumes. The resulting parameters are as follows: P Component , cP Benzene 205.3 96 0.48 Ethanol 125.3 59.2 0.79 For benzene at infinite dilution in ethanol, values of molecular volume and the parachor for ethanol must be multiplied by 8 times the viscosity of ethanol. Thus, from Eq. (3-42), using P = 8(0.79)(125.3) = 792 and = 8(0.79)(59.2) = 374 for ethanol, the solvent in this case, DB,A = 155 10-8 . 3181.29 792 0.5 / 205.30.42 0.79 0.92 374 0.23 = 2.51 10 -5 cm2 / s For ethanol at infinite dilution in benzene, Eq. (3-42) gives, DA,B = 155 10 . -8 3181.29 205.30.5 / 125.30.42 0.48 0.92 96 0.23 = 3.4 10 -5 cm2 / s Use Eqs. (3-45) and (3-46) of Vignes to compute DA,B and DB,A as a function of composition, DA,B = DA,B DB,A = DB,A xB DB,A DA,B xA 1+ 1+ ln A ln xA ln B ln xB (1) T ,P xA xB (2) T ,P The partial derivatives are evaluated numerically with, ln A ln A = ln xA ln xA and ln B ln B = ln xB ln xB Exercise 3.16 (continued) Analysis: (continued) Using the data from Exercise 2.23, xA DA,B , DB,A , ln A ln B ln ln B cm2/s cm2/s ln xA ln xB 0.0000 3.4x10-5 0.0374 2.0937 0.0220 0.0673 -0.501 -0.466 1.66x10-5 1.78x10-5 0.0972 1.6153 0.0519 0.2057 -0.773 -0.757 0.727x10-5 0.777x10-5 0.3141 0.7090 0.2599 0.4170 -0.785 -0.783 0.645x10-5 0.651x10-5 0.5199 0.3136 0.5392 0.949x10-5 0.985x10-5 0.6143 -0.664 -0.651 0.1079 0.8645 0.7087 -0.414 -0.353 1.56x10-5 1.72x10-5 0.8140 0.9193 0.0002 1.3177 0.9392 -0.186 -0.121 2.08x10-5 2.25x10-5 0.9591 -0.0077 1.3999 1.0000 2.51x10-5 These results give the following plot showing that the liquid diffusivities do not vary linearly with composition. Exercise 3.17 Subject: Estimation of the diffusivity of an electrolyte. Given: 1-M aqueous NaOH at 25oC (298 K). Assumptions: Dilute solution. Find: Diffusivity of the Na+ (A) and OH- (B) ions. Analysis: Eq. (3-47) of Nernst and Haskell applies, with n+ = 1 and n- = 1. From Table 3.7, + = 50.1 for Na+ , and - = 197.6 for OH-. From Eq. (3-47), RT DA,B = F2 1 1 + n+ n- 1 8.314 ( 298 ) 96,500 2 1 1 + 1 1 1 1 + 50.1 197.6 + + 1 = = 2.1 10-5 cm 2 /s - Note that DA,B may be 10 to 20% higher for 1-M solution. Exercise 3.18 Subject: Estimation of the diffusivity of an electrolyte and comparison with experiment. Given: Experimental value of 1.28 x 10-5 cm2/s for 2 M aqueous NaCl at 18oC (291 K). Assumptions: Dilute solution Find: Diffusivity of Na+ (A) and Cl- (B) ions. Analysis: Eq. (3-47) of Nernst and Haskell applies, with n+ = 1 and n- = 1. From Table 3.7, for 25oC, + = 50.1 for Na+ , and - = 76.3 for Cl-. Correction to ionic conductances for 18oC = T/334 water = 291/(334)(1.05) = 0.83 1 1 1 1 RT + 8.314 ( 291) + n+ n- 1 1 DA,B = = = 1.3 10-5 cm 2 /s 1 1 1 1 96,500 2 + F2 + 50.1(0.83) 76.3(0.83) + - This value compares very well with the experimental value. Exercise 3.19 Subject: Estimation of diffusivity of N2 (A) in H2 (B) in pores of a solid catalyst. Given: Catalyst at 300oC (573 K) and 20 atm, with porosity of 0.45 and tortuosity of 2.5. Assumptions: Only mass transfer mechanism is ordinary molecular diffusion. Correction for high pressure is not necessary because of high temperature. Find:. Diffusivity Analysis: Use Eq. (3-49), with = 0.45 and =2.5. Estimate diffusivity of nitrogen in hydrogen at 573 K and 20 atm from Eq. (3-36) of Fuller, Schettler, and Giddings, with, M A,B = From Table 3.1, 2 1 1 + 28 2.016 = 3.76 VA = 18.5, DA,B = From Eq. (3-49), Deff = VB = 612 . 0.00143(573)1.75 = 0124 cm2 / s . 1/ 2 1/ 3 1/ 3 2 (20)(3.76) [18.5 + 612 ] . DA,B (0.45)(0.124) = = 0.022 cm 2 / s 2.5 Exercise 3.20 Subject: Diffusion of hydrogen through the steel wall of a spherical pressure vessel. Given: Gaseous hydrogen (A) stored at 150 psia and 80oF in a 4-inch inside diameter spherical pressure vessel of steel, with a 0.125-inch wall thickness. Solubility of hydrogen in steel at these conditions = 0.094 lbmol/ft3. Diffusivity of hydrogen in steel at these conditions = 3.0 x 10-9 cm2/s. Air outside the vessel with zero partial pressure of hydrogen. Assumptions: Henry's law for solubility of hydrogen in steel. Ideal gas law. Find:. (a) Initial rate of mass transfer of hydrogen. (b) Initial rate of pressure decrease inside the vessel. (c) Time in hours for the pressure inside the vessel to decrease to 50 psia at constant temperature of 80oF. Analysis: As the hydrogen diffuses through the wall, the pressure in the vessel will decrease. (a) Assume finite-difference form of Fick's law with no bulk flow effect. For a spherical shell, Eq. (3-62) gives, nA = DA,B AGM cA1 - cA 2 r2 - r1 (1) DA,B = 3 10 -9 (3600) . = 116 10-9 ft 2 / h 2 (2.54 12) r2 = (2 + 0.125)/12 = 0.1771 ft, r1 = 2/12 = 0.1667 ft 1/2 2 AGM = (A2A1) , A2 = 4(r2) = 4(3.14)(0.1771)2 = 0.394 ft2 , A1 = 4(r1)2 = 4(3.14)(0.1667)2 = 0.349 ft2 , AGM = (0.394 x 0.349)1/2 = 0.371 ft2 cA 1 = 0.094 lbmol / ft 3 , cA 2 = 0 Substituting into Eq. (1), nA = 11.6 10 -9 ( 0.371) (b) ( 0.1771 - 0.1667 ) ( 0.094 - 0 ) = 3.89 10 -8 lbmol/h Let m = number of moles of hydrogen in the vessel at any time. By the ideal gas law, m = PV/RT (2) T = 540oR , V = D3/6 = 3.14(4/12)3/6 = 0.233 ft3 Initially, P = 150 psia and m = (150)(0.233)/(10.73)(540) = 0.00603 mol Exercise 3.20 (continued) Analysis: (b) (continued) Differentiating Eq. (2) with respect to time to determine the rate of change of pressure due to the decrease in the number of moles of hydrogen as it leaves the vessel by diffusion through the walls, dm V dP =n=- dt RT dt (3) Solving Eq. (3) for the rate of change of pressure, the initial rate of change is, dP nRT (3.89 10-8 )(10.73)(540) = -0.00097 psia/h =- =- dt V 0.233 (c) As the pressure in the vessel decreases, the solubility of hydrogen in steel and, therefore, the rate of diffusion of hydrogen through the steel will decrease. Assume Henry's law for the solubility. Then, cA 1 = 0.094 P 150 (4) From Eqs. (1), (3), and (4), the rate of diffusion of hydrogen as a function of pressure is, n = 3.89 10 -8 P V dP (0.233) dP dP =- =- = -0.000040 150 RT dt (10.73)(540) dt dt (5) Integrating Eq. (5) from P = 150 psia to 50 psia, t 389 10 -8 . dt = - 150(0.000040) 0 50 150 dP 150 = 10986 = ln . P 50 Therefore, t = 154,000 (1.0986) = 169,500 h , a very long time! Exercise 3.21 Subject: Mass transfer of gases through a dense polymer membrane Given: A 0.8-micron thick polyisoprene membrane. Partial pressures of methane and hydrogen on the two sides of the membrane. Membrane properties in Table 14.9. Assumptions: Applicability of the solution-diffusion model. Find:. Mass-transfer fluxes for methane and hydrogen Analysis: The mass-transfer flux of a gas species is given by Eq. (3-52), Ni = Hi Di pi z (1) Membrane thickness = 0.8 micron = 0.8 x 10-6 m Values of Di and Hi = Si are given in Table 14.9 and are used with the given partial pressures in Eq. (1) to give the following results: Species Methane Hydrogen p, Pa 2.45 x 106 1.80 x 106 H, mol/m3-Pa 1.14 x 10-4 0.17 x 10-4 D, m2/s 8.0 x 10-11 109 x 10-11 Flux, N, mol/m2-s 0.028 0.042 Exercise 3.22 Subject: Diffusion of NaCl into stagnant water at 25oC (298 K). Given: An initial 0.10-inch (0.254-cm) thickness of NaCl sitting below a 3-ft (91.44 cm) depth of water. Solubility of NaCl in water of 36 g NaCl/100 g water. Diffusivity of NaCl in water = 1.2 x 10-5 cm2/s. Assumptions: Water acts as a semi-infinite slab. Equilibrium at salt-water interface. Find:. Time and concentration profile of salt in the water, when: (a) 10% of the salt has dissolved. (b) 50% of the salt has dissolved. (c) 90% of the salt has dissolved. Analysis: Basis: area for mass transfer = 1 cm2 = A Volume of salt = 1(0.254) = 0.254 cm3 From Perry's Handbook, density of NaCl = 2.165 g/cm3 Mass of salt = 2.165 (0.254) = 0.550 g Molecular weight of NaCl = 58.44 g/mol Moles of salt = 0.550/58.44 = 0.00941 mol/cm3 At the solid-liquid interface, NaCl concentration = 36/(100 + 36) x 100% = 26.5 wt% From Perry's Handbook, density of this solution at 25oC = 1.2 g/cm3 or cs = 0.265 (1.2)/58.44 = 0.00544 mol NaCl/cm3 Initial concentration of salt in the bulk of the water = co = 0 For diffusion into a semi-infinite slab, Eq. (3-79) applies, for the number of moles of NaCl, N, transferred in time t. Solving that equation for time t gives, t= 4 A ( cs - co ) DNaCl 2 2 N 2 = 4(1) ( 0.00544 - 0 ) (1.2 10 2 2 N 2 (3.14) -5 ) = 2.21 109 N 2 s (1) Using Eq. (1), the following times are computed: Part (a) 10% transferred (b) 50% transferred (c) 90% transferred Moles transferred 0.000941 0.004705 0.008469 Time, s 1,960 48,900 159,000 Time, h 0.544 13.6 44.0 Use Eq. (3-75) to compute concentration profiles, where z = 0 at the salt-water interface, z z = 0.00544 erfc = 0.00544 erfc( x ) 2 DNaCl t 2 1.2 10 -5 t c = cs erfc (2) Exercise 3.22 (continued) Analysis: (continued) Solving Eq. (2) for values of x, which fixes the value of c, the following results are obtained: z, cm for value of c______ x 0.5 1.0 1.5 2.0 erfc (x) 0.4795 0.1573 0.0339 0.0047 c, mol/cm3 0.00261 0.000856 0.000184 0.0000256 10% transf. 0.157 0.313 0.470 0.627 50% transf. 90% transf. 0.782 1.408 1.565 2.816 1.173 4.224 3.130 5.632 Since the water layer is 91.44 cm thick, we see that the penetration of the salt is far less than this, such that the assumption of a semi-infinite slab is valid. Exercise 3.23 Subject: Diffusion of water into 4-inch (10.16-cm) thick wood, with sealed edges, from air of 40% relative humidity. Given: Equilibrium moisture content for air conditions = 10 lb water/100 lb of dry wood. Diffusivity of water in wood = 8.3 x 10-6 cm2/s. Initially, wood is dry. Find:. Time for water to penetrate to the center of the wood Analysis: To determine the time, apply the semi-infinite slab solution of Eq. (3-75). Consider center concentrations of moisture equal to 1%, 0.1%, and 0.01% of the equilibrium surface concentration. From Eq. (3-75), = ccenter - co ccenter z 5.08 = = erfc = erfc = erfc x cs - co cs 2 Dwater t 2 8.3 10-6 t (1) Using Eq. (1), the following results are obtained, 0.01 0.001 0.0001 x 1.82 2.33 2.75 t, s 234,000 144,000 104,400 Time, h 65 40 29 If the penetration depth is defined as x = 2, then time = 54 h. Regardless, it takes a long time Exercise 3.24 Subject: Determination of moisture diffusivity in clay brick from measurements of drying rate. Given: Brick dimensions of 2 in (5.08 cm) by 4 in (10.16 cm) by 6 in (15.24 cm). Uniform initial moisture content of 12 wt%. Equilibrium surface moisture content maintained at 2 wt%. After 5 h, average moisture content is 8 wt%. Assumptions: Diffusivity is uniform in all directions. Find: (a) Diffusivity of moisture in clay. (b) Additional time for average moisture content to reach 4 wt%. Analysis: (a) By definition in Eq. (3-80), the unaccomplished fractional moisture concentration change after 5 hours = Eavg = (cs - cavg)/( cs - co) = (2-8)/(2-12) = 0.60. Apply Newman's method for a rectangular slab, given by Eq. (3-86). Thus, Eavg = Ex Ey Ez (1) where the three E values are average values. These values are plotted in Fig. 3.9 as a function of the Fourier number for mass transfer, using half dimensions, a, b, and c in the x, y, and z directions. Thus, using units of cm and seconds, the three Fourier numbers are: Dt/a2 = D (5 x 3600)/(2.54)2 = 349 D to get Ex Dt/b2 = D (5 x 3600)/(5.08)2 = 698 D to get Ey Dt/c2 = D (5 x 3600)/(7.62)2 = 1047 D to get Ez An iterative procedure is used to determine the diffusivity, D. 1. Assume a value of D in cm2/s. 2. Compute the three Fourier numbers. 3. From Fig. 3.9, read off values of Ex , Ey , and Ez for the corresponding Fourier numbers. 4. Calculate Eavg from Eq. (1). If it does not equal 0.60, repeat the four steps. This procedure leads to the following results, where the first assumed value of D is obtained by reading the Fourier number for Eavg = 0.61/3 = 0.84 for Fig. 3.9 and using it with the Fourier number for the y dimension to compute D. Assumed D, cm2/s 3.2 x 10-5 3.4 x 10-5 Dt/a2 0.011 0.012 Dt/b2 0.022 0.024 Dt/c2 0.033 0.036 Ex 0.92 0.91 Ey 0.84 0.83 Ez 0.81 0.80 Eavg 0.626 0.604 Therefore, D = 3.4 x 10-5 cm2/s. (b) At 4 wt% average moisture, Eavg = (2-4)/(2-12) = 0.2 The three Fourier numbers, with time in hours, are 2.37 x 10-3 t, 4.74 x 10-3 t, and 7.11x 10-3 t. By iteration, to satisfy Eq. (1), using Fig. 3.9, time = 32 h or additional time = 27 h. Exercise 3.25 Subject: Diffusion of moisture from a 2-inch (5.08 cm) diameter ( a = 2.54 cm radius) spherical ball of clay into air. Given: Initial moisture content of clay = 10 wt%. Equilibrium moisture content of clay-air interface = 3 wt%. Diffusivity of water in clay = D = 5 x 10-6 cm2/s. Assumptions: Diffusivity is uniform throughout clay. Find: Time for average moisture content to drop to 5 wt%. Analysis: By definition in Eq. (3-80), the unaccomplished fractional moisture concentration change after time, t is: Eavg = (cs - cavg)/( cs - co) = (3-5)/(3-10) = 0.286 The lower curve in Fig. 3.9 relates that change to the Fourier number. From the figure, N Fo M 5 10-6 Dt = 0.08 = 2 = t = 7.75 10 -7 t 2 a 2.54 Solving Eq. (1), gives t = 1.03 x 105 s = 28.7 h Exercise 3.26 Subject: Absorption of pure oxygen into a downward flowing laminar water film Given: Film, at 10 atm and 25oC, flows down a wall, with L = 1 m high and W = 6 cm (0.06 m) wide. Flow is laminar, without ripples, at a Reynolds number of 50. Assumptions: Diffusivity of O2 in water = 2.5 x 10-5 cm2/s = 2.5 x 10-9 m2/s. Solubility of O2 in water = a mole fraction of 2.3 x 10-4 = xO 2 . Negligible vaporization of water. Find: Rate of absorption Analysis: The exercise is similar to Example 3.13. By material balance on the O2 in the water film, Eq. (3-103) gives: nO2 = u y W cO2 where, W = 0.06 m and cO2 o L - cO2 o (1) = oxygen concentration in the water at the top of the film = 0. (2) 1/ 3 From Eq. (3-92), the bulk average film velocity is, g 2 uy = 3 3 where from Eq. (3 - 92), film thickness = = 2 g (3) The Reynolds number for film flow is given by (3-93). For a water viscosity, L= 0.89 cP = 0.00089 kg/m-s at 25oC, the film flow rate per unit width of film is, 0.00089 = 50 = 0.01112 kg / m - s 4 4 From Eq. (3), the film thickness for = 1,000 kg/m3 and g = 9.807 m/s2 is, = N Re 3(0.00089)(0.01112) = (1,000) 2 (9.807) 1/ 3 = 1.45 10 -4 m (1,000)(9.807)(1.45 10 -4 ) 2 From Eq. (2), u y = = 0.0772 m / s 3(0.00089) Combining Eqs. (3-103) and (3-113), we obtain an equation for the O2 concentration in the water at the bottom of the film, noting that with pure oxygen gas, all of the mass-transfer resistance resides in the liquid water film: A k cavg W cO2 = cO2 - cO2 - cO2 exp - (4) L i i o uy Exercise 3.26 (continued) Analysis: (continued) The total concentration in the liquid is approximately that of water = c = /Mwater = 1,000/18.02 = 55.5 kmol/m3. The concentration of oxygen in the water at the interface = cO2 = xO2 c = 2.3 x 10-4 (55.5) = 0.0128 kmol/m3. i The average mass-transfer coefficient is given by (3-110), provided that the factor , given by Eq. (3-100), is greater than 0.1. This equation requires the Schmidt number. From Eq. (3-101), N Sc = From Eq. (3-100), = 8/3 8/3 = = 1.033 (> 0.1) N Re N Sc / L (50)(356) 145 10-4 / 1 . DO2 ,H 2 O = 0.00089 = 356 (1,000)(2.5 10-9 ) k cavg From Eq. (3-110), uy 0.0772(1.45 10 -4 ) = 0.241 + 51213 = . 0.241 + 51213(1.033) = 6.2 10-5 m / s . L 1 Area for mass transfer = A = LW = (1)(0.06) = 0.06 m2. From Eq. (4), 0.06 0.06 = 0.0128 - 0.0128 - 0 exp - = 0.01275 kmol / m3 (0.0772)(145 10-4 ) . 6.2 10 -5 cO 2 L Thus, the water is better than 99% saturated with oxygen. From Eq. (1), nO2 = (0.0772) (1.45 10 -4 ) ( 0.06 )( 0.01275 - 0 ) = 8.6 10-9 kmol/s Exercise 3.27 Subject: Absorption of pure carbon dioxide into a downward flowing laminar water film. Given: Results in Example 3.13, including: DAB = 1.96 x 10-5 cm2/s, u y = 0.0486 m/s, = 1.15 x 10-2 cm, and cCO2 ( ) o =0 Find: Height, y, from the top of the film where the water is 50% saturated with CO2. Analysis: At height, y = L, from the top of the film, cCO 2 y / cCO 2 i = 0.5 . If we use Eqs. (3-99) and (3-100), we do not have to compute the average mass-transfer coefficient. Compute the value of from (3-99): 1 - 0.5 = 0.5 = 0.7857 exp ( -5.1213 ) + 0.09726 exp ( -39.661 ) + 0.036093exp ( -106.25 ) Solving this nonlinear equation for with a spreadsheet, = 0.0894 From a rearrangement of Eq. (3-100), 3 (1.15 10-2 ) (4.86)(0.0894) 2 L= 32 u y 2 DAB = 2 (1.96 10 -5 ) = 4.4 cm Exercise 3.28 Subject: Evaporation of water (A) from a film on a flat plate into air (B) at 1 atm (101.3 kPa) and 25oC (298 K) that is flowing past the plate. Given: Free-stream air velocity = uo = 2 m/s. Plate length = L = 2 inches = 0.0508 m. Diffusivity of water vapor in air = DA,B = 0.25 cm2/s = 0.25 x 10-4 m2/s Assumptions: Ideal gas law. Laminar boundary-layer flow of the air. Negligible absorption of air by the water film. Equilibrium at the air-water interface. Find: Evaporation flux at 50% of L = x = 0.0254 m from the leading edge of the plate. Analysis: First check to see if the boundary layer is laminar. From Perry's Handbook, for air at 25oC, = 0.018 cP = 1.8 x 10-5 kg/m-s. From the ideal gas law, = PM/RT = (101.3)(29)/(8.314)(298) = 1.186 kg/m3 From Eq. (3-127), at x = 0.0254 m, xu (0.0254)(2)(1186) . N Re x = o = = 3,350 -5 18 10 . Therefore, the boundary layer is laminar, since the criterion is less than 5 x 105. Since the liquid film is pure water, the only mass-transfer resistance for evaporation of water into air is in the air boundary layer. From Eq. (3-105), the molar mass-transfer flux of water is, nH 2 O = k cx cH 2 O - cH 2 O (1) i o A For a laminar boundary layer, the local mass-transfer coefficient can be obtained from Eq. (3132), which can be rearranged to give: k cx = 0.332 From Eq. (3-101), N Sc = DA,B N Re x x 1/ 2 N Sc 1/ 3 (2) 18 10 -5 . = = 0.607 DA,B (1186) 0.25 10 -4 . From Eq. (2), k cx = 0.332 0.25 10 -4 0.0254 3,350 1/ 2 0.607 1/ 3 = 0.16 m / s At 25oC, vapor pressure of water = 0.46 psia, Therefore, mole fraction of water at the interface = yH 2 O = 0.46/14.7 = 0.0313. cH 2 O = yH 2 O /Mair = 0.0313(1.186)/29 = 0.0128 i i i kmol/m3.From Eq. (1), nH 2O A = (0.016)[0.00128 - 0] = 2.05 10-5 kmol/s-m 2 Exercise 3.29 Subject: Sublimation of a thin, flat plate of naphthalene (A) into air (B) at 100oC (373 K) and 1 atm (101.3 kPa) that is flowing past the plate. Given: Length of plate, L, is 1m. Reynolds number at the trailing edge of the plate = N Re L = 5 105 = limit of laminar boundary-layer flow. Find: (a) Average rate of sublimation. (b) Local rate of sublimation at x = 0.5 m from leading edge of plate. Analysis: This exercise is similar to Example 3.14, which also deals with the sublimation of naphthalene into air at the same temperature and pressure. Therefore, the physical properties used in Example 3.14 apply here. The only mass-transfer resistance is in the gas phase. (a) From Eq. (3-105), the average rate of sublimation per unit area is, nA,B / A = k cavg cA i - cA o (1) From Eqs. (3-137) and (3-138), the average mass-transfer coefficient is given by, k cavg = 0.664 DA,B N Re L L 1/ 2 N Sc 1/ 3 = 0.664 0.94 10 -5 5 105 1 1/ 2 2.41 1/ 3 = 0.0059 m / s From Example (3.14), cA i =4.3 x 10-2 kmkol/m3 and cA o = 0. From Eq. (1), nA,B / A = 0.0059 4.3 10 -2 - 0 = 2.54 10 -4 kmol / s - m2 (b) From Eqs. (3-132) and (3-133), using physical properties in Example (3.14), with a Reynolds number of one-half of 5 x 105, since x = 0.5 m, k cx = 0.332 DA,B N Re x x 1/ 2 N Sc 1/ 3 = 0.332 0.94 10 -5 2.5 105 0.5 1/ 2 2.41 1/ 3 = 0.00417 m / s From Eq. (3-105), the local rate of sublimation per unit area is, nA,B / A = kcx ( cA )i - ( cA )o = 0.00417 ( 4.3 10-2 - 0 ) = 1.79 10 -4 kmol/s-m 2 Exercise 3.30 Subject: Sublimation of a straight, circular tube of naphthalene (A) into air (B) at 100oC (373 K) and 1 atm that flows through it in laminar flow. Given: Tube diameter, D, is 5 cm (0.05 m). Reynolds number = NRe = 1,500. Physical properties in Example 3.14. Assumptions: Fully established laminar-flow velocity profile for air entering the tube. Negligible pressure drop. Equilibrium at naphthalene-air interface. Find: Tube length, x, for average naphthalene mole fraction in exiting air, yA L , to equal 0.005. Analysis: From Example 3.14, yA i = PAs / P = 10 / 760 = 0.0132 By material balance for sublimation of naphthalene into air, the rate of mass transfer over the tube length must be: nA = ux S cA x - cA 0 (1) This rate of mass transfer is given by (3-105), which with a log-mean driving force is, nA = k cavg A cA i - cA ln 0 - cA i - cA 0 x x cA i - cA cA i - cA (2) For the entering air, cA 0 = 0. Therefore, combining Eqs. (1) and (2), and solving for k cavg , k cavg = cA i - cA ux S ln A cA i - cA 0 x (3) From the tube geometry, S = D2/4 = (3.14)(0.05)2/4 = 0.00196 m2 A = Dx = (3.14)(0.05)x = 0.157 x m2 For laminar flow, Eq. (3-150) applies, which contains the Peclet number, Eq. (3-102): N Pe M = N Re N Sc = (1,500)(2.41) = 3,620 From Eq. (3-150), using the definition of the Sherwood number in Eq. (3-148), but for average conditions, Exercise 3.30 (continued) Analysis: (continued) k cavg = 0.0668 N Pe M / x / D DA,B D N Sh avg = A,B 3.66 + D D 1 + 0.04 N / x/D Pe M 2/3 = 0.0668 3620 / x / 0.05 0.94 10-5 3.66 + = 2/3 0.05 1 + 0.04 3620 / x / 0.05 = 0.000688 + 0.00227 x 1 + 1.28 / x 2 / 3 (4) N Re = Dux N (1,500)(0.000215) , therefore, u x = Re = = 68 cm / s = 0.68 m / s 0.0327(29) D (5) 1,000 ln cA i - cA cA i - cA 0 x = ln yA i - yA yA i - yA 0 x = ln 0.0132 - 0.0 = 0.476 0.0132 - 0.005 Combining Eqs. (3) and (4), 0.000688 + Rearranging Eq. (5), 0.00227 0.00404 - =0 x x 1 + 1.28 / x 2 / 3 0.00227 (0.68)(0.00196) 0.00404 = (0.476) = 0.157 x x x 1 + 128 / x 2 / 3 . (5) f x = 0.000688 + Solving the nonlinear Eq. (6), x = 3.73 m (6) Exercise 3.31 Subject: Evaporation of a spherical water (A) drop into still, dry air (B). Given: Initial drop diameter = dp = 1 mm = 0.1 cm. Air temperature = T = 38oC. Drop temperature = 14.4oC. Pressure = P = 1 atm. Assumptions: Ideal gas law. Quasi-steady-state evaporation of water into air. That is, the rate of accumulation of water vapor in the air is negligible compared to the rate of diffusion of the water through the air. Find: (a) A Sherwood number of 2 for mass transfer from the drop surface to the air. (b) The initial mass of the drop. (c) The initial rate of evaporation of the drop in g/s. (d) The time in seconds for the drop diameter to be reduced to 0.2 mm = 0.02 cm. (e) The initial rate of heat transfer to the drop. If insufficient to supply the necessary heat, what will happen? Analysis: (a) With a pure water drop, the only resistance to mass transfer is in the air by ordinary molecular diffusion from the drop surface out to infinity. With the quasi-steady-state assumption, Eq. (3-74) for diffusion through a spherical region becomes, cA DAB 2 cA = 2 r =0 t r r r Therefore, d 2 dcA r =0 dr dr Integrating Eq. (2) twice, C1 + C2 with boundary conditions of: cA = cA s at r = rs and cA = 0 at r = r With these boundary conditions, the two constants of integration are evaluated to give for the region outside the drop, r (3) cA = cA s s r cA = - The rate of mass transfer from the drop surface to the bulk air in terms of a mass-transfer coefficient, Eq. (3-105), can be equated to the diffusion rate from the drop surface, as given by Fick's first law, as in Eq. (3-63), to give the following equation, which is similar to Eq. (3-106), (2) (1) Exercise 3.31 (continued) Analysis: (a) (continued) nA = k c A cA s - cA = - DAB A Differentiating Eq. (3), dcA dr (4) r = rs dcA dr r = rs = cA s rs - 1 r2 =- r = rs cA s rs (5) Substituting Eq. (5) into Eq. (4), k c A cA s - 0 = - DAB A - cA s rs (6) Simplifying and rearranging Eq. (6) into the form of a Sherwood number, like Eq. (3-112), with a characteristic length of the drop diameter, NSh = kc (2rs ) kc d p = =2 DAB DAB d3 p 6 (7) (0.1)3 = 5.23 10-4 g 6 (b) Initial mass of a spherical drop = m = V = = (1.0) (3.14) (c) Initial rate of evaporation can be obtained for Fick's law or from Eq. (3-105) with a masstransfer coefficient. From Eq. (3-105), nA = k c A cA s - cA = k c Ac yA s - 0 (8) The binary gas diffusivity is obtained from Eq. (3-36) using T = 38oC = 311 K and P = 1 atm. 2 M A,B = = 22.2 1 1 + 18 29 From Table 3.1, V A = 131, . V B = 19.7 DAB = 0.00143(311)1.75 = 0.273 cm2 / s (1)(22.2)1/ 2 [(131)1/ 3 + (19.7)1/ 3 ]2 . Exercise 3.31 (continued) Analysis: (c) (continued) From Eq. (7), kc = 2 DAB (2)(0.273) = = 5.46 cm/s dp 0.1 2 Area for mass transfer = area of the drop = A = d p = (3.14)(0.1)2 = 0.0314 cm2 From the ideal gas law, the total gas concentration is, c= P 1 = = 3.92 10-5 mol/cm3 RT (82.06)(311) The mole fraction of water in the air at the surface of the drop is obtained from the vapor pressure of water at 14.4oC = 12.3 torr, with a total pressure of 1 atm = 760 torr, yA s = PAs 12.3 = = 0.0162 P 760 Substituting into Eq. (8), nA = (5.46)(0.0314)(3.92 10 -5 )(0.0162 - 0) = 1.09 10 -7 mol/s = 1.96 x 10-6 g/s (d) At any instant of time, the rate of evaporation by mass transfer = rate of loss of drop. Let N = moles of water drop at any time, t. Then, kc Ac yA s - 0 = - Vdrop = 3 d p ( ) dVdrop dN = - water dt M water dt dVdrop dt = (9) 6 , Therefore, 2 3d p d ( d p ) 6 dt 2 A = d p , From Eq. (7), kc = 2 DAB dp Combining these equations and solving for dt, dt = - water d p 4 M water DABcyA s d dp (10) Integrating Eq. (10) from the initial drop diameter, d p1 , to the later drop diameter, d p2 , Exercise 3.31 (continued) Analysis: (continued) 2 2 d p1 - d p2 water (1.0) = t= 4 M water DABcyA s 2 4(18)(0.273)(3.92 10 -5 )(0.0162) ( 0.1) - ( 0.02 ) 2 2 2 = 380 s (e) By analogy to mass transfer, the rate of heat transfer from the air to the drop is, 2k k dp o -6 From Perry's Handbook, thermal conductivity of air at 38 C = k = 58 x 10 cal/s-cm2-(oC/cm) h = 2(58 x 10-6)/0.1 = 0.00116 cal/s-cm2-oC. Ts - T = 38 - 14.4 = 23.6oC. A = 0.0314 cm2 Therefore, the initial rate of heat transfer = Q = 0.00116(0.0414)(23.6) = 0.00086 cal/s Q = hA Ts - T , where N Nu = 2 = or h = The rate of heat transfer required to maintain the drop at 14.4oC, based on the initial drop diameter is, hd p Qneeded = ( nevap M water ) H vap + ( CP ) water vapor (Ts - T ) = (1.96 10-6 ) [589 + 0.44(38 - 14.4) ] = 0.001176 cal/s since, from Perry's Handbook, the heat of vaporization of water at 14.4oC is 589 cal/g and the specific heat of steam is 0.44 cal/g-oC. Because the heat needed is greater than the rate of heat transfer, the temperature of the drop will decrease with time initially. Exercise 3.32 Subject: Dissolution of benzoic acid when water flows in turbulent motion.through a straight, circular tube of it Given: Tube has an initial inside diameter, D, of 2 inches (5.08 cm), with a length, L, of 10 ft (305 cm). Water at 25oC flows through the tube, with a velocity, ux , of 5 ft/s (152 cm/s). Physical properties in Example 3.15. Assumptions: Fully developed turbulent flow. Find: Average concentration of benzoic acid in the water exiting the tube. Analysis: The rate of increase in benzoic acid (A) in the water (B), by the continuity equation, equals the rate of mass transfer. The only resistance to mass transfer is in the water. Using Eq. (3-105), with a log mean driving force, nA = ux S cA out - cA in = k c A cA wall - cA = kc A cA wall - cA in - cA wall - cA out ln Because cA in = 0, Eq. (1) simplifies to, cA wall kc A = ln or solving, ux S cA wall - cA out cA out = cA wall 1 - exp - kc A ux S (2) cA wall - cA in cA wall - cA out (1) LM Area for mass transfer = A = DL = 3.14(5.08)(305) = 4,870 cm2 Cross-sectional area for flow = S = D2/4 = 3.14(5.08)2/4 = 20.25 cm2 Use the Chilton-Colburn j-factor method to obtain kc . Combining Eqs. (3-165) and (3-166), 0.023G N Re kc = N Sc N Re = -0 . 2 2/3 0.023ux N Re = N Sc -0.2 2/3 = 0.023ux N Re N Sc -0 . 2 2/3 (3) DG Dux (5.08)(152)(10) . = = = 86,800 (confirms turbulent flow) (4) 0.0089 N Sc = 0.0089 = = 970 DAB (10)(9.18 10-6 ) . (5) Exercise 3.32 (continued) Analysis: (continued) Substitution of the results of Eqs. (4) and (5) into (3) gives, k c = 0.023(152) 86,800 970 -0 . 2 2/3 = 0.00367 cm/s From Example 3.15, equilibrium exists at the inside wall of the tube such that the concentration of benzoic acid in the water is at its equilibrium value of cA wall = 0.0034 g/cm3 From Eq. (2), cAout =0.0034 1 - exp - (0.00367)(4,870) (152)(20.25) = 0.000020 g/cm3 Therefore, the concentration of benzoic acid in the exiting water is way below the solubility value. Exercise 3.33 Subject: Sublimation of a long, circular cylinder of naphthalene (A) to air (B) flowing normal to it. Given: Long, circular cylinder of 1-inch (0.0254 m) diameter, D. Air at 100oC (373 K) and 1 atm flows at a Reynolds number, NRe , of 50,000. Physical properties given in Example 3.14. Assumptions: Negligible bulk flow effect. Find: Average sublimation flux in kmol/s-m2. Analysis: Only mass-transfer resistance is in the air. From Eq. (3-105), the flux is, nA / A = N A = k c cA s - cA (1) Using the Chilton-Colburn analogy, Eq. (3-169) applies, which combined with Eq. (3-165) gives, 0.0266G N Re kc = N Sc -0.195 2/3 (2) From the Reynolds number, mass velocity, G, is, using = 0.0215 cP (2.15 x 10-5 kg/s-m2), G= N Re (50,000)(2.15 10-5 ) = = 42.3 kg / s - m2 D 0.0254 From Example 3.14, gas density = 0.0327(29) = 0.948 kg/m3 and NSc = 2.41. 0.0266G N Re From Eq. (2), k c = N Sc Using properties in Example 3.14, cA s -0.195 2/3 0.0266(42.3) 50,000 = 2/3 0.948 2.41 -0.195 = 0.080 m/s PAs 10 = yA s c = c= (0.0327) = 4.3 10-4 kmol / m3 and cA = 0 P 760 From Eq. (1), N A = kc cA s - cA = (0.080)(4.3 10-4 - 0) = 3.44 10 -5 kmol/s-m2 ( ) Exercise 3.34 Subject: Sublimation of a sphere of naphthalene (A) into air (B) flowing past it. Given: Sphere of 1-inch (0.0254 m) diameter, D. Air at 100oC (373 K) and 1 atm flows at a Reynolds number, NRe , of 50,000. Physical properties given in Example 3.14. Assumptions: Negligible bulk flow effect. Find: Initial average sublimation flux in kmol/s-m2 and comparison to that in a bed packed with the same spheres at a void fraction of 0.5. Analysis: Only mass-transfer resistance is in the air. From Eq. (3-105), the flux is, nA / A = N A = k c cA s - cA (1) Using the Chilton-Colburn analogy, Eq. (3-170) applies, which combined with Eq. (3-165) gives, 0.37G N Re kc = N Sc -0 . 4 2/3 (2) From the Reynolds number, mass velocity, G, is, using = 0.0215 cP (2.15 x 10-5 kg/s-m2), G= N Re (50,000)(2.15 10-5 ) = = 42.3 kg / s - m2 D 0.0254 From Example 3.14, gas density = 0.0327(29) = 0.948 kg/m3 and NSc = 2.41. 0.37(42.3) 50,000 From Eq. (2), k c = 2/3 0.948 2.41 -0 . 4 = 0121 m/s . Using properties in Example 3.14, PAs 10 cA s = yA s c = c= (0.0327) = 4.3 10-4 kmol / m3 and cA = 0 P 760 From Eq. (1), N A = kc cA s - cA = (0.121)(4.3 10-4 - 0) = 5.20 10-5 kmol/s-m2 From Eq. (3-171), packed bed correlation can't used with NRe = 50,000. ( ) Exercise 3.35 Subject: Experimental data for stripping of CO2 (A) from water by air (B) in a wetted-wall tube. Given: Location in the tube where P = 10 atm, T =25oC (298 K), CO2 flux from water into air = NA = 1.62 lbmol/h-ft2, partial pressure of CO2 at gas-liquid interface = pA i = 8.2 atm, and partial pressure of CO2 in bulk air = pA b = 0.1 atm. Diffusivity of CO2 in air = DAB = 1.6 x 10-2 cm2/s. Assumptions: Neglect stripping of water, but account for the bulk flow effect. Mass-transfer resistance of the water film does not need to be considered. Ideal gas law. Find: Using the film theory, find the mass-transfer coefficient, kc , for the gas phase and the film thickness, . Analysis: Solve the problem in cgs units. From Eq. (3-105) and a rearrangement of Eq. (3-189), kc = N A 1 - yA LM c yA i - yA b = DAB (1) NA = 1.62 lbmol/h-ft2 = 0.00022 mol/s-cm2 From the ideal gas law, c =P/RT = (10)/(82.06)(298) = 0.000409 mol/cm3 yA i = pA i P = 8.2 = 0.82 10 yA b = pA b P = 0.1 = 0.01 10 (1 - yA ) LM = (1 - yA b ) - (1 - yA i ) (1 - 0.01) - (1 - 0.82) = = 0.475 (1 - yA b ) (1 - 0.01) ln ln (1 - 0.82) (1 - yA i ) From Eq. (1), kc = (0.00022)(0.475) = 0.315 cm/s (0.000409) ( 0.82 - 0.01) From Eq. (1), = DAB / kc = 1.6 x 10-2/0.315 = 0.051 cm Exercise 3.36 Subject: Experimental data for absorption of CO2 (A) from air into water (B) using a packed column of Pall rings. Given: Location in a packed column where pA i = 150 psia, cA b = 0 , and NA = 0.017 lbmol/h-ft2 (2.31 x 10-6 mol/s-cm2). Diffusivity of CO2 in water = DAB = 2.0 x 10-5 cm2/s. Henry's law for CO2 between air and water is p, psia = 9,000 x. Assumptions: Negligible bulk flow effect. Only the mass-transfer resistance in the water need be considered. Neglect stripping of water. Find: (a) Liquid-phase mass-transfer coefficient, kc , and film thickness, , for the film theory. (b) Contact time for penetration theory. (c) Average eddy residence time and probability distribution for surface renewal theory. Analysis: Solve the problem in cgs units. (a) From Eq. (3-105) and a rearrangement of Eq. (3-189), kc = NA D = AB c xA i - xA b (1) Total liquid-phase concentration, since almost pure water, = c = /MB = 1.0/18 = 0.0556 mol/cm3 From given Henry's law, xA i = pA i /9,000 = 150/9,000 = 0.0167 and xA b = 0.0. From Eq. (1), kc = 2.31 10-6 = 0.0025 cm/s (0.0556) ( 0.0167 - 0 ) From Eq. (1), = DAB /kc = 2 x 10-5/0.0025 = 0.0080 cm (b) From a rearrangement of Eq. (3-194), tc = 4 DAB 4(2 10 -5 ) = = 4.08 s (3.14)(0.0025) 2 kc2 1 DAB 2 10-5 (c) From Eq. (3-201), t = = 2 = = 3.2 s s kc (0.0025) 2 From Eq. (3-196), the probability distribution is: {t} = s exp(- st ) = 0.313 exp(-0.313t ) Exercise 3.36 (continued) Exercise 3.37 Subject: Determination of diffusivity of H2S (A) in water (B) from experimental data for absorption into a laminar jet, using the penetration theory. Given: Laminar jet of D = 1 cm and L = 7 cm. Temperature, T = 20oC (293 K). Solubility of H2S in water = cA i = 100 mol/m3 (100 x 10-6 mol/cm3). Absorption rates, nA, as a function of jet volumetric flow rate, Q. Assumptions: Jet is contacted by pure H2S. Therefore, all mass-transfer resistance is in the liquid. Negligible bulk flow effect. Concentration of H2S in bulk water cA b = 0. Find: Diffusivity by penetration theory. Analysis: Use cgs units. From a rearrangement of Eq. (3-192), DAB = nA 2 2A cAi - cAi ( ) tc (1) Surface area of jet = A = DL = (3.14)(1)(7) = 22.0 cm2 Assume the contact time is the time of exposure of the jet = jet volume/jet flow rate = V/Q Jet volume = V = D2L/4 = (3.14)(1)2(7)/4 = 5.5 cm3. Therefore, tc = 5.5/Q. Therefore, Eq. (1) becomes, DAB Jet rate, cm3/s 0.143 0.568 1.278 2.372 3.571 5.142 nA = 2(22) 100 10 -6 - 0 2 2 (314)(5.5) . 4 nA = 89.2 10 Q Q Absorption rate, mol/s x 106 1.5 3.0 4.25 6.15 7.20 8.75 Diffusivity, DAB , cm2/s 1.40 x 10-5 1.41 x 10-5 1.26 x 10-5 1.42 x 10-5 1.29 x 10-5 1.33 x 10-5 Average DAB = 1.35 x 10-5 cm2/s. Experimental value at 25oC is 1.61 x 10-5 cm2/s. Note that a material balance shows that the bulk concentration of H2S in the water only reaches a high of 10.5 x 10-6 mol/cm3. Thus, the diffusivities would only be slightly higher than above. Exercise 3.38 Subject: Vaporization of water (A) into air (B) in a wetted-wall column. Given: Column inside diameter = D = 1.46 cm and length = L = 82.7 cm. Air volumetric flow rate = Q = 720 cm3/s at 24oC and 1 atm (760 torr). Water enters at 25.15oC and exits at 25.35oC. Partial pressures of water vapor in the air are 6.27 torr entering and 20.1 torr leaving. Diffusivity of water vapor in air = DAB = 0.22 cm2/s at 0oC and 1 atm. Assumptions: Negligible bulk flow effect. Ideal gas law. Negligible pressure drop in column. Find: (a) Rate of mass transfer of water into air. (b) Overall mass transfer coefficient, KG. Analysis: Use cgs units. No mass transfer resistance in the liquid phase because pure water. (a) The rate of mass transfer can be obtained from a material balance on the water content of the air using the entering and exiting partial pressures of the water vapor. For the entering air: Mole fraction of water vapor = yA = pA / P = 6.27/760 = 0.00825 Molecular weight = M = 0.00825(18) + (1-0.825)(29) = 28.9 Density = =PM/RT = (1)(28.9)/(82.06)(298) = 0.00118 g/cm3 Mole flow rate of gas = n = Q/M =(720)(0.00118)/28.9 = 0.02941 mol/s Mole flow rate of water in entering air = n yA = 0.02941(0.00825) = 2.426 x 10-4 mol/s For the exiting air: Mole fraction of water vapor = yA = pA / P =20.1/760 = 0.02645 1 - 0.00825 = 0.02996 mol/s By ratio with entering gas, mole flow rate of gas = m = 0.02941 1 - 0.02645 Mole flow rate of water in exiting air = m yA = 0.02996(0.02645) = 7.924 x 10-4 mol/s Increase in water content of the gas = rate of mass transfer of water = nA = 7.924 x 10-4 - 2.426 x 10-4 = 5.5 x 10-4 mol/s (b) From Eqs. (3-211) and (3-218), noting that there is no resistance in the liquid phase, KG = k p = ( NA pAi - pAb ) = LM nA AP yAi - yAb ( ) (1) LM Exercise 3.38 (continued) Analysis: (b) (continued) A = DL = 3.14(1.46)(82.7) = 379 cm2 At 25oC, the interface mole fraction of water vapor = yA i = yA i - yA in - yA i - yA out ln From Eq. (1), yA i - yA in yA i - yA out PAs 0.46 psi = = 0.0313 P 14.7 psi y A i - yA b LM = = 0.0313 - 0.00825 - 0.0313 - 0.02645 = 0.0117 0.0313 - 0.00825 ln 0.0313 - 0.02645 KG = 5.5 10-4 = 1.24 10-4 mol/s-cm2-atm (379)(1)(0.0117) Note that the correction for bulk flow amounts to only about 2.5% because the average mole fraction of air between the interface and the bulk is 0.975. Exercise 3.39 Subject: Absorption of ammonia (A) from air (B) into 2 N H2SO4 in a wetted-wall column, with countercurrent flow. Given: Column inside diameter = D = 0.575 inch (0.0479 ft) and length = L = 32.5 in. (2.71 ft) Air flow rate = n = 0.260 lbmol/h at 1 atm. Partial pressures of ammonia in the air are 0.0807 atm entering and 0.0205 leaving. Air enters at 77oF and exits at 84oF. Aqueous acid enters at 76oF and exits at 81oF. Change in acid strength is negligible. Assumptions: Reaction of the basic ammonia with the sulfuric acid is instantaneous such that the partial pressure, pA i , of ammonia at the gas-liquid interface is zero. Therefore, the only mass transfer resistance is in the gas phase. Air flow rate refers to the total entering gas. Find: Mass-transfer coefficient, kp , for the gas phase. Negligible bulk flow effect. Analysis: From a rearrangement of Eq. (3-207), kp = nA A pAb - pAi ( ) = LM nA A pA b - 0 ( ) = LM nA A pAb ( ) (1) LM The rate of mass transfer is obtained by material balance for the decrease in partial pressure of the ammonia in the air. First compute the flow rate of just the air in the gas. 10 - 0.0807 . = 0.239 lbmol/h P 1.0 pA in 0.0807 Ammonia in entering gas = 0.239 = 0.021 lbmol/h nB = 10 - 0.0807 . pBin nB = n = 0.260 Ammonia in exiting gas = P - pA in pA out pBout nB = 0.0205 0.239 = 0.005 lbmol/h 10 - 0.0205 . Rate of mass transfer of ammonia = nA = 0.021 - 0.005 = 0.016 lbmol/h Area for mass transfer = A = DL = 3.14(0.0479)(2.71) = 0.408 ft2 pA b = pA in - pA out 0.0807 - 0.0205 = = 0.0439 atm pA in 0.0807 ln ln 0.0205 pA out LM From Eq. (1), k p = nA / A pAb ( ) LM = 0.016 /(0.408)(0.0439) = 0.893 lbmol/h-ft2-atm The bulk flow effect amounts to a correction of about 5%. Exercise 3.40 Subject: Overall mass-transfer coefficient for a packed cooling tower, where water (A) is evaporated into air (B). Given: Cross-sectional area for air flow = S = 0.5 ft2. Height over which data were taken = H = 7 ft. Air rate = n = 0.401 lbmol/h. Water rate = 20 lbmol/h. At the bottom, P = 14.1 psia, T = 120oF, PAs = 169 psia, and yA b = 0.001609. At the top, PAs = 1995 psia, yA b = 0.0882, T = 126oF, . . and P = 14.3 psia. Assumptions: All mass-transfer resistance resides in the gas phase because pure water is used. Assume air rate pertains to moisture-free air. Bulk-flow effect may be appreciable. Find: Overall mass-transfer coefficient, Ky Analysis: From a rearrangement of Eq. (3-234) and including the bulk-flow effect, nA 10 - yA LM . Ky = (1) A yA i - yA b LM However, no data are given on the packing to enable the determination of the area, A. Therefore, let a = area, A, for mass transfer per unit volume, V, of tower. Then Eq. (1) becomes, Kya = nA 10 - yA . V yA i - yA b LM LM (2) Therefore, Kya will be determined instead of Ky . The rate of mass transfer of water into the air is obtained by material balance, with the water mole fraction data and the air flow rate (water-free basis), where, nA = n yA top 1.0 - yA top - yA bottom 1.0 - yA bottom = 0.401 0.0882 0.001609 - = 0.0381 lbmol/h 1.0 - 0.0882 1.0 - 0.001609 The volume over which the data were taken = V = SH = 0.5(0.7) = 0.35 ft2 Interface mole fractions of water vapor are, yA itop = PAstop / Ptop = 1995 / 14.3 = 0.1395 . yA i bottom = PAsbottom / Pbottom = 1.69 / 14.1 = 01199 . Exercise 3.40 (continued) Analysis: (continued) y A i - yA b LM = yA itop - yA btop - yA ibottom - yAbbottom ln yA itop - yA btop yA ibottom - yA bbottom = 01395 - 0.0882 - 01199 - 0.001609 . . = 0.0802 01395 - 0.0882 . ln 01199 - 0.001609 . For the bulk-flow effect, (1.0 - yA)LM at the top = 1 - yA i ln top - 1 - yA b top top 1 - yA i = 1 - yA b . 1 - 01395 - 1 - 0.0882 = 0.886 1 - 0.1395 ln 1 - 0.0882 top (1.0 - yA)LM at the bottom = 1 - yA i bottom - 1 - yA b bottom bottom ln 1 - yA i = 1 - yA b 1 - 01199 - 1 - 0.001609 . = 0.938 1 - 0.1199 ln 1 - 0.001609 bottom The average value of (1.0 - yA)LM is 0.912. From Eq. (2), K ya = 0.0381(0.912) = 1.24 lbmol/h-ft3-unit mole fraction (0.35)(0.0802) For the same packing and gas and liquid velocities, this value can be used to design a large unit. Exercise 4.1 Subject: Degrees of freedom analysis for a three-phase equilibrium stage. Given: Equilibrium stage of Figure 4.35, with two feeds (one vapor, one liquid), vapor entering from stage below, liquid entering from stage above, three exiting streams (one vapor, two liquid), and heat transfer. Assumptions: Equilibrium stage Find: (a) (b) (c) (d) Analysis: List and count of variables. List and count of equations. Number of degrees of freedom. List of reasonable set of design variables. (a) With 4 streams in and 3 streams out, and heat transfer, Number of variables = NV = 7(C+3) + 1 = 7C + 22 The variables are 7 total flow rates, 7 temperatures, 7 pressures, 1 heat transfer rate and C mole fractions for each of the 7 streams. (b) The equations are: C Component material balances 1 Energy balance 2 Pressure identity equations for the 3 exiting streams 2 Temperature identity equations for the 3 exiting streams 7 Mole fraction sums (one for each stream) y yi 2C Phase equilibrium equations: KiI = iI and KiII = II xi xi Total number of equations = NE = 3C + 12 (c) Degrees of freedom = ND= NV - NE = (7C + 22) - (3C + 12) = 4C +10 (d) A possible set of specifications is: For each entering stream: Total flow rate, temperature, pressure, and C-1 mole fractions, which totals 4(C + 2) = 4C + 8 For the remaining 2 variables, choose any combination of Q, temperature of one of the three exiting streams, and/or pressure of one of the three exiting streams. Exercise 4.2 Subject: Determination of uniqueness of three different operations. Given: (a) An adiabatic equilibrium stage with known vapor and liquid feed streams, and known stage temperature and pressure. (b) Same as (a), except that stage is not adiabatic. (c) Partial condenser using cooling water, with known vapor feed (except for flow rate), outlet pressure, and inlet cooling water flow rate. Assumptions: Exiting streams in equilibrium. Find: (a) Whether composition and amounts of exiting vapor and liquid can be computed. (b) Same as part (a). (c) Whether cooling water rate can be computed. Analysis: (a) With two steams in and two out, number of variables = NV = 4(C + 3) = 4C + 12 Equations are: C Component material balances 1 Energy balance 1 Pressure identity for 2 exiting streams 1 Temperature identity for 2 exiting streams 4 Mole fraction sums for 4 streams C Phase equilibrium equations Therefore, number of equations = NE =2C + 7 Degrees of freedom = ND= NV - NE = (4C + 12) - (2C + 7) = 2C +5 Given specifications are: 2C + 4 variables for the two feed streams. Only one specification left. Therefore can not specify both T and P for exiting streams. (b) If stage in part (a) is not adiabatic, add Q as a variable to give NV = 4C + 13. The number of equations stays the same, i. e. NE =2C + 7. Thus, have one additional degree of freedom, giving, ND= NV - NE = 2C +6. Can now specify both T and P for exiting streams. (c) First, consider just the partial condensation of the vapor into two exiting streams by heat transfer, Q, without considering the cooling water. For three streams, NV =3(C + 3) + 1 (for Q) = 3C +10. Equations are: C Component material balances 1 Energy balance 1 Pressure identity for 2 exiting streams 1 Temperature identity for 2 exiting streams 3 Mole fraction sums for 3 streams C Phase equilibrium equations Therefore, NE = 2C + 6 and ND = NV - NE = C + 4 Exercise 4.2 (continued) Analysis: (c) (continued) Specified for the feed vapor are only C + 1 variables, because the feed rate is not specified. Also the outlet pressure of the condenser is specified. This give C + 2. We are short two variables from being able to compute Q. An energy balance on the cooling water gives: Q = mCP Tout - Tin But only Tin is given. Thus, with Q unknown and Tout not given, we are three variable short of being able to compute the water rate, m. Thus, the problem can not be solved uniquely. We could solve it if the vapor feed rate, the outlet temperature of the partial condensate, and the cooling water outlet temperature were specified. Exercise 4.3 Subject: Degrees of freedom analysis for an adiabatic, two-phase flash. Given: Continuous, adiabatic flash of one feed into vapor and liquid products Assumptions: Exiting streams are in equilibrium Find: (a) (b) (c) (d) (e) Analysis: Number of variables. All equations relating variables. Number of equations. Number of degrees of freedom Preferred specifications (a) Variables are those appearing in Figure 4.36 N V = 3C + 9 (b) C Component material balances 1 Energy balance 1 Pressure identity for two exiting streams 1 Temperature identity for two exiting streams 3 Mole fraction sums for three streams C Vapor-liquid equilibrium equations (c) NE = 2C + 6 (d) ND = NV - NE = (3C + 9) - (2C + 6) = C + 3 (e) Specify the feed completely ( feed rate, temperature, pressure and C - 1 mole fractions) plus exiting pressure. Exercise 4.4 Subject: Degrees of freedom analysis for a non-adiabatic three-phase flash. Given: Continuous, non-adiabatic flash of a liquid feed to produce a vapor and two liquid phases as shown in Figure 4.33. Assumptions: The three exiting phases are in equilibrium. Find: Number of degrees of freedom. Analysis: The variables are the heat transfer rate and four each of stream flow rates, temperatures, pressures, and C mole fractions. Thus, NV = 4C + 13. The equations are: C Component material balances 1 Energy balance 2 Pressure identities for three exiting streams 2 Temperature identities for three exiting streams 4 Mole fraction sums for four streams 2C Phase equilibrium relations: KiI = yi / xiI and KiII = yi / xiII Therefore, NE = 3C + 9 Number of degrees of freedom = ND = NV - ND = (4C + 13) - (3C + 9) = C + 4 Exercise 4.5 Subject: Application of Gibbs phase rule to seven-phase system of Figure 4.31. Given: One gas and six liquid phases in equilibrium. Assumptions: Gas phase includes N2 , O2 , and argon. Find: Number of degrees of freedom by Gibbs phase rule. Possible set of specifications to fix system. Analysis: From Eq. (4-1), Number of degrees of freedom = C - number phases + 2 Number of components = 9 (N2 , O2 , argon, n-hexane, aniline, water, phosphorus, gallium, and mercury. Number of phases = 7 Number of degrees of freedom = 9 - 7 + 2 = 4 Specify T, P, and mole fractions of argon and oxygen in the air. Exercise 4.6 Subject: Partial vaporization of a mixture of benzene and ethyl alcohol. Given: Vapor-liquid equilibrium data (T-x-y) for benzene-ethyl alcohol at 1 atm and vapor pressure data for benzene and ethyl alcohol. Initial mixture contains 25 mol% benzene. Assumptions: Phase equilibrium Find: (a) Bubble point. (b) Composition of vapor at bubble point. (c) Composition of liquid at 25 mol% vaporization. (d) Composition of liquid at 90 mol% vaporization. (e) Same as part (c), except vapor is removed and then additional 35 mol% is vaporized. (f) Plot of temperature vs. mol% vaporized for parts (c) and (e). (g) Repeat of parts (a) to (f) assuming ideal solutions Analysis: See plot of T-x-y data on next page, as drawn with a spreadsheet. Curved (instead of straight) lines connecting the points would be a good improvement. (a) For a benzene mole fraction of 0.25, a vertical line from M intersects the liquid line at N at 69.4oC, which is the bubble point. (b) The benzene mole fraction in the vapor at 69.4oC, obtained from the left-most vapor line at P, is 0.56. (c) To find the benzene mole fraction in the liquid at 25 mol% vaporization, extend a dashed, vertical line upward from the bubble point at N, as shown in the figure on the next page, until point B is reached. At this point, using the inverse lever-arm rule, the ratio of the AB line length to the BC line length is 25/75. The benzene mole fraction in the equilibrium liquid at point A is 0.175 at a temperature of 71.2oC. (d) To find the benzene mole fraction in the liquid at 90 mol% vaporization, extend a dashed, vertical line upward from the bubble point, as shown in the figure on the next page, until point Eis reached. At this point, using the inverse lever-arm rule, the ratio of the DE line length to the EF line length is 90/10. The benzene mole fraction in the equilibrium liquid at point D is 0.045 at a temperature of 75.1oC. (e) To find the benzene mole fraction in the liquid when the liquid from part (c) is removed from the vapor and further vaporized, proceed as follows. If we start with 100 moles, then the liquid at 25 mol% vaporized is 75 moles with a benzene mole fraction of 0.17. If an additional 35 mol% is vaporized, we will have 35 moles of vapor in equilibrium with 40 moles of liquid. Therefore, extend a dashed, vertical line upward from point A, as shown in the figure on the next page, until point H is reached. At this point, using the inverse lever-arm rule, the ratio of the GH line length to the HI line length is 35/40. The benzene mole fraction in the liquid at G is 0.05 at a temperature of 74.9oC. Exercise 4.6 (continued) Analysis: (f) Extending parts (c) and (e) to other percent vaporizations, the following data are obtained, where for the extension of part (e), the first 25 mol% of vapor is removed in each case. Mol% vaporization 0 25 50 75 100 Part (c) temperature, oC 69.4 71.2 72.9 74.3 75.3 Part (e) temperature, oC 69.4 71.2 73.8 75.5 76.3 (g) Thus, To calculate T-x-y curves from vapor pressure data, using Raoult's and Dalton's laws, Eq. (2-44 ) applies, as well as the sum of the mole fractions in the phases in equilibrium. letting A = benzene and B = ethyl alcohol, KA = s yA PA T = xA P , , KB = s yB PB T = xB P (1, 2) (3, 4) yA + y B = 1 xA + x B = 1 Exercise 4.6 (continued) Analysis: (g) (continued) Equations (1) to (4) can be reduced to the following equations for the mole fractions of benzene in terms of the K-values: xA = 1 - KB KA - KB , y A = KA x A (5, 6) If the given vapor pressure data are fitted to Antoine equations, we obtain: PAs = exp 15.5645 - PBs = exp 21.0192 - 2602.34 T + 211.271 (7) 505106 . (8) T + 272.702 Where vapor pressure is in torr and temperature is in oC. Solving, Eqs. (1) to (8), T, oC Ps of A, torr Ps of B, torr KA KB xA yA 78.41 78.50 78.60 78.70 78.80 78.90 79.00 79.10 79.20 79.30 79.40 79.50 79.60 79.70 79.80 79.90 80.00 80.10 80.11 721.3 723.3 725.5 727.8 730.0 732.3 734.6 736.8 739.1 741.4 743.7 746.0 748.3 750.6 752.9 755.2 757.5 759.9 760.0 760.0 762.8 765.9 769.1 772.2 775.4 778.6 781.7 784.9 788.1 791.4 794.6 797.8 801.1 804.3 807.6 810.9 814.2 814.4 0.9491 0.9517 0.9547 0.9576 0.9606 0.9636 0.9665 0.9695 0.9725 0.9755 0.9785 0.9816 0.9846 0.9876 0.9907 0.9937 0.9968 0.9998 1.0000 1.0000 1.0037 1.0078 1.0119 1.0161 1.0202 1.0244 1.0286 1.0328 1.0370 1.0413 1.0455 1.0498 1.0541 1.0584 1.0627 1.0670 1.0713 1.0716 0.000 0.071 0.147 0.220 0.290 0.357 0.422 0.484 0.544 0.602 0.658 0.712 0.764 0.814 0.862 0.909 0.954 0.997 1.000 0.000 0.067 0.140 0.210 0.278 0.344 0.408 0.469 0.529 0.587 0.644 0.699 0.752 0.804 0.854 0.903 0.951 0.997 1.000 The T-x-y plot is shown on the next page. It is drastically different from the one based on experimental data. Thus, Raoult's law K-values for the benzene-ethyl benzene system are greatly in error. Consequently, no calculations are made for parts (a) to (f). They would predict only very small differences between yA and xA over a very small temperature range. Exercise 4.6 (continued) Raoult's law prediction of T-x-y curves for Benzene-Ethyl alcohol at 1 atm Exercise 4.6 (continued) Exercise 4.7 Subject: Steam (B) distillation of stearic acid (A). Given: T = 200oC. Vapor pressure of pure stearic acid at 200oC = 0.40 kPa. Assumptions: Partial pressure of stearic acid in vapor = 70% of the pure vapor pressure. Find: Kilograms of acid distilled per kilogram of steam added as a function of total pressure from 3.3 kPa to 101.3 kPa. Analysis: pA = 0.7(0.4) = 0.28 kPa pB = P - pA = P - 0.28 yi = pi / P M A = 284.5 , M B = 18.02 (1) (2) Using Eqs. (1) and (2), kg A yA M A pA M A 0.28(284.5) 4.42 = = = = kg B yB M B pB M B (P-0.28)(18.02) P - 0.28 (3) Solving Eq. (3) for values of P from 3.3 to 101.3 kPa gives the following results: P, kPa 101.3 75 50 25 15 10 5 3.3 kg A/kg B 0.0438 0.0592 0.0890 0.1790 0.3006 0.4553 0.9376 1.4650 Exercise 4.7 (continued) Exercise 4.8 Subject: Vapor-liquid equilibrium for benzene (A) - toluene (B) system at 1 atm Given: Average relative volatility = 2.5. Vapor pressure data. Assumptions: Raoult's and Dalton's laws. Find: x-y diagram for A,B = 2.5. x-y diagram for Raoult's law using vapor pressure data. (a) Temperature for 25 mol% vaporization of a 70 mol% A/30 mol% B mixture. Composition of condensed vapor and liquid residue. (b) Plot of Raoult's law K-values as a function of temperature. Analysis: For a constant relative volatility, Eq. (4-8) applies. For A,B = 2.5, yA = A,B xA 1 + xA A,B - 1 = 2.5xA 1 + 15xA . Solving this equation for values of xA = 0 to 1.0 gives the following: xA 0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.0 yA 0.0000 0.2174 0.3846 0.5172 0.6250 0.7143 0.7895 0.8536 0.9091 0.9574 1.0000 Exercise 4.8 (continued) Analysis: (continued) To calculate y-x and T-x-y curves from vapor pressure data, using Raoult's and Dalton's laws, Eq. (2-44 ) applies, as well as the sum of the mole fractions in the phases in equilibrium. Thus, KA = s yA PA T = xA P , , KB = s yB PB T = xB P (1, 2) (3, 4) yA + y B = 1 xA + x B = 1 Exercise 4.8 (continued) Analysis: (continued) Equations (1) to (4) can be reduced to the following equations for the mole fractions of benzene (A) in terms of the K-values: xA = 1 - KB KA - KB , y A = KA x A (5, 6) If the given vapor pressure data in Exercise 4.6 for benzene, and this exercise for toluene are fitted to Antoine equations, we obtain: PAs = exp 15.5645 - PBs = exp 17.2741 - 2602.34 T + 211.271 (7) 3896.3 (8) T + 255.67 Where vapor pressure is in torr and temperature is in oC. Solving, Eqs. (1) to (8), T, oC Ps of A, torr Ps of B, torr KA KB xA yA 80.1 82.5 85.0 87.5 90.0 92.5 95.0 97.5 100.0 102.5 105.0 107.5 110.0 110.5 759.9 817.4 880.8 948.0 1019.1 1094.1 1173.4 1256.9 1345.0 1437.6 1535.0 1637.3 1744.7 1766.8 290.0 314.9 342.7 372.5 404.4 438.5 474.9 513.7 555.2 599.3 646.2 696.1 749.1 760.1 0.9998 1.0755 1.1590 1.2474 1.3409 1.4396 1.5439 1.6539 1.7697 1.8916 2.0198 2.1544 2.2957 2.3248 0.3816 0.4144 0.4510 0.4901 0.5321 0.5769 0.6249 0.6760 0.7305 0.7885 0.8503 0.9159 0.9856 1.0001 1.000 0.886 0.775 0.673 0.579 0.490 0.408 0.331 0.259 0.192 0.128 0.068 0.011 0.000 1.000 0.953 0.899 0.840 0.776 0.706 0.630 0.548 0.459 0.363 0.259 0.146 0.025 0.000 Plots of y-x and T-x-y based on the above table from Raoult's law calculations are shown on the next page. When the y-x plot is compared to the previous y-x plot based on a constant relative volatility, it is seen that, for a given value of x for benzene, the values of y for benzene are in fairly close agreement. From the above table, the Raoult's law A,B = PAs / PBs ranges from 2.62 at 80.1oC to 2.32 at 110.5oC. Analysis: (continued) Exercise 4.8 (continued) Exercise 4.8 (continued) Analysis: (continued) (a) To find the temperature at 25 mol% vaporization, starting with a liquid mixture of 70 mol% benzene and 30 mol% toluene, extend a dashed, vertical line upward from point M on the T-y-x diagram on the previous page until point B is reached. At this point, using the inverse lever-arm rule, the ratio of the AB line length to the BC line length is 25/75. The temperature is 88oC. The benzene mole fraction of the equilibrium vapor when condensed is the same as the equilibrium vapor at point C or 0.88. The benzene mole fraction in the residue liquid is the same as the equilibrium liquid at point A or 0.65. (b) The Raoult's law K-values are included in the above table, and are plotted below. Exercise 4.9 Subject: Vapor-liquid equilibrium for n-heptane (A) - toluene (B) system at 1 atm Given: Vapor pressure data for n-heptane and toluene, and experimental T-y-x data. Assumptions: Raoult's and Dalton's laws Find: (a) x-y plot based on n-heptane, the most volatile component. (b) T-x bubble-point plot. (c) A,B and K-values plotted against temperature. (d) x-y plot based on the average A,B. (e) Comparison of x-y and T-x-y plots with experimental data. Analysis: (a) To calculate y-x and T-x-y curves from vapor pressure data, using Raoult's and Dalton's laws. Eq. (2-44 ) applies, as well as the sum of the mole fractions in the phases in equilibrium. Thus, s yA PA T KA = = xA P , , s yB PB T KB = = xB P (1, 2) (3, 4) yA + y B = 1 xA + x B = 1 Equations (1) to (4) can be reduced to the following equations for the mole fractions of n-heptane (A) in terms of the K-values: xA = 1 - KB KA - KB , y A = KA x A (5, 6) If the given vapor pressure data in Exercise 4.8 for toluene, and this exercise for n-heptane are fitted to Antoine equations, we obtain: PAs = exp 15.7831 - PBs = exp 17.2741 - 2855.27 T + 213.64 3896.3 T + 255.67 (7) (8) Where vapor pressure is in torr and temperature is in oC. Solving, Eqs. (1) to (8), Analysis: (a) (continued) Exercise 4.9 (continued) KA KB xA yA A,B T, oC Ps of A, torr Ps of B, torr 98.4 99.0 100.0 101.0 102.0 103.0 104.0 105.0 106.0 107.0 108.0 109.0 110.0 110.5 760.0 773.0 795.9 819.2 843.1 867.6 892.6 918.1 944.2 970.9 998.1 1026.0 1054.4 1068.9 528.7 538.3 555.2 572.5 590.2 608.4 627.1 646.2 665.8 685.9 706.5 727.5 749.1 760.1 1.0000 1.0172 1.0472 1.0780 1.1094 1.1415 1.1744 1.2080 1.2424 1.2775 1.3133 1.3500 1.3874 1.4064 0.6956 0.7083 0.7305 0.7533 0.7766 0.8006 0.8251 0.8503 0.8761 0.9025 0.9296 0.9573 0.9856 1.0001 1.000 0.944 0.851 0.760 0.671 0.585 0.501 0.418 0.338 0.260 0.184 0.109 0.036 0.000 1.000 1.438 0.961 1.436 0.891 1.434 0.819 1.431 0.745 1.428 0.668 1.426 0.588 . 1.423 0.506 1.421 0.420 1.418 0.332 1.415 0.241 1.413 0.147 1.410 0.050 1.408 0.000 1.406 From this table, an x-y plot is given below. (b) From the above table, a T-x-y plot is given below. The x-curve is the bubble-point curve, while the y-curve is the dew-point curve. (c) A graph of relative volatility and K-values as a function of temperature is given on the next page. (d) From the above table, the arithmetic average relative volatility, using the extreme values is:(A,B)avg = (1.438 + 1.406)/2 = 1.422 Analysis: (a) (continued) Exercise 4.9 (continued) Exercise 4.9 (continued) Analysis: (c) and (d) (continued) Relative Volatility and K-Values For a constant relative volatility, Eq. (4-8) applies. For A,B = 1.422, A,B xA 1.422 xA yA = = 1 + xA A,B - 1 1 + 0.422 xA Solving this equation for values of xA = 0 to 1.0 gives the following: xA yA 0 0.0000 0.1 0.1364 0.2 0.2623 0.3 0.3787 0.4 0.4867 0.5 0.5871 0.6 0.6808 0.7 0.7684 0.8 0.8505 0.9 0.9275 1 1.0000 Exercise 4.9 (continued) Analysis: (c) and (d) (continued) y-x Plot for an average relative volatility 1 0.9 Mole fraction n-heptane in vapor 0.8 0.7 0.6 0.5 0.4 0.3 0.2 0.1 0 0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1 Mole fraction n-heptane in liquid (e) Raoult's law calculations compared to experimental are as follows: T, oC 110.75 106.80 104.50 102.95 101.35 99.73 98.90 98.50 98.35 Raoult's law xA yA -0.018 -0.026 0.276 0.350 0.459 0.547 0.589 0.672 0.729 0.793 0.876 0.910 0.954 0.967 0.992 0.995 1.007 1.005 Experimental xA yA 0.025 0.048 0.129 0.205 0.250 0.349 0.354 0.454 0.497 0.577 0.692 0.742 0.843 0.864 0.940 0.948 0.994 0.993 The Raoult's law values are in very poor agreement with the experimental values. Analysis: (e) (continued) Exercise 4.9 (continued) Analysis: (e) (continued) Exercise 4.9 (continued) Comparison with Experimental Data Exercise 4.10 Subject: Continuous, single stage distillation of A and B to produce a distillate and bottoms. Given: Saturated liquid feed of 50 mol% A and 50 mol% B fed to a still at 40 mol/h. Relative volatility = A,B = 2. Bottoms rate = 30 mol/h (a) Total condenser with a reflux ratio = 1. (b) No reflux. Assumptions: Still is an equilibrium stage. Find: (a) Composition of the two products. (b) Composition of the two products. Analysis: From the definition of the relative volatility, A,B = yA x B yA 1 - xA = =2 xA y B xA 1 - yA (1) (a) Distillate = D = F - W = 40 - 30 = 10 mol/h Material balance for A: 0.5(40) = 20 = yA(10) + xA(30) (2) Solving Eqs. (1) and (2) simultaneously by eliminating yA , we obtain: 2 3 xA + 3 xA - 2 = 0 (3) Solving Eq. (3), a quadratic equation, get only one postive root: xA = 0.4575, xB = 0.5425 for the bottoms Substitution into Eq. (2), gives, yA = 0.6275, yB = 0.3725 for the distillate (b) Note that the solution to Part (a) was independent of the reflux ratio. Accordingly, the solution to Part (b) is the as for Part (a) Exercise 4.11 Subject: Distillation of an acetone (A) - water (B) mixture that is partially vaporized. Given: Feed is 57 mol% A and 43 mol% B as a liquid at 125oC and 687 kPa. It is flashed across a valve to the column pressure of 101.3 kPa, with a resulting temperature of 60oC. Vapor-liquid equilibrium data at column pressure. Enthalpy data at column conditions. Compositions of the distillate and bottoms. Assumptions: Feed is at equilibrium downstream of the feed valve. Will have to check if feed valve operates adiabatically. Given heat capacities are for the liquid and are constant. Heats of vaporization are constant. No effect of pressure on enthalpy. Find: Mole ratio of liquid to vapor in the feed downstream of the valve. Construct an H-x-y diagram. Analysis: From the equilibrium data, at 60oC, xA = 0.50 and yA = 0.85 . Take a basis of F = feed rate = 1 kmol/s. Total material balance around feed valve: F = 1 = V + L (1) Acetone material balance around feed valve: 0.57(1) = 0.85V + 0.50L Solving Eqs. (1) and (2) simultaneously, V = 0.2 kmol/s and L = 0.8 kmol/s Therefore, after the valve, moles L/moles V = 0.8/0.2 = 4 Now check whether valve is operating adiabatically. Enthalpy of liquid entering valve = 0 (as given) Enthalpy of feed after the valve, using given enthalpies = 27,200(0.2) + (-5,270)(0.8) = 1224 kJ/s Therefore, the enthalpy increases across the valve by 1224 kJ/s To construct an enthalpy diagram for 1 atm pressure, take as an enthalpy datum, A and B as liquids at 25oC. This is a different datum than that used to get the given enthalpy of the hot feed. Pure A: boils at 56.7oC. Since CP of liquid =134 kJ/kmol-K, hL at 56.7oC = 134(56.7-25)=4248 kJ/kmol-K hV at 56.7oC =4248 + latent heat = 4248 + 29750 = 33998 kJ/kmol-K Pure B: boils at 100oC. Since CP of liquid =75.3 kJ/kmol-K, hL at 100oC = 75.3(100-25)=5648 kJ/kmol-K hV at 100oC =5648 + latent heat = 5648 + 42430 = 48078 kJ/kmol-K Equilibrium liquid mixture of 50 mol% A and 50 mol% B has a bubble point at 60oC. Therefore, hL = 0.5(134)(60-25) + 0.5(75.3)(60-25) = 3663 kJ/kmol-K Equilibrium vapor mixture of 85 mol% A and 15 mol% B has dew point of 60oC. Therefore, hV = 0.85[(134)(60-25) + 29750] + 0.15[(75.3)(60-25) + 42430] = 36034 kJ/kmol-K Calculations for other equilibrium mixtures are done in a similar manner and are summarized in the following table: (2) Analysis: (continued) Exercise 4.11 (continued) xA xB hL, kJ/kmol yA yB hV, kJ/kmol T, oC 56.7 1.000 0.000 4248 1.000 0.000 33998 57.1 0.920 0.080 4151 0.944 0.056 34656 60.0 0.500 0.500 3663 0.850 0.150 36034 61.0 0.330 0.670 3408 0.837 0.163 36296 63.0 0.176 0.824 3254 0.805 0.195 36880 71.7 0.068 0.932 3703 0.692 0.308 39069 100.0 0.000 1.000 5648 0.000 1.000 48078 From this table, the h-x-y plot follows, with tie lines to connect the vapor-liquid equilibrium along the dew-point and bubble-point lines. Enthalpy-Composition Plot Exercise 4.12 Subject: Vaporizer and condenser heat duties for benzene (A) -toluene (B) mixtures, using an enthalpy-concentration diagram. Given: P = 1 atm. Vapor pressure data. Saturated liquid and vapor enthalpy data. Assumptions: Raoult's law. Find: (a) Construct an h-x-y plot. (b) Heat duty for 50 mol% vaporization of a 30 mol% A mixture, starting from liquid saturation temperature. Heat duty to condense the vapor and subcool it 10oC. Analysis: (a) First, compute the vapor and liquid equilibrium compositions at 1 atm and temperatures from 60 to 100oC using Raoult's law with the vapor pressure data. Eq. (2-44 ) applies, as well as the sum of the mole fractions in the phases in equilibrium. Thus, KA = s yA PA T = xA P , , KB = s yB PB T = xB P (1, 2) (3, 4) yA + y B = 1 xA + x B = 1 Equations (1) to (4) can be reduced to the following equations, xA = 1 - KB KA - KB , y A = KA x A (5, 6) Vapor pressure data in Exercises 4.6 for benzene, and 4.8 for toluene give Antoine equations, PAs = exp 15.5645 - 2602.34 , T + 211.271 PBs = exp 17.2741 - 3896.3 T + 255.67 (7, 8) Where vapor pressure is in torr and temperature is in oC. Solving, Eqs. (1) to (8), T, oC Ps of A, torr Ps of B, torr KA KB xA yA 80.1 759.9 290.0 0.9998 0.3816 1.000 85.0 880.8 342.7 1.1590 0.4510 0.775 90.0 1019.1 404.4 1.3409 0.5321 0.579 95.0 1173.4 474.9 1.5439 0.6249 0.408 100.0 1345.0 555.2 1.7697 0.7305 0.259 105.0 1535.0 646.2 2.0198 0.8503 0.128 110.5 1766.8 760.1 2.3248 1.0001 0.000 This covers the temperature range of co-existence of vapor and liquid. 1.000 0.899 0.776 0.630 0.459 0.259 0.000 Analysis: (a) (continued) Molecular weights are MA = 78 and MB = 92 For a given temperature, compute saturated liquid-phase mixture enthalpies in kJ/kg of mixture from, xA M A hLA + (1 - xA ) M BhLB hL = (9) xA M A + (1 - xA ) M B Similarly for the vapor, hV = yA M A hVA + (1 - yA ) M BhVB yA M A + (1 - yA ) M B (10) Exercise 4.12 (continued) Will have to interpolate and extrapolate given saturated enthalpy data. Liquid enthalpy data are linear with temperature, therefore, it is found that: hLA = 185T - 32 , . hLB = 185T - 34 . (11, 12) Vapor enthalpy data are not quite linear, but fit the following quadratic equations: hVA = 427 + 0.85T + 0.0025T 2 , T, oC xA yA hVB = 411 + 0.85T + 0.0025T 2 (hL)B , kJ/kg 114.2 123.3 132.5 141.8 151.0 160.3 170.4 (hV)A , kJ/kg 511.1 517.3 523.8 530.3 537.0 543.8 551.5 (hV)B , kJ/kg 495.1 501.3 507.8 514.3 521.0 527.8 535.5 hL , kJ/kg 116.2 124.7 133.6 142.5 151.5 160.5 170.4 (13, 14) hV , kJ/kg 511.1 515.4 519.7 523.8 527.7 531.5 535.5 80.1 85.0 90.0 95.0 100.0 105.0 110.5 1.000 0.775 0.579 0.408 0.259 0.128 0.000 1.000 0.899 0.776 0.630 0.459 0.259 0.000 (hL)A , kJ/kg 116.2 125.3 134.5 143.8 153.0 162.3 172.4 Plots of h in kJ/kg mixture as a function of saturated vapor and liquid mole fractions, and y-x are given on the next page. (b) Take a basis of 1 kmol of 30 mol% A - 70 mol% B feed mixture. Then, kg A = (0.30)(78) = 23.4 kg and kg B = (0.70)(92) = 64.4 kg or 87.8 kg total feed. Use y-x diagram to obtain compositions of vapor and liquid for 50 mol% vaporized. From the equation above Eq. (4-6), the slope of the q-line is [(V/F)-1]/(V/F) = (0.5-1.0)/0.5 = -1. The construction is shown on the y-x diagram, where the intersection with the equilibrium curve gives xA = 0.22 and yA = 0.38. The mass of liquid = (0.22)(0.5)(78) + (0.78)(0.5)(92) = 44.5 kg. The mass of vapor = 87.8 - 44.5 = 43.3 kg. On the h-x-y diagram, Point A is the saturated liquid feed with hL = 150 kJ/kg of feed. Point C is the liquid remaining after 50 mol% vaporization, withhL, = 158 kJ/kg. Since 44.5/87.8 or 0.507 of the feed is left as liquid, this is equivalent to (0.507)(158) = 80 kJ/kg feed. Point D is the vapor, with hV = 540 kJ/kg vaporized. Since 0.493 Exercise 4.12 (continued) Analysis: (b) (continued) of the feed is vaporized, this is equivalent to (0.493)(540) = 266 kJ/kg feed. Therefore, the energy required for partial vaporization = 266 + 80 - 150 = 196 kJ/kg of feed. Point B is the combined vapor and liquid phases after partial vaporization. Point E is condensed vapor as saturated liquid, with an enthalpy of 145 kJ/kg. This is equivalent to (0.493)(145) = 71 kJ/kg of feed. Therefore, the condenser duty = 266 - 71 = 195 kJ/kg feed. Point F is 10oC subcooled condensate, where the enthalpy change from saturation, based on a liquid specific heat of 1.85 kJ/kg-oC, is 1.85(10)(0.493) = 9 kJ/kg feed. Therefore, the condenser duty is now 195 + 9 = 204 kJ/kg feed. Exercise 4.12 (continued) Analysis: (a) Enthalpy Composition Diagram Exercise 4.13 Subject: Azeotrope for the chloroform-methanol system at 101.3 kPa. Given: Vapor-liquid equilibrium data from Section 13, p. 11 of Perry's Handbook, 6th edition. Find: From data, construct y-x and T-x-y plots. Azeotrope conditions Analysis: See plots below. From these plots, a minimum-boiling azeotrope occurs at 53.5oC with a composition of 65 mol% chloroform and 35 mol% methanol. Exercise 4.14 Subject: Azeotrope for the water-formic acid system at 101.3 kPa. Given: Vapor-liquid equilibrium data from Section 13, p. 14 of Perry's Handbook, 6th edition. Find: From data, construct y-x and T-x-y plots. Azeotrope conditions Analysis: See plots below. From these plots, a maximum-boiling azeotrope occurs at 107.6oC with a composition of 42 mol% water and 58 mol% formic acid. Exercise 4.15 Subject: Partial vaporization of a water (A) -isopropanol (B) mixture at 1 atm. Given: Vapor-liquid equilibrium data at 1 atm and vapor-pressure data. Find: (a) (b) (c) (d) (e) Construct T-x-y and y-x diagrams. Composition of vapor when a 60 mol% A - 40 mol% B mixture is at its bubble point. Composition of vapor and liquid for 75 mol% vaporization of mixture in Part (a). K-values and -values at 80 and 89oC. Comparison of parts (a), (b), and (c) to results from using Raoult's and Dalton's laws. Analysis: (a) The following are plots of the given equilibrium data, including the purecomponent normal boiling points. Exercise 4.15 (continued) Analysis: (a) continued Exercise 4.15 (continued) Analysis: (b) From the y-x plot on the previous page, the composition of the first bubble of vapor is 57 mol% isopropanol and 43 mol% water. See the q-line on the diagram. (c) For 75 mol% vaporization, use the inverse lever-arm rule on the T-x-y diagram or plot a q-line on the y-x diagram. For the latter, from the equation above Eq. (4-6), the slope of the q-line is [(V/F)-1]/(V/F) = (0.75-1.0)/0.75 = -0.333. The construction is shown on the y-x diagram, where the intersection with the equilibrium curve gives xA = 0.14 and yA = 0.50. (d) Can not compute the K-values or at 80oC, because this temperature is below the lowest boiling mixture, which is the azeotrope. At 89oC, the T-x-y diagram gives the following compositions from the line shown on the above diagram: yB = 0.35, yA = 0.65 xB = 0.035, xA = 0.965 From Eq. (2-19) for the definition of the K- value, KB = yB 0.35 = = 10 xB 0.035 KA = yA 0.65 = = 0.67 xA 0.965 From Eq. (2-21) for the definition of the relative volatility, , noting that at 89oC and 1 atm, isopropanol is more volatile, K 10 B,A = B = = 15 KA 0.67 (e) To calculate T-x-y curves from vapor pressure data, using Raoult's and Dalton's laws, Eq. (2-44 ) applies, as well as the sum of the mole fractions in the phases in equilibrium. Thus, s s yA PA T yB PB T KA = = , KB = = (1, 2) xA P xB P yA + y B = 1 , xA + x B = 1 (3, 4) Equations (1) to (4) can be reduced to the following equations for the mole fractions of benzene in terms of the K-values: 1 - KB , y A = KA x A (5, 6) KA - KB If the given vapor pressure data are fitted to Antoine equations, we obtain: xA = Exercise 4.15 (continued) Analysis: (e) continued 392196 . T + 230.91 8010.6 T + 353.238 PAs = exp 18.4854 - PBs = exp 25.0173 - (7) (8) Where vapor pressure is in torr and temperature is in oC. Solving, Eqs. (1) to (8), T, oC 82.5 84.0 86.0 88.0 90.0 92.0 94.0 96.0 98.0 100.0 Ps of B Ps of A 760.0 392.1 809.5 416.2 879.9 450.2 955.7 486.6 1037.3 525.3 1125.0 566.6 1219.3 610.6 1320.5 657.4 1429.1 707.2 1545.6 760.0 KB 1.0000 1.0651 1.1578 1.2575 1.3649 1.4803 1.6043 1.7375 1.8804 2.0336 KA 0.5159 0.5476 0.5924 0.6402 0.6912 0.7456 0.8035 0.8650 0.9305 1.0000 xB 1.000 0.874 0.721 0.583 0.458 0.346 0.245 0.155 0.073 0.000 yB 1.000 0.931 0.835 0.733 0.626 0.513 0.394 0.269 0.138 0.000 B-A 1.938 1.945 1.954 1.964 1.975 1.985 1.997 2.009 2.021 2.034 These results are plotted below. Raoult's law is badly in error when compared to the experimental data. For part (b), Raoult's law predicts a bubble-point vapor with an isopropanol mole fraction of 0.56. By coincidence, this compares well with the result determined with the experimental data. For part (c), however, Raoult's law predicts isopropanol mole fractions of 0.28 for the liquid and 0.43 for the vapor. These are drastically different from the values of 0.14 and 0.50, respectively from the experimental data. Raoult's law can not be used for the isopropanol-water system, for which it also fails to predict an azeotrope. Exercise 4.15 (continued) Analysis: (e) continued Exercise 4.16 Subject: Vaporization of mixtures of n-hexane (H) and n-octane (C) at 1 atm Given: T-x-y diagram in Fig. 4.3, and y-x diagram in Fig. 4.4. 100 kmol mixture. Find: Temperature, kmol of vapor, mole fractions of H in liquid and vapor at equilibrium for various flash conditions. Analysis: Let zH = mole fraction of n-hexane in the feed and = V/F. Use inverse lever-arm rule as displayed by Line DEF in Fig. 4.3. The results for parts (a) through (f) are as follows: Given (a) zH = 0.5, = 0.2 (b) zH = 0.4, yH = 0.6 (c) zH = 0.6, xC = 0.7 (d) zH = 0.5, = 0.0 (e) zH = 0.5, = 1.0 (f) zH = 0.5, T = 200oF T, oF 196 220 210 188 230 200 V, kmol 20 48.6 73.7 0.0 100 31 yH 0.80 0.60 0.70 0.84 0.50 0.77 xH 0.43 0.21 0.30 0.50 0.14 0.38 Exercise 4.17 Subject: Derivation of equilibrium flash equations for a binary mixture (1, 2). Given: Eqs. (5), (6), and (3) of Table 4.4. Find: Derive given equations for x1, x2, y1, y2, and = V/F. Analysis: First derive the equation for = V/F. From Eq. (3), Table 4.4, z1 1 - K1 1 - z1 1 - K2 + =0 1 + K1 - 1 1 + K2 - 1 Solving Eq. (1) for , and simplifying, = (1) z1 K1 - K2 / 1 - K2 - 1 - z1 1 - K1 - 1 - z1 1 - K2 = z1 1 - K1 K2 - 1 + 1 - z1 1 - K2 K1 - 1 K1 - 1 (3) Substituting Eq. (3) into Eq. (5) of Table 4.4 and simplifying gives the required equation for x1. Then use y1 = K1x1 and simplify, followed by x2 = 1 - x1 and y2 = 1- y1 . Exercise 4.18 Subject: Conditions for Rachford-Rice equation to be satisfied. Given: Eq. (3), Table 4.4, which is the Rachford-Rice equation. Find: Conditions under which the equation can be satisfied for 0 V 1. F Analysis: A necessary, but not sufficient, condition is that at least one K-value is < 1 and at least one K-value is > 1. If all K-values are > 1, the sum: zi 1 - Ki will be negative and can not be zero. i =1 1 + Ki - 1 C If all K-values are < 1, the numerator in the sum will be positive for each term. With between 0 and 1, the term (Ki - 1) will always be < 1. Therefore, the denominator will be positive also and the sum will be positive and can not be zero. Exercise 4.19 Subject: Flash vaporization of a benzene (A) - toluene (B) mixture for A-B = 2.3. Given: Feed is 40 mol% A and 60 mol% B. Find: Percent of A in the equilibrium vapor if 90% of the toluene leaves in the liquid by graphical means. A,B xA Analysis: For constant relative volatility, Eq. (4-8) applies, yA = 1+ xA A,B - 1 Solving this equation for yA as a function of xA , xA 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 yA 0.2035 0.3651 0.4964 0.6053 0.6970 0.7753 0.8429 0.9020 0.9539 A plot of the calculated equilibrium curve is given below. To use this plot for a graphical solution of the equilibrium, draw a q-line, using the following equation above Eq. (4-6), for an assumed value of = V/F and check the resulting % recovery of toluene in the liquid. Vary until the % recovery = 90%. Then compute, for the corresponding , the % recovery of benzene in the vapor. yA = -1 1 -1 1 xA + zA = xA + 0.40 (1) Analysis: (continued) Exercise 4.19 (continued) Basis: F = 100 moles, 60 moles toluene (B). Want 0.9(60) = 54 moles B in liquid. Therefore, 60 - 54 = 6 moles B in vapor. Therefore, want (nB)V = yBV = (1 - yA)100 = 6. Then compute % recovery of benzene in vapor = (nA)V/40 x 100% = yAV/40 x 100% = 2.5 yA x 100%. The following are typical values for the trial and error procedure, with the final result at the bottom. Assumed 0.3 0.2 0.15 0.142 yA 0.54 0.56 0.575 0.58 xA (nB)V , moles 0.35 13.8 0.36 8.8 0.37 6.4 0.375 6.0 % recovery of A in vapor 40.5 28.0 21.6 20.6 Exercise 4.20 Subject: Flash vaporization of a benzene (A) - toluene (B) mixture. Given: Feed is 40 mol% A and 60 mol% B. Vapor pressure data. Assumptions: Raoult's law (ideal solutions). Pressure = 1 atm. Find: Percent of A in the equilibrium vapor if 90% of the toluene leaves in the liquid. Analysis: Basis: F = 100 mole with 60 moles B and 40 moles A. Want 0.9(60) = 54 moles B in liquid. Therefore, 60 - 54 = 6 moles B in vapor. Therefore, want (nB)V = yBV = (1 yA)100 = 6. Then compute % recovery of benzene in vapor = (nA)V/40 x 100% = yAV/40 x 100% = 2.5 yA x 100%. The following trial and error procedure can be used, based on material balance and equilibrium equations: (1) Guess a temperature. (2) Read vapor pressures from Fig. 2.4 and compute K-values from Raoult's law (Eq. (3), Table 2.3), Ki = Pi s / P . (3) Solve for = V/F using the fifth equation in Exercise 4.17, zA KA - KB / 1 - KB - 1 0.40 KA - KB / 1 - KB - 1 = = KA - 1 KA - 1 (4) Solve for yA from the third equation in Exercise 4.17, y A = KA K B - KA / K B - KA (5) Compute (nB)V = (1 - yA)100 . If the value is 6, then temperature guess is correct. Otherwise, guess another T, and repeat steps (1) to (5). If 6, compute % recovery of benzene in the vapor from 2.5 yA x 100%. Guess T, oF 195 205 205.5 Ps of A, psia 20.0 23.4 23.6 Ps of B, psia 8.0 9.5 9.6 KA 1.36 1.59 1.605 KB 0.544 0.646 0.653 -0.79 0.113 0.161 yA 0.596 0.585 Moles B in vapor 4.6 6.7 By interpolation, T = 205.3oF to obtain 6 moles of B in the vapor. This corresponds to = 0.145 and yA = 0.588. From above, % recovery of benzene in the vapor = 2.5(0.588)(0.145)100% = 21.3% Exercise 4.21 Subject: Equilibrium flash of a seven-component mixture. Given: Feed mole fractions and K-values. Find: = V/F by: zi 1 - Ki (a) Rachford- Rice equation, f 1 {} = = i =1 1 + Ki - 1 C C C i =1 g1 {i , } C zi Ki (b) Alternative flash equation, f 2 {} = = i =1 1 + Ki - 1 Make plots of f{} vs. for each method and compare. (a) I 1 2 3 4 5 6 7 i =1 g2 {i , } Analysis: Calculations with a spreadsheet, for values of from 0 to 1.0 in intervals of 0.1: zF 0.0079 0.1321 0.0849 0.2690 0.0589 0.1321 0.3151 K 16.2 5.2 2.6 1.98 0.91 0.72 0.28 f{}: =0 g{i, -0.120 -0.555 -0.136 -0.264 0.005 0.037 0.227 -0.805 0.1 -0.048 -0.391 -0.117 -0.240 0.005 0.038 0.244 -0.508 0.2 -0.030 -0.302 -0.103 -0.220 0.005 0.039 0.265 -0.345 0.3 -0.022 -0.245 -0.092 -0.204 0.005 0.040 0.289 -0.227 0.4 -0.017 -0.207 -0.083 -0.189 0.005 0.042 0.319 -0.130 0.5 -0.014 -0.179 -0.075 -0.177 0.006 0.043 0.354 -0.042 0.6 -0.012 -0.158 -0.069 -0.166 0.006 0.044 0.399 0.045 0.7 -0.010 -0.141 -0.064 -0.156 0.006 0.046 0.457 0.137 0.8 -0.009 -0.127 -0.060 -0.148 0.006 0.048 0.535 0.245 0.9 -0.008 -0.116 -0.056 -0.140 0.006 0.049 0.645 0.380 1 -0.007 -0.107 -0.052 -0.133 0.006 0.051 0.810 0.568 (b) I 1 2 3 4 5 6 7 zF 0.0079 0.1321 0.0849 0.2690 0.0589 0.1321 0.3151 K 16.2 5.2 2.6 1.98 0.91 0.72 0.28 f{}: =0 g{i, 0.128 0.687 0.221 0.533 0.054 0.095 0.088 0.805 0.1 0.051 0.484 0.190 0.485 0.054 0.098 0.095 0.457 0.2 0.032 0.373 0.167 0.445 0.055 0.101 0.103 0.276 0.3 0.023 0.304 0.149 0.412 0.055 0.104 0.113 0.159 0.4 0.018 0.256 0.135 0.383 0.056 0.107 0.124 0.078 0.5 0.015 0.222 0.123 0.357 0.056 0.111 0.138 0.021 0.6 0.013 0.195 0.113 0.335 0.057 0.114 0.155 -0.018 0.7 0.011 0.174 0.104 0.316 0.057 0.118 0.178 -0.041 0.8 0.010 0.158 0.097 0.299 0.058 0.123 0.208 -0.049 0.9 0.009 0.144 0.090 0.283 0.058 0.127 0.251 -0.038 1 0.008 0.132 0.085 0.269 0.059 0.132 0.315 0.000 The values of f{} are plotted on the next page, where it is observed that the Rachford-Rice and Alternative equations give the same result of = 0.55. However, the alternative equation also has a trivial root at = 1.0. With a Newton procedure, the alternative equation may converge to the trivial root. Therefore, the Rachford-Rice equation is preferred because of its uniqueness. Analysis: (continued) Exercise 4.21 (continuous) Exercise 4.22 Subject: Equilibrium flash of a hydrocarbon mixture. Given: 100 kmoles of 25 mo1% nC4 , 40 mol% nC5, and 35 mol% nC6 . K-values in Fig. 2.8 Assumptions: Amounts are per hour. Find: Pressure and liquid and vapor compositions for equilibrium at 240oF to recover, in the liquid phase, 80% of the nC6 in the feed. Analysis: For 80% recovery of nC6 , the liquid product must contain (0.35)(0.80)(100) = 28 kmole/h of nC6 . Must solve by trial and error by assuming values of pressure to obtain the Kvalues from Fig. 2.8. Then solve the Rachford-Rice equation (Eq. (3), Table 4.4), f {} = zi 1 - Ki =0 i =1 1 + Ki - 1 C for = V/F by a nonlinear solver, such as Newton's method. Compute V from Eq. (4), Table 4.4. Then solve Eqs. (5) and (6), Table 4.4 for the equililbrium vapor and liquid compositions. Solve for the liquid, L, from Eq. (7). Repeat this procedure until 28 kmol/h of nC6 are found in the equilibrium liquid as computed from nnC 6 = xnC 6 L , noting that each assumed pressure L requires an iterative procedure to solve for from the Rachford-Rice equation. The calculations are summarized in the following table: Assumed P, psia K-values: nC4 nC5 nC6 V/F L, kmol/h x of nC4 nL of nC4 , kmol/h 100 2.40 1.20 0.53 0.706 29.4 0.524 15 110 2.30 1.08 0.50 0.482 51.8 0.461 24 117 2.25 1.00 0.48 0.335 66.5 0.424 28 Therefore, the converged pressure is 117 psia. The equilibrium vapor and liquid compositions in terms of amounts are: Component nC4 nC5 nC6 . Vapor flows, kmol/h 13 13 7 Liquid flows, kmol/h 12 27 28 Exercise 4.23 Subject: Equilibrium flash vaporization of a hydrocarbon mixture. Given: Equimolar mixture of C2, C3, nC4, and nC5. K-values from Fig. 2.8 and 2.9 Find: Amounts and compositions of equilibrium liquid and vapor at 150oF and 205 psia. Conditions of T and P where 70% of C2 and no more than 5% of nC4 is in the vapor. Analysis: Take as a basis, a feed of 100 lbmol/h. From Fig. 2.8, at 150oF and 205 psia, the Kvalues are as given in the table below. Then solve the Rachford-Rice equation (Eq. (3), Table 4.4), f {} = zi 1 - Ki =0 i =1 1 + Ki - 1 C for = V/F by a nonlinear solver, such as Newton's method. Compute V from Eq. (4), Table 4.4. Then solve Eqs. (5) and (6), Table 4.4 for the equililbrium vapor and liquid compositions. The calculations are summarized in table below, which also includes other conditions of T and P to obtain 70% of C2 and no more than 5% of nC4 in the vapor. Thus, we desire (0.7)(25) = 17.5 lbmol/h of C2 in the vapor and 25 - 17.5 = 7.5 lbmol/hr of C2 in the liquid. At the same time we desire no more than (0.05)(25) = 1.25 lbmol/h of nC4 in the vapor, corresponding to 25 - 1.25 = 23.75 lbmol/h of nC4 in the liquid. In searching for these other conditions, we note that at the base conditions, 75.8% of the C2 goes to the vapor, which is very close to the desired 70%. But 30% of the nC4 also goes to the vapor, which is much higher than the desired 5%. The relative volatility of C2 to nC4 at the base conditions is: C 2 , nC4 = KC 2 KnC4 = yC 2 / xC 2 ynC4 / xnC 4 = (nC 2 )V / (nC2 ) L (nnC 4 )V / (nnC4 ) L = (19.0 / 6.0) = 7.39 (7.5 / 17.5) But we need a relative volatility of: C 2 , nC4 = (nC2 )V / (nC 2 ) L (nnC4 )V / (nnC 4 ) L = (17.5 / 7.5) = 44.3 (1.25 / 23.75) For ideal solutions, where the Raoult's law K-value applies, Eq. (2-21) combined with Eq. (3 in Table 2.3, gives relative volatility as independent of pressure and equal to the ratio of vapor pressures, which depend only on temperature. In general, as the temperature is reduced, the relative volatility increases. Assume that the hydrocarbon mixture, although not an ideal solution, Exercise 4.23 (continued) Analysis: (continued) follows these same trends. Thus, to increase the relative volatility, the pressure has little effect. We must decrease the temperature to obtain the desired relative volatility, and then adjust the pressure to obtain the required compositions. The base case and calculations leading to the desired separation are summarized in the following table: T, oF P, psia K-values: C2 C3 nC4 nC5 of C2 to nC4 % C2 to vapor % nC4 to vapor Base Case 150 205 4.1 1.5 0.56 0.215 7.3 75.8 30 -70 14.7 4.0 0.46 0.055 0.077 73 -40 14.7 7.5 1.11 0.165 0.028 45 Desired Case -40 16.4 6.7 1.00 0.148 0.025 45 70 4.9 Thus, at -40oF and 16.4 psia, the desired 70% of the ethane is found in the vapor product, with only 5% of the n-butane. The compositions of the vapor and liquid products for the base case and the desired case are as follows: Base Case: Desired Case: y l, lbmol/h x y l, lbmol/h x Component , lbmol/h , lbmol/h C2 19.0 0.44 6.0 0.11 17.5 0.69 7.5 0.10 C3 13.4 0.31 11.6 0.20 6.4 0.25 18.6 0.25 7.5 0.17 17.5 0.31 1.22 0.05 23.78 0.32 nC4 nC5 3.5 0.08 21.5 0.38 0.22 0.01 24.78 0.33 Total: 43.4 1.00 56.6 1.00 25.34 1.00 74.66 1.00 Exercise 4.24 Subject: Cooling of a reactor effluent with recycle liquid from a partial condensation. Given: Reactor effluent temperature of 1000oF and composition in lbmol/h of 2000 H2, 2000 CH4, 500 benzene, and 100 toluene. Partial condensation conditions or 100oF and 500 psia, and component K-values at these conditions. Two heat exchangers in a recycle loop. Find: (a) Composition and flow rate of vapor leaving flash drum in Fig. 4.38. (b) Proof that vapor flow rate is independent of quench rate. Analysis: (a) Assume that vapor rate is independent of quench rate. Therefore, conduct the flash calculation on just the reactor effluent at the flash drum conditions of temperature and pressure. Use the Rachford-Rice equations (Eqs. (3) and (6), Table 4.4): f {} = zi 1 - Ki =0 i =1 1 + Ki - 1 C (1) (2) yi = zi Ki 1 + Ki - 1 Nonlinear Eq. (1) is solved for = V/F , followed by calculation of V = F, and then calculations of vapor mole fractions from Eq. (2). The given input for Eq. (1) is: Component Hydrogen Methane Benzene Toluene f, lbmol/h 2,000 2,000 500 100 zi 0.4348 0.4348 0.1087 0.0217 Ki 80 10 0.010 0.004 One method for solving Eq. (1) is to use a spreadsheet to make a plot of f{} vs. in increments of 0.1 from 0.0 to 1.0. Then, use smaller increments in in the vicinity of f{}= 0 to obtain the solution. The results are shown in the two figures on the next page. The converged solution is = 0.8697. Therefore, V = 0.8697(4,600) = 4,000.6 lbmol/h. The composition of the equilibrium vapor from Eq. (2) is as follows: Component Hydrogen Methane Benzene Toluene zi 0.4348 0.4348 0.1087 0.0217 yi 0.4990 0.4925 0.0078 0.0007 xi 0.0062 0.0493 0.7820 0.1625 Exercise 4.24 (continued) Analysis: (a) (continued) Exercise 4.24 (continued) Analysis: (b) When the flash conditions of temperature and pressure are fixed, the compositions of the equilibrium vapor and liquid are independent of any recycle of equilibrium liquid or vapor. To prove this, draw a material balance envelope around the system in Fig. 4.38 as shown below. Now, the flash equations are the same as in Table 4.3, except for the energy balance, Eq. (6). But, that equation is only solved after all of the other equations are solved. Thus, the results for the compositions of the net vapor and liquid products are the same as when there is no recycle. Exercise 4.25 Subject: Partial condensation of a gas mixture at 120oF and 300 psia. Given: Gas at 392oF and 315 psia, with a composition in kmol/h of 72.53 N2, 7.98 H2, 0.13 benzene, and 150 cyclohexane. The gas is cooled and partial condensed to 120oF and 300 psia, followed by phase separation. Find: Equilibrium vapor and liquid flow rates and compositions. Analysis: The flash calculations are made conveniently with a process simulator, using an appropriate K-value correlation. The following results were obtained with CHEMCAD, using the Chao-Seader, Grayson-Streed (CSGS) method for K-values. Component Hydrogen Nitrogen Benzene Cyclohexane CSGS Ki 79.7 7.54 0.024 0.022 fi , kmol/h 72.53 7.98 0.13 150.00 i , kmol/h 70.82 6.36 0.0016 1.67 li , kmol/h 1.71 1.62 0.1284 148.33 Exercise 4.26 Subject: Rapid determination of phase condition without making a flash calculation. Given: A hydrocarbon mixture at 200oF and 200 psia, with a composition in lbmol/h of 125 C3, 200 nC4, and 175 nC5, and K-values at these conditions. Find: Phase(s) present without making a flash condition. Analysis: From Eq. (4-12), have a subcooled liquid if C i =1 zi Ki < 1 . C From Eq. (4-13), have a superheated vapor if Component C3 nC4 nC5 Total: i =1 zi < 1. Ki zi/Ki 0.122 0.432 0.673 1.227 > 1 fi 125 200 175 500 zi 0.25 0.40 0.35 1.00 Ki 2.056 0.925 0.520 zi K i 0.514 0.370 0.182 1.066 > 1 Therefore stream is partially vaporized. Both vapor and liquid phases present. Exercise 4.27 Subject: Determination of reflux-drum pressure for a specified temperature and total distillate (vapor and liquid phases) composition Given: Overhead partial condensing system of a distillation column that produces vapor distillate, liquid distillate, and liquid reflux. Of the total distillate, 10 mol% is vapor. Reflux drum temperature is 100oF, and composition of total distillate in mole fractions is 0.10 C2, 0.20 C3, and 0.70 nC4. Reflux pressure is not given, but K-values at 100oF and 200 psia are given. Assumptions: K-values are inversely proportional to pressure. Find: Pressure in the reflux drum. Analysis: As shown in Exercise 4.24, the compositions of net equilibrium vapor and liquid are independent of recycle or, in this case, reflux. Therefore, the flash equations can be applied using the total distillate composition as the feed composition. Therefore, =V/F = 0.10. The Kvalues are given by: KC2 = 2.7 200 200 200 , KC3 = 0.95 , KnC 4 = 0.34 P P P Substituting these equations into Eq. (3), Table 4.4, 200 200 200 01 1 - 0.95 . 01 1 - 0.34 . P P P f { P} = + + =0 200 200 200 1 + 01 2.7 . - 1 1 + 0.1 0.95 - 1 1 + 0.1 0.34 -1 P P P 01 1 - 2.7 . (1) Eq. (1) is a nonlinear equation that can be solved by various means. Using a spreadsheet, in a manner similar to that used to solve Exercise 4.24, we obtain P = 126 psia. Exercise 4.28 Subject: Comparison of flash calculations using three different K-value correlations. Given: A stream at 7.2oC and 2,620 kPa with the overall composition given below. Find: Phase conditions Analysis: Using the CHEMCAD process simulator, the following results are obtained using the Soave-Redlich-Kwong (SRK), Peng-Robinson (PR), and Benedict-Webb-Rubin-Starling (BWRS) correlations: K-values: Component N2 C1 C2 C3 nC4 nC5 nC6 SRK 17.5 5.73 1.07 0.327 0.098 0.036 0.0096 PR 17.6 5.71 1.12 0.337 0.102 0.032 0.0110 BWRS 16.3 6.00 0.99 0.292 0.087 0.029 0.0071 The K-values for the SRK and PR correlations are in reasonably good agreement, deviating from each other by less than 15 %. Except for C1, the BWRS correlation predicts lower values, with the biggest deviation for nC6. Product compositions: Componen t N2 C1 C2 C3 nC4 nC5 nC6 Total: SRK: f, kmol/h , kmol/h 1.0 0.70 124.0 52.92 87.6 11.13 161.6 6.81 176.2 2.31 58.5 0.25 33.7 0.05 642.60 74.17 PR: l, kmol/h , kmol/h 0.30 0.69 71.08 52.06 76.46 10.45 154.79 6.42 173.89 2.16 58.25 0.26 33.65 0.04 568.43 72.08 BWR: l, kmol/h , kmol/h 0.31 0.67 71.94 53.04 77.15 9.63 155.18 5.67 174.04 1.88 58.24 0.21 33.66 0.03 570.52 71.13 l, kmol/h 0.33 70.96 77.97 155.93 174.32 58.29 33.67 571.47 All three correlations predict about the same V/F ratio, which ranges from 0.1107 to 0.1154. Exercise 4.29 Subject: Equilibrium flash calculations at different temperatures and pressures Given: Mixture of 100 kmol of 60 mol% benzene (A), 25 mol% toluene (B), and 15 mol% oxylene (C). Sources of vapor pressure data. Assumptions: Ideal solutions using vapor pressure with Raoult's law. Find: Amounts and compositions of vapor and liquid products at: (a) 100oC and 1 atm. (b) 100oC and 2 atm. (c) 105oC and 0.1 atm. (d) 150oC and 1 atm. Analysis: Instead of Figure 2.4 for the vapor pressures of benzene and toluene and three vaporpressure data points for o-xylene, use the built-in vapor pressure data in the CHEMCAD process simulator with ideal K-values. The results are as follows: Case Vapor, kmol Liquid, kmol Vapor mol frac: Benzene Toluene o-Xylene Liquid mol frac: Benzene Toluene o-Xylene (a) 100oC, 1 atm 67.64 32.36 0.698 0.224 0.078 0.395 0.305 0.300 (b) 100oC, 2 atm. 0 100 (c)105oC, 0.1 atm. 100 0 0.60 0.25 0.15 (d) 150oC, 1 atm 100 0 0.60 0.25 0.15 0.60 0.25 0.15 Only in the Case (a), are two phases formed. At 1 atm, the bubble point is 91.3oC and the dew point is 107.5oC. Exercise 4.30 Subject: Prove that, at equilibrium, vapor is at its dew point and liquid is at its bubble point. Analysis: After equilibrium is achieved, separate the vapor from the liquid and analyze the separate phases. zi xi = (1) For the liquid: Apply Eq. (5), Table 4.4, 1 + Ki - 1 At the bubble point, = V/F = 0 and, therefore, from Eq. (1), xi = zi . Also, then, Ki xi = Ki zi = yi and, therefore, C i =1 Ki xi = C i =1 Ki zi = C i =1 yi = 1, which is the bubble-point equation, Eq. (4-12). For the vapor: Apply Eq. (6), Table 4.4, yi = Ki zi 1 + Ki - 1 (2) At the dew point, = V/F = 1 and, therefore, from Eq. (2), yi = zi . Also, then, xi = yi /Ki = zi /Ki C C yi C zi and, therefore, = = xi = 1 , which is the bubble-point equation, Eq. (4-12). i =1 Ki i =1 Ki i =1 Exercise 4.31 Subject: Bubble-point temperature of feed to a distillation column. Given: 2.8. Find: Feed at 1.72 MPa (250 psia) with a composition in kmol/h below. K-values in Fig. Bubble-point temperature. C Analysis: Iterate on temperature until the bubble-point equation, Eq. (4-12), is satisfied, i =1 Ki zi = 1 (1) For the first guess, take the temperature that gives the K-value for nC4 = 1.0, that is 225oF. This result and one for 200oF is as follows: T = 225oF T = 200oF Component fi , kmol/h zi Ki K i zi Ki K i zi C2 1.5 0.03 4.8 0.144 4.3 0.129 C3 10.0 0.20 2.1 0.420 1.9 0.380 nC4 18.5 0.36 1.0 0.360 0.81 0.292 nC5 17.5 0.34 0.44 0.150 0.34 0.116 nC6 3.5 0.07 0.21 0.015 0.15 0.011 Sum: 51.0 1.00 1.089 0.928 o By linear interpolation, T = 211 F for Eq. (1) to be satisfied. Exercise 4.32 Subject: Bubble and dew point pressures of binary mixture at constant temperature. Given: Mixture of 50 mol% benzene (A) and 50 mol% toluene (B) at 90oC (194oF). Vapor pressures from Fig. 2.4 (19.5 psia for A and 7.9 psia for B). Assumptions: Raoult's law for K-values. Find: Bubble and dew point pressures. Analysis: Substitution of Raoult's law, Eq. (3) in Table 2.3, into Eqs. (4-12) and (4-13) for the bubble and dew points, respectively, gives, Bubble point: C i =1 C i =1 Ki zi = C i =1 Pi s zi = 10 or . P C i=1 C i=1 Pi s zi = P (1) Dew point: C zi zi P = = 10 or . Ki i =1 Pi s zi 1 = s Pi P (2) Eq. (1) gives 13.70 psia for the bubble point. Eq. (2) gives 15.8 psia for the dew point. Exercise 4.33 Subject: Bubble point, dew point, and flash of a water (W) - acetic acid (A) mixture. Given: Equimolar mixture of W and A at 1 atm. Correlations of liquid-phase activity coefficients for W and A as a function of liquid-phase mole fractions and temperature: Assumptions: Modified Raoult's law, Eq. (4), Table 2.3 applies. Find: Dew point, bubble point, and equilibrium vapor and liquid at a temperature halfway between the bubble and dew points. Analysis: The Rachford-Rice flash equations can be used from Table 4.4: f {} = zi 1 - Ki =0 i =1 1 + Ki - 1 C (1) (2) (3) zi Ki 1 + Ki - 1 zi xi = 1 + Ki - 1 yi = where, zW = 0.5 and zA = 0.5 and the modified Raoult's law is: iL Pi s Ki = P (4) Antoine vapor pressure equations are given in Perry's Handbook for water and acetic acid: 1730.63 T ( C ) + 233.426 1936.01 log PAs = 8.02100 - o T ( C ) + 258.451 s log PW = 8.07131 - o (5) (6) The equations for the liquid-phase activity coefficients, given in the Chemical Engineering Science article of 1967 by Sebastiani and Lacquaniti, are incorrect. They are of the RedlichKister form (see Walas, S. M., "Phase Equilibria in Chemical Engineering", Butterworth, 1985, page 184) and should be: 2 log W = xA B + C 4 xW - 1 + C xW - xA 6 xW - 1 2 log A = xW B + C 4 xW - 3 + C xW - xA 6 xW - 5 (7) (8) Exercise 4.33 (continued) Analysis: (continued) where, 64.24 T ( K) 43.27 B = 01735 - . T ( K) C = 01081 . A = 0.1182 + (9) (10) (11) Since P = 1 atm and the normal boiling points of water and acetic acid are 100oC and 118.1oC, respectively, it might be expected that the dew and bubble point of the mixture would be in the vicinity of 100oC, unless the liquid-phase activity coefficients are much different from 1. To check this, activity coefficients are computed from Eqs. (7) and (8) with a spread sheet at 100oC, with the following result as a plot. It is seen that in the vicinity of mole fractions equal to 0.5, the coefficients are not large, but are about 1.2. Exercise 4.33 (continued) Analysis: (continued) At the bubble point, = V/F = 0, and Eq. (1), combined with (4), becomes: f {T} = zW 1 - s W PW {T} P s {T} + zA 1 - A A =0 P P (12) Also, at the bubble point, xi = zi = 0.5. Then, the only unknown in Eq. (12) is T. Solving nonlinear Eq. (12), with Eqs. (5) to (11), by trial and error with a spreadsheet, starting from a guess of T = 100oC, quickly leads to a bubble-point temperature of 101.6oC. The composition of the vapor bubble is obtained from Eq. (2), which at the bubble point reduces to yi = xiKi = ziKi.. The K-values at the bubble point are computed to be KW = 1.364 and KA = 0.636, giving yW = 0.682 and yA = 0.318. At the dew point, = V/F =1, and Eq. (1), combined with (4), becomes: f{xW, xW, T} = zW P P - 1 + zA -1 = 0 s W {xW , xA , T} PW {T} A {xW , xA , T} PAs {T} where because yi = zi = 0.5, T , xW , and xA = (1 - xW) are left as unknowns. The liquid phase mole fractions are from Eq. (3), xi = zi /Ki . Solving these equations by trial and error with a spread sheet, starting from T = 105oC, xW = 0.4 and xA = 0.6, quickly leads to a dew-point temperature of 105.8oC. The K-values at the dew point are computed to be KW = 1.578 and KA = 0.732, with xW = 0.3169 and xA = 0.6832. The equilibrium flash calculation is carried out at T = (101.6 + 105.8)/2 = 103.7oC. In this case, the values of , xW, and xA are computed from Eqs. (1) and (3), where the vapor pressures are computed from Eqs. (5) and (6) to be 867 torr for W and 473 for A. Values of yW and yA are obtained from Eq. (2). Using, again, a spreadsheet with a trial and error procedure, the following result is quickly obtained: V/F = 0.49 xW = 0.4100 xA = 0.5900 yW = 0.5937 yA = 0.4063 Exercise 4.34 Subject: Bubble point, dew point, and flash of a toluene (1) - n-butanol (2) mixture. Given: Feed of z1 =0.4 and z2 =0.6 at 1 atm. Liquid-phase activity coefficients for 1 and 2 as a function of liquid-phase mole fractions from the van Laar equations. Assumptions: Modified Raoult's law, Eq. (2-72) applies. Find: Dew point, bubble point, and equilibrium vapor and liquid at a temperature halfway between the bubble and dew points. Analysis: The Rachford-Rice flash equations can be used from Table 4.4: f {} = zi 1 - Ki =0 i =1 1 + Ki - 1 C (1) zi Ki (2) 1 + Ki - 1 zi xi = (3) 1 + Ki - 1 The modified Raoult's law is: iL Pi s Ki = (4) P Antoine vapor pressure (in torr) equations are obtained by fitting the vapor pressure data for toluene that are given in Exercise 4.8 and from Perry's Handbook for n-butanol: yi = P1s = exp 17.2741 - log P2s = 7.36366 - 3896.3 T ( C ) + 255.67 o o (5) (6) 1305198 . T ( C ) + 173.427 The van Laar equations, Table 2.9, with the given constants are: ln 1 = 0.855 0.855x1 1+ 1306 x2 . 1306 . 1306 x2 . 1+ 0.855x1 (7) ln 2 = (8) Exercise 4.34 (continued) Analysis: (continued) Since P = 1 atm and the normal boiling points of toluene and n-butanol are 110.8oC and 117oC, respectively, it might be expected that the dew and bubble point of the mixture would be in the vicinity of 110oC, unless the liquid-phase activity coefficients are much different from 1. To check this, activity coefficients are computed from Eqs. (7) and (8), with a spread sheet, with the following result as a plot. It is seen that in the vicinity of mole fractions equal to 0.5, the coefficients are not large, but are about 1.3. At the bubble point, = V/F = 0, and Eq. (1), combined with (4), becomes: f {T} = z1 1 - 1 P s {T} P s {T} 1 + z2 1 - 2 2 =0 P P (9) Exercise 4.34 (continued) Analysis: (continued) Also, at the bubble point, x1 = z1 = 0.4 and x2 = z2 = 0.6. Then, the only unknown in Eq. (9) is T. Solving nonlinear Eq. (9), by trial and error with a spreadsheet, starting from a guess of T = 100oC, quickly leads to a bubble-point temperature of 106.9oC. The composition of the vapor bubble is obtained from Eq. (2), which at the bubble point reduces to yi = xiKi = ziKi.. The K-values at the bubble point are computed to be K1 = 1.363 and K2 = 0.758, giving y1 = 0.545 and y2 = 0.455. At the dew point, = V/F =1, and Eq. (1), combined with (4), becomes: f{x1, x2, T} = z1 P P - 1 + z2 -1 = 0 s 1 {x1 , x2 } P {T} 2 {x1 , x2 } P2s {T} 1 where because y1 = z1 = 0.4 and y2 = z2 = 0.6, x1 , and x2 =(1 - x1) are left as unknowns. The liquid phase mole fractions are from Eq. (3), xi = zi /Ki . Solving these equations by trial and error with a spread sheet, starting from T = 105oC, x1 = 0.2 and x2 = 0.8, quickly leads to a dewpoint temperature of 109.7oC. The K-values at the dew point are computed to be K1 = 1.793 and K2 = 0.772, with x1 = 0.2231 and x2 = 0.7769. The equilibrium flash calculation is carried out at T = (109.7 + 106.9)/2 = 108.3oC. In this case, the values of , x1, and x2 are computed from Eqs. (1) and (3), where the vapor pressures are computed from Eqs. (5) and (6) to be 714 torr for 1 and 539 torr for 2. Values of y1 and y2 are obtained from Eq. (2). Using, again, a spreadsheet with a trial and error procedure, the following result is quickly obtained: V/F = 0.604 x1 = 0.2949 x2 = 0.7051 y1 = 0.4689 y2 = 0.5311 Exercise 4.35 Subject: Bubble point, dew point, and azeotrope of an ethyl acetate (A) - ethyl alcohol (E) mixture. Given: Liquid mixture of 80 mol% A - 20 mol% E at 101.3 kPa (1 atm). Liquid-phase activity coefficients for A and E as a function of liquid-phase mole fractions from the van Laar equations. Assumptions: Modified Raoult's law, Eq. (2-72) applies. Find: (a) Bubble-point temperature and vapor composition. (b) Dew point. (c) Temperature and composition of possible azeotrope. Analysis: The Rachford-Rice flash equations can be used from Table 4.4: f {} = zi 1 - Ki =0 i =1 1 + Ki - 1 C (1) (2) yi = xi = zi Ki 1 + Ki - 1 zi (3) 1 + Ki - 1 The modified Raoult's law from Eq. (2-72) is: iL Pi s Ki = (4) P Antoine vapor pressure (in torr) equations are obtained from Section 13 of Perry's Handbook: log PAs = 7.10179 - 1244.951 T ( C ) + 217.881 1281590 . log PEs = 7.58670 - o T ( C ) + 193.768 o (5) (6) The van Laar equations, Table 2.9, with the given constants are: ln A = 0.855 0.855xA 1+ 0.753xE 0.753 0.753xE 1+ 0.855xA (7) ln E = (8) Exercise 4.35 (continued) Analysis: (continued) (a) Since P = 1 atm and the normal boiling points of ethyl acetate and ethyl alcohol are o 77.1 C and 78.4oC, respectively, it might be expected that the bubble and dew points of the mixture would be in the vicinity of 70oC, unless the liquid-phase activity coefficients are much different from 1. To check this, activity coefficients are computed from Eqs. (7) and (8), with a spread sheet, with the following result as a plot. It is seen that the coefficients are not large, but are as high as 2.35. At the bubble point, = V/F = 0, and Eq. (1), combined with (4), becomes: f {T} = zA 1 - A PAs {T} P s {T} + zE 1 - E E =0 P P (9) Exercise 4.35 (continued) Analysis: (continued) Also, at the bubble point, xA = zA = 0.8 and xE = zE = 0.2. Then, the only unknown in Eq. (9) is T. Solving nonlinear Eq. (9), by trial and error with a spreadsheet, starting from a guess of T = 70oC, quickly leads to a bubble-point temperature of 73.5oC. The composition of the vapor bubble is obtained from Eq. (2), which at the bubble point reduces to yi = xiKi = ziKi.. The Kvalues at the bubble point are computed to be KA = 0.913 and KE = 1.350, giving yA = 0.730 and yE = 0.270. (b) At the dew point, = V/F =1, and Eq. (1), combined with (4), becomes: f{xA, xE, T} = zA P P - 1 + zE -1 = 0 s A {xA , xE } PA {T} E {xA , xE } PEs {T} where because yA = zA = 0.8 and yE = zE = 0.2, xA , and xE =(1 - xA) are left as unknowns. The liquid phase mole fractions are from Eq. (3), xi = zi /Ki . Solving these equations by trial and error with a spread sheet, starting from T = 70oC, xA = 0.8 and xE = 0.2, quickly leads to a dewpoint temperature of 74.3oC. The K-values at the dew point are computed to be KA = 0.922 and KE = 1.508, with xA = 0.8674 and xE = 0.1326. (c) To determine the existence of an azeotrope, where yi = xi , a series of bubble-point calculations can be made, using the procedure in part (a), starting from say, xA = 0.05 in increments of 0.05. If, in the range of xA from 0.05 to 0.95, the K-value of A switches from more than 1 to less than 1, then an azeotrope exists in this range. The calculations can then be refined. The result from a spreadsheet is a minimum-boiling azeotrope at 72.46oC with a composition of 54.4 mol% A and 45.6 mol% B. This compares to experimental values from Perry's Handbooof 71.8oC at a composition of 54 mol% A. Exercise 4.36 Subject: Bubble point, dew point, and azeotrope of a water (W) - formic acid (F) mixture. Given: Liquid mixture of 50 mol% W - 50 mol% F at 107oC. Liquid-phase activity coefficients for W and F as a function of liquid-phase mole fractions from the van Laar equations. Assumptions: Modified Raoult's law, Eq. (2-72) applies. Find: (a) Bubble-point pressure. (b) Dew point pressure. (c) Azeotropic pressure and composition at 107oC. Analysis: The Rachford-Rice flash equations can be used from Table 4.4: f {} = zi 1 - Ki =0 i =1 1 + Ki - 1 C (1) (2) yi = xi = zi Ki 1 + Ki - 1 zi (3) 1 + Ki - 1 The modified Raoult's law from Eq. (2-72) is: Ps Ki = iL i (4) P Antoine vapor pressure (in torr) equations are obtained from Section 13 of Perry's Handbook: 1730.630 T ( C ) + 233.426 1295.260 log PFs = 6.94459 - o T ( C ) + 218.00 s log PW = 8.07131 - o (5) (6) The van Laar equations, Table 2.9, with the given constants are: ln W = -0.2935 ( -0.2935) xA 1+ ( -0.2757) xE -0.2757 ( -0.2757) xE 1+ ( -0.2935) xA (7) ln F = (8) Exercise 4.36 (continued) Analysis: (continued) (a) Since T = 107oC and the normal boiling points of water and formic acid are 100oC and 100.8oC, respectively, it might be expected that the bubble and dew point pressures of the mixture would be in the vicinity of 1 atm, unless the liquid-phase activity coefficients are much different from 1. To check this, activity coefficients are computed from Eqs. (7) and (8), with a spread sheet, with the following result as a plot. It is seen that the coefficients lie between 0.7 and 1.0 At the bubble point, = V/F = 0, and Eq. (1), combined with (4), becomes: f { P} = zW 1 - s W PW {107 O C} P s {107 O C} + zF 1 - F F =0 P P (9) Exercise 4.36 (continued) Analysis: (continued) Also, at the bubble point, xW = zW = 0.5 and xF = zF = 0.5. Then, the only unknown in Eq. (9) is P. Equation (9) is linear in P and, thus, can be solved directly to give a bubble-point pressure of 719 torr. The K-values at the bubble point are 1.015 for W and 0.984. The composition of the vapor bubble is obtained from Eq. (2), which at the bubble point reduces to yi = xiKi = ziKi . This gives yW = 0.508 and yF = 0.492. (b) At the dew point, = V/F =1, and Eq. (1), combined with (4), becomes: f{xW, xF, P} = zW P P - 1 + zF -1 = 0 s o W {xW , xF } PW {107 C} F {xW , xF } PFs {107 o C} where because yW = zW = 0.5 and yF = zF = 0.5, xW , and xF =(1 - xW) are left as unknowns. The liquid phase mole fractions are from Eq. (3), xi = zi /Ki . Solving these equations by trial and error with a spread sheet, starting from P = 760 torr, xW = 0.5 and xF = 0.5, quickly leads to a dew-point pressure of 725.2 torr. The K-values at the dew point are computed to be KW = 1.006 and KF = 0.977, with xW = 0.497 and xF = 0.512. (c) To determine the existence of an azeotrope, where yi = xi , a series of bubble-point calculations can be made, using the procedure in part (a), starting from say, xW = 0.05 in increments of 0.05. If, in the range of xW from 0.05 to 0.95, the K-value of W switches from more than 1 to less than 1, then an azeotrope exists in this range. The calculations can then be refined. The result from a spreadsheet is a maximum-boiling azeotrope at 718.2 torr with a composition of 44.5 mol% W and 55.5 mol% F. Exercise 4.37 Subject: Bubble point, dew point, and equilibrium flash of a ternary mixture. Given: Mixture of 45 mol% n-hexane (1), 25 mol% n-heptane, and 30 mol% n-octane. Assumptions: Applicability of S-R-K method for estimating K-values. Find: (a) Bubble-point temperature at pressures of 5, 1, and 0.5 atm. Dew-point temperature at pressures of 5, 1, and 0.5 atm. (b) Temperature and phase compositions for flash of V/F = 0.5 at 5, 1, and 0.5 atm. (c) n-octane in the vapor if 90% of the n-hexane is vaporized at 1 atm. Analysis: The following results are obtained with the CHEMCAD simulator. (a) Bubble-point and dew-point temperatures: Pressure, atm Bubble-point T, oF Dew-point T, oF (b) 50 mol% vaporization of feed: Pressure, atm Temperature, oF Vapor mole fractions: n-Hexane n-Heptane n-Octane Liquid mole fractions: n-Hexane n-Heptane n-Octane Moles V/mole F Moles L/mole F 0.5 162 0.615 0.231 0.154 0.285 0.269 0.446 0.5 0.5 1 201 0.599 0.233 0.168 0.301 0.267 0.432 0.5 0.5 5 318 0.554 0.239 0.207 0.346 0.261 0.393 0.5 0.5 0.5 147 178 1 187 216 5 307 330 (c) By iteration on the isothermal flash calculation, for 90 mol% of n-hexane to vapor, need a temperature of 210.2oF at 1atm. This gives 80.1 mol% vaporization with 65% vaporization of n-octane. Exercise 4.38 Subject: Vaporization of column bottoms in a partial reboiler. Given: 150 kmol/h of bubble-point liquid, L1, at 758 kPa, with a molar composition of 10% propane, 40% n-butane, and 50% n-pentane leaving bottom stage of a distillation column and passing to a reboiler where all by 50 kmol/h is vaporized to VB. Assumptions: Pressure in reboiler = 758 kPa. S-R-K method for K-values and enthalpies. Find: Compositions and amounts of boilup and bottoms, B, and reboiler duty, QR, from a simulation program Analysis: Use flash module of the CHEMCAD simulator. Feed (L1) temperature is computed from a bubble-point calculation at 758 kPa to be 74.9oC. Flash conditions are P = 758 kPa and V/F = (150 - 50)/150 = 0.6667. The result is a temperature of 87.5oC with a reboiler duty of 2.22 x 106 kJ/h, and compositions as follows in terms of component flow rates: Component: Bottoms, kmol/h Boilup, kmol/h Propane 1.94 13.06 n-Butane 15.39 44.61 n-Pentane 32.67 42.33 Total 50.00 100.00 Exercise 4.39 Subject: Bubble-point and equilibrium flash temperatures for a ternary mixture. Given: Mixture at 50 psia with a composition in mole fractions of 0.005 methane, 0.595 ethane, and 0.400 n-butane. Find: (a) Bubble-point temperature. (b) Temperature and phase compositions for 25 mol% vaporization. Analysis: Instead of using Figs. 2.8 and 2.9 for K-values, use S-R-K method with the CHEMCAD simulator, with the following results: (a) Bubble-point temperature = -61oF. Use of Figs. 2.8 and 2.9 gives -60oF. (b) Flash temperature for 25 mol% vaporization = -43.6oF, with compositions: Component Vapor mole fraction Liquid mole fraction Methane 0.0179 0.0007 Ethane 0.9559 0.4747 n-Butane 0.0262 0.5246 Exercise 4.40 Subject: Heating and expansion of a hydrocarbon mixture. Given: 100 lbmol/h of a mixture at 150oF and 260 psia, with a mole fraction composition of 0.03 ethane, 0.20 propane, 0.37 n-butane, 0.35 n-pentane, and 0.05 n-hexane. Mixture is heated to 260oF at 250 psia, followed by expansion to 100 psia. Assumptions: Expansion is adiabatic. S-R-K method for K-values and enthalpies Find: Using a simulation program, find for each stream in the process, the mol% vapor, and vapor and liquid phase mole fractions. Analysis: Using the CHEMCAD process simulator, the following results are obtained. The feed is all liquid. The streams leaving the heater and the valve are all vapor. The final temperature is 235oF. Exercise 4.41 Subject: Equilibrium vapor and liquid streams leaving the feed stage of a distillation column Given: Feed stream, F, and streams VF+1 and LF-1 as summarized below. Pressure = 785 kPa. Assumptions: VF+1 is at its dew point and LF-1 is at its bubble point. Adiabatic conditions. Find: Composition and amounts of streams VF and LF. Analysis: Using the S-R-K method for K-values, with the CHEMCAD process simulator, the following results are obtained: Stream Temperature, oC Phase condition Flow rate, kmol/h: Propane n-Butane n-Pentane Total F 64.3 Liquid 32.0 64.0 64.0 160.0 VF+1 73.1 Vapor 58.8 98.0 39.2 196.0 LF-1 68.5 Liquid 15.0 45.0 40.0 100.0 VF 70.0 Vapor 69.6 89.9 34.5 194.0 LF 70.0 Liquid 36.2 117.1 108.7 262.0 Exercise 4.42 Subject: Adiabatic flash across a valve of a hydrocarbon mixture. Given: Feed mixture, of composition below, at 250oF and 500 psia. Pressure exiting valve = 300 psia. Find: (a) (b) (c) (d) Phase condition of feed. Temperature downstream of valve. Mole fraction vaporized across valve. Mole fraction compositions of vapor and liquid phases downstream of valve. Analysis: Use CHEMCAD process simulator with S-R-K method for K-values and enthalpies. Results are as follows: Stream Temperature, oF Pressure, psia Phase condition Mole fraction of feed Mole fractions: Ethylene Ethane Propylene Propane Isobutane n-Butane Feed 250 500 Liquid 1.00 0.02 0.03 0.05 0.10 0.20 0.60 Vapor from valve 207.6 300 Vapor 0.4043 0.0348 0.0491 0.0658 0.1260 0.1953 0.5290 Liquid from valve 207.6 300 Liquid 0.5957 0.0099 0.0171 0.0393 0.0823 0.2032 0.6482 Exercise 4.43 Subject and to Find: Algorithm for flash calculation when = V/F and P are specified. Given: Isothermal flash algorithm of Fig. 4-19a, and equations of Table 4.4. Analysis: Specify feed rate and composition, and values of = V/F and P. Use the isothermal flash algorithm of Fig. 4-19a as an inner loop. Guess the flash temperature and enter the inner loop. If the calculated = V/F is not the specified value, guess a new value of T = say 1.05 times the initial guess of T, and repeat the inner loop. For the next and subsequent iterations, k, apply the false position method to provide a new guess of T: T k + 2 = T k +1 + spec - k +1 T k +1 - T k / k +1 - k This assumes that T is a linear function of = V/F. Iterate until the computed = V/F is within say 0.1% of the specified value. Exercise 4.44 Subject and to Find: Algorithms for flash calculations with 6 different sets of specified variables given in the table below. Given: Isothermal flash algorithm of Fig. 4-19a, and equations of Table 4.4. Assumption: All flashes are adiabatic. Analysis: The equations to be solved for each algorithm are those for the standard adiabatic flash procedure, where the specifications are outlet P and Q = 0. Rachford-Rice Eq. (3), Table 4.4: f1 = zi 1 - Ki =0 i =1 1 + Ki - 1 C (1) Adiabatic energy balance, Eq. (4-19): f 2 = hV + (1 - )hL - hF 1,000 (2) For each of the 6 algorithms, we must choose a tear variable, the output variable for f1, and the output variable for f2. If K-values are composition-dependent, then outer loop iterations with f1 are necessary as in Fig. 4.19. Note that the specification of hF is equivalent to specifying TF or Q = 0. In some cases, it may be necessary to solve f1 and f2 simultaneously. Output Variable in Tear f1 f2 Variable TV TV PV TV PV TV hF{TF} PV hF hF{TF} TV hF hF{TF} hF Case 1 2 3 4 5 6 Specifications hF , P hF , T hF , , T , P T, P Find , , P T, P hF , P hF , T hF , As an example of one of the algorithms, consider Case 1, which is equivalent to the standard adiabatic flash specification. The algorithm is shown in diagram form on the following page. Exercise 4.44 (continued) Analysis: (continued) Exercise 4.45 Subject: Comparison of solvents for single-equilibrium-stage liquid-liquid extraction. Given: Feed, F = 13,500 kg/h of 8 wt% acetic acid (B) in water (A) at 25oC. Four solvents (C), each with a different distribution coefficient, KC, in mass fractions, xB, for acetic acid, as given in the table below, according to Eq. (2-20), where (1) is the extract of flow rate E and (2) is the raffinate of flow rate R, where for a single equilibrium stage, the raffinate is to contain only 1 wt% B ( ( K DB = xBE ) / xBR ) (1) Assumptions: Water is insoluble in the solvent and the solvent is insoluble in water. Find: The kg/h, S, of each solvent required. ( ( Analysis: In Eq. (1), xBR ) = 0.01. Therefore, xBE ) = 0.01 K DB (2) (3) A total material balance gives: An acetic acid material balance gives: F = 13,500 = E + R - S ( ( ( ( xBF ) F = (0.08)(13,500) = 1,080 = xBE ) E + xBR ) R = xBE ) E + 0.01R (4) 92 wt% of the feed is water, or (0.92)(13,500) = 12,420 kg/h. Since all of the water appears in the raffinate, which is 99 wt% water, R = 12,420/0.99 = 12,546 kg/h. Eq. (3) becomes: S = E - 954 (5) and Eq. (4) becomes: ( xBE ) E = 954.5 (6) ( Eqs. (2), (5), (6) are three equations in three unknowns: S, E, and xBE ) . For each solvent, solve ( Eq. (2) for xBE ) . Solve (6) for E. Solve (5) for S. The results are: Solvent Methyl acetate Isopropyl ether Heptadecanol Chloroform KD 1.273 0.429 0.312 0.178 ( xBE ) 0.01273 0.00429 0.00312 0.00178 E, kg/h 74.984 222,500 305,940 536,260 S, kg/h 74,030 221,500 305,000 535,300 Although methyl acetate is the best solvent, the solvent rates required are very large. To reduce the solvent rate, use a countercurrent, multiple-stage system. Exercise 4.46 Subject: Liquid-liquid extraction of ethylene glycol from water by furfural with one stage. Given: Feed, F = 45 kg, of 30 wt% ethylene glycol (B) and 70 wt% water (A). Phase equilibrium diagrams of Fig. 4.14a and 4.14e for 25oC. Find: (a) Minimum amount of solvent. (b) Maximum amount of solvent. (c) % glycol extraction and amounts of solvent-free extract and raffinate for 45 kg solvent. (d) Maximum possible glycol purity in extract. Maximum purity of water in raffinate. Analysis: For a single stage, all mixtures of feed, F, and solvent, S, lie on a straight line between these two points as shown in the following ternary diagram of Fig. 4.14a. (a) The minimum amount of solvent corresponds to the maximum solubility of the solvent in the feed. This is point M1 in the diagram below. By the inverse lever arm rule, S/F = 0.097. Therefore, Smin = 0.0972(45) = 4.4 kg. In this case, no extract is obtained. (b) The maximum amount of solvent corresponds to maximum solubility of the feed in the solvent. This is point M2 in the diagram below. By the inverse lever arm rule, S/F = 11.15. Thus, Smax = 11.2(45) = 504 kg. In this case, no raffinate is obtained. Exercise 4.46 (continued) Analysis: (continued) (c) With 45 kg of solvent, S/F = 1. Therefore, the mixing point, M3 is at the mid point between F and S in the diagram below. A tie line drawn through point M3 determines the raffinate, R, and the extract, E. By total material balance, R + E = F + S = 45 + 45 = 90. By the inverse lever arm rule, R/E = 0.61. Combining these equations gives: E = 55.9 kg, R = 34.1 kg. From the diagram, the composition of the extract is: 19 wt% B, 76 wt % C, and 5 wt% A. Therefore, the extract contains (0.19)(55.9) = 10.62 kg B. The composition of the raffinate is 8 wt% B, 84 wt% A, and 8 wt% C. The % extraction of glycol (B) = 10.62/[(0.30)(45)] x 100% = 78.7%. The amount of solvent-free extract = (55.9)(1 - 0.76) = 13.4 kg. The amount of solvent-free raffinate = (34.1)(1 - 0.08) = 31.4 kg. Exercise 4.46 (continued) Analysis: (continued) (d) The maximum possible glycol purity in the extract occurs when the minimum amount of solvent is used, as discussed in part (a), giving raffinate R1. In this case, have say just one drop of extract, corresponding to point E1, which connects to R1 by a tie line in the diagram below. At point E1, have 48 wt% glycol. The maximum purity of water in the raffinate occurs when the maximum amount of solvent is used, as discussed in part (b), giving extract E2. In this case, have say just one drop of raffinate, corresponding to point R2, which connects to E2 by a tie line in the diagram below. At point R2, have 90 wt% water. Exercise 4.47 Subject: Representation of the composition of a ternary mixture on a triangular diagram. Given: A triangular diagram where each vertex represents a pure component. Assumptions: Consider the case of an equilateral triangular diagram. Prove: The composition of any point inside the triangle is proportional to the length of the respective perpendicular drawn from the point to the side of the triangle opposite the vertex in question. Analysis: In the triangular diagram below, the pure components are E, R, and S. A mixture of these components is represented by the point M. For the equilateral triangle shown, the sum of the lengths of the three perpendiculars drawn from an interior point, such as M, to the three sides equals the altitude, is the same length from each side. Therefore, divide each altitude into 100 divisions and number these divisions starting with 0 at the base to 100 at the apex. Thus, if the divisions represent wt% (or mol%), the sum of the perpendiculars equals 100%. In the diagram below, each division is 10%. The perpendicular s measures 28% and is the composition of S. The perpendicular r measures 18% and is the composition of R. The perpendicular e measures 54% and is the composition of E. This proof is extended to a triangle of any shape on pages 11 to 18 of "Chemical Process Principles, Part I" by Hougen, Watson, and Ragatz. Exercise 4.48 Subject: Liquid-liquid extraction of acetic acid from chloroform by water at 18oC and 1 atm. Given: Equilibrium data for the ternary mixture Assumptions: Equilibrium stages. Find: (a) Compositions and weights of raffinate and extract when 45 kg of a 35 wt% chloroform (C) and 65 wt% acetic acid (A) feed mixture (F) is extracted with 22.75 kg of water (S) in a single-stage extraction. (b) Compositions and weights of raffinate and extract if the raffinate from part (a) is extracted again with half its weight of water. (c) Composition of the raffinate from part (b) if all the water is removed from it. Use a right-triangle diagram of the equilibrium data, which is easily produced Analysis: with a spreadsheet. (a) In the diagram on the next page, the feed and solvent are represented by points F and S, respectively. The mixing point is M1, which is the sum of the feed and solvent (67.75 kg) with an overall composition of 33.58 wt% S, 23.25 wt% C, and 43.17 wt% A. A tie line passing through point M1 locates, on the equilibrium curve, the extract E1 and the raffinate R1. If the inverse lever arm rule is used to obtain E1 and R1, and their compositions are read from the diagram, the following results are obtained: Feed Solvent Extract Raffinate F S E1 R1 Amount, kg 45 22.75 55 12.75 Composition, wt%: Chloroform 35 0 11.5 74.5 Acetic acid 65 0 48.0 22.0 Water 0 100 40.5 3.5 (b) If R1 is mixed with half its weight of solvent (6.375 kg), the mixing point is M2, shown in the diagram on the next page, with an overall composition of 14.6 wt% A, 49.7 wt% C, and 35.7 wt% S. . A tie line passing through point M2 locates, on the equilibrium curve, the extract E2 and the raffinate R2. If the inverse lever arm rule is used to obtain E2 and R2, and their compositions are read from the diagram, the following results are obtained: Feed Solvent Extract Raffinate Water-free (R1) S E2 R2 Raffinate Amount, kg 12.75 6.375 8.895 10.23 10.10 Composition, wt%: Chloroform 74.5 0 1.0 91.0 93.2 Acetic acid 22.0 0 24.0 6.7 6.8 Water 3.5 100 74.9 1.3 0.0 (c) The water-free raffinate is included in the above table. Exercise 4.48 (continued) Exercise 4.49 Subject: L-L extraction of acetic acid from water by isopropyl ether at 25oC and 1 atm. Given: Equilibrium data for the ternary mixture Assumptions: Equilibrium stages. Find: (a) Compositions and weights of raffinate and extract when 100 kg of a 30 wt% acetic acid (A) and 70 wt% water (W) feed mixture (F) is extracted with 120 kg of isopropyl ether (E) in a single-stage extraction. Weight % of A in the extract if E is removed. (b) Compositions and weights of raffinate and extract if 52 kg of A and 48 kg of W are contacted with 40 kg of E Analysis: Use a right-triangle diagram of the equilibrium data, which is easily produced with a spreadsheet. (a) In the diagram on the next page, the feed and solvent are represented by points F1 and S, respectively. The mixing point is M1, which is the sum of the feed and solvent (220 kg) with an overall composition of 13.6 wt% A, 31.8 wt% W, and 54.6 wt% E. A tie line passing through point M1 locates, on the equilibrium curve, the extract E1 and the raffinate R1. If the inverse lever arm rule is used to obtain E1 and R1, and their compositions are read from the diagram, the following results are obtained, including the amount and composition of the ether-free extract. Feed Solvent Extract Raffinate Ether-free F1 S E1 R1 Extract Amount, kg 100 120 133.4 86.6 16.1 Composition, wt%: Water 70 0 2.9 76.3 24.2 Acetic acid 30 0 9.1 20.6 75.8 Isopropyl ether 0 100 88.0 3.1 0.0 (b) In the diagram on the next page, the feed and solvent are represented by points F2 and S, respectively. The mixing point is M2, which is the sum of the feed and solvent (140 kg) with an overall composition of 37.1 wt% A, 34.3 wt% W, and 28.6 wt% E. A tie line passing through point M2 locates, on the equilibrium curve, the extract E2 and the raffinate R2. If the inverse lever arm rule is used to obtain E2 and R, and their compositions are read from the diagram, the following results are obtained: Feed Solvent Extract Raffinate F2 S E2 R2 Amount, kg 100 40 55.0 85.0 Composition, wt%: Water 48 0 9.3 50.5 Acetic acid 52 0 28.9 42.5 Isopropyl ether 0 100 61.8 7.0 Exercise 4.49 (continued) Exercise 4.50 Subject: Separation of paraffins from aromatics by liquid-liquid extraction with 90 mol% diethylene glycol (DEG) and 10 mol% water at 325oF and 300 psia. Given: 280 lbmol/h of 42.86 mol% nC6 , 28.57 mol% nC7 , 17.86 mol% benzene, and 10.71 mol% toluene. 500 lbmol/h of solvent solution. Assumptions: One equilibrium stage. UNIFAC L/L method for activity coefficients. Find: Flow rates and compositions of extract and raffinate phases. Selectivity of DEG. Analysis: Using the LLVF (three-phase flash) model of CHEMCAD, the result is: Flow rate, lbmol/h: Feed Solvent Extract Raffinate Diethylene glycol 0 450 447.5 2.5 Water 0 50 49.8 0.2 nC6 120 0 18.7 101.3 nC7 80 0 8.8 71.2 Benzene 50 0 31.6 18.4 Toluene 30 0 16.5 13.5 Total 280 500 572.9 207.1 DEG is more selective for aromatics. Exercise 4.51 Subject: Liquid-liquid extraction of organic acids from water with ethyl acetate (EA). Given: 110 lbmol/h of feed containing 5, 3, and 2 lbmol/h, respectively, of formic, acetic, and propionic acids in water. 100 lbmol/h of ethyl acetate solvent. One equilibrium stage. Assumptions: UNIFAC method for liquid-phase activity coefficients. T = 25oC. Find: Flow rates and compositions of extract and raffinate. Order of selectivity of EA. Analysis: Using the LLVF (three-phase flash) model of the CHEMCAD simulation program: Flow rate, lbmol/h: Feed Solvent Extract Raffinate Ethyl acetate 0 100 99.6 0.4 Water 100 0 14.7 85.3 Formic acid 5 0 2.1 2.9 Acetic acid 3 0 2.6 0.4 n-Propionic acid 2 0 1.9 0.1 Total 110 100 120.9 89.1 The order of selectivity is n-propionic acid highest, then acetic acid, and then formic acid. Exercise 4.52 Subject: Given: Leaching of oil from soybean flakes by a hexane solvent 100,000 kg/h of soybean flakes, containing 19% by weight oil. 200,000 kg/h of a hexane solvent. Oil content of the flakes to be reduced to 0.5 wt% in a single stage. Expected contents of the variable underflow in Example 4.9. Assumptions: Ideal leaching stage so that compositions of overflow and underflow liquids are equal. Find: Using an algebraic method, the compositions and flow rates, of overflow and underflow, and the percentage of oil in the feed that is recovered in the overflow. Analysis: The flakes contain (0.19)(100,000) = 19,000 kg/h of oil and (100,000 19,000) = 81,000 kg/h of oil-free solids. However, all of the oil is not leached. For convenience in the calculations, lump the unleached oil with the oil-free solids to give an effective solids. The flow rate of unleached oil = (81,000)(0.5/99.5) = 407 kg/h. Therefore, take the solids flow rate as (81,000 + 407) = 81,407 kg/h. Take the oil in the feed as just the amount that is leached or (19,000 407) = 18,593 kg/h. Therefore, in the feed, F, YA = (81,407/18,593) = 4.38 XB = 1.0 The sum of the liquid solutions in the underflow and overflow include 200,000 kg/h of hexane and 18,593 kg/h of leached oil. Therefore, for the underflow and overflow, XB = YB = [18,593/(200,000 + 18,593)] = 0.0851 This is a case of variable solution underflow. Using data in the table that accompanies Example 4.9, and the conversion shown there of that table to XA as a function of XB, we obtain XA = 2.09. Using the algebraic method, since the flow rate of solids in the underflow = 81,407 kg/h, the flow rate of liquid in the underflow = 81,407/2.09 = 38,950 kg/h. The total flow rate of underflow is U = 81,407 + 38,950 = 120,357 kg/h. By mass balance, the flow rate of overflow = 300,000 120357 = 179,643 kg/h. Exercise 4.52 (continued) Now compute the compositions of the underflow and overflow. In the overflow, with XB = 0.0851, the mass fractions of solute B and solvent C are, respectively, 0.0851 and (1 0.0851) = 0.9149. In the underflow, using XA = 2.09 and XB = 0.0851, the mass fractions of solids, B, and C, are, respectively, [2.09/(1 + 2.09)] = 0.676, 0.0851(1 0.676) = 0.0276, and (1 0.676 0.0276) = 0.2964. From above, the oil flow rate in the feed is 19,000 kg/h. The oil flow rate in the overflow = YBV = 0.0851(179,643) = 15,288 kg/h. Thus, the percentage of the oil in the feed that is recovered in the overflow = 15,288/19,000 = 0.805 or 80.5%. Exercise 4.53 Subject: Leaching of Na2CO3 from a solid by water. Given: Lo = 3,750 kg/h of a solid containing 1,350 kg/h Na2CO3, contacted with So = 4,000 kg/h of water. Assumptions: Ideal leaching stage so that compositions of overflow and underflow liquid are equal. Underflow contains 40 wt% water on a solute-free basis. Find: Compositions and flow rates, L1 and S1, respectively, of overflow and underflow. Analysis: Let x = mass fraction of Na2CO3 (solute). Lo + So = 3,750 + 4,000 = 7,750 = L1 + S1 (a) Total balance: Na2CO3 balance: x Lo Lo = 1,350 = x L1 L1 + xS1 S1 Insoluble solids balance: (3,750 - 1,350) = 2,400 = (1 - 0.40) 1 - xS1 S1 = 0.60 1 - xS1 S1 Equilibrium: (1) (2) (3) x L1 = xS1 S1 S1 - (0.60) 1 - xS1 S1 = xS1 0.4 + 0.6 xS1 (4) Eqs. (1) to (4) can be reduced to one quadratic equation in xS1 . Solving, xS1 = 01189 ; x L1 = 0.2523 ; S1 = 4,540 kg / h ; and L1 = 3,210 kg / h . The material balance may be summarized as follows: Component Insoluble oxide Na2CO3 Water Total Lo , Solids, kg/h 2.400 1,350 0 3,750 So , Solvent, kg/h 0 0 4,000 4,000 L1 , Overflow, kg/h 0 810 2,400 3,210 S1 , Underflow, kg/h 2,400 540 1,600 4,540 The % extraction of Na2CO3 = 810/1,350 x 100% = 60% Exercise 4.54 Subject: Leaching of Na2CO3 from a solid by water. Given: Lo = 3,750 kg/h of a solid containing 1,350 kg/h Na2CO3, contacted with So = 4,000 kg/h of water. Assumptions: Ideal leaching stage so that compositions of overflow and underflow liquid are equal. Underflow contains 40 wt% water on a solute-free (dissolved Na2CO3-free) basis. Only 80% of the Na2CO3 is dissolved. Find: Compositions and flow rates, L1 and S1, respectively, of overflow and underflow. Analysis: Because only 80% of the Na2CO3 is dissolved, the underflow will contain (3,750 1,350) = 2,400 kg/h of insoluble oxide plus (0.20)(1,350) = 270 kg/h of solid Na2CO3 or a total of 2,670 kg/h of total solids. Let: V = kg/h of overflow L = kg/h of liquid in the underflow. Total underflow = L + 2,670 kg/h y = mass fraction of dissolved Na2CO3 in overflow or underflow liquid (equilibrium) Total mass balance: Na2CO3 mass balance: Lo + So = 4,000 + 3,750 = 7,750 = V + L + 2,670 1,350 = 270 + y(V + L) (1) (2) Solving Eqs. (1) and (2), V + L = 5,080 kg/h and y = 0.2126 Flow rate of underflow on a solute-free basis = (1 - y)L + 2,670 = 0.7874 L + 2,670 kg/h Therefore the flow rate of water in the underflow = 0.40(0.7874 L + 2,670) Water mass balance: 4,000 = (1 - 0.2126)V + 0.40(0.7874 L + 2,670) (3) Solving linear Eqs. (1) and (3), L = 2,260 kg/h and V = 2,820 kg/h The material balance may be summarized as follows: Component Insoluble oxide Insoluble Na2CO3 Soluble Na2CO3 Water Total Lo , Solids, kg/h 2.400 1,350 0 3,750 So , Solvent, kg/h 0 0 0 4,000 4,000 Overflow, kg/h 0 0 600 2,220 2,820 Underflow, kg/h 2,400 270 480 1,780 4,930 The % extraction of Na2CO3 = 600/1,350 x 100% = 44.4% Exercise 4.55 Subject: Crystallization from a mixture of benzene and naphthalene. Given: 6,000 lb/h of 40 wt% benzene and 60 wt% naphthalene in liquid state at 50oC, which is cooled to 15oC. Assumptions: Equilibrium at 15oC. Equilibrium phase diagram of Fig. 4.23. Find: Flow rates and compositions of crystals and mother liquor. Analysis: The feed is (0.60)(6,000) = 3,600 lb/h of naphthalene and 2,400 lb/h of benzene. In the figure below: A is the feed solution at 50oC. It lies in the homogeneous solution region. C is the feed solution at 15oC. It now lies in the two-phase region and separates into naphthalene crystals at D and mother liquor at B. From the diagram, the mother liquor at 15oC contains 31.5wt% naphthalene and, therefore, 68.5 wt% benzene. Because all of the benzene is in the mother liquor, the mother liquor flow rate is 2,400/(0.685) = 3,500 lb/h. Therefore, the naphthalene crystals flow rate is 6,000 - 3,500 = 2,500 lb/h. The % crystallization of naphthalene is 2,500/3,600 x 100% = 69.4%. Exercise 4.56 Subject: Crystallization from a mixture of benzene and naphthalene. Given: 6,000 lb/h of 40 wt% benzene and 60 wt% naphthalene in liquid state at 50oC to be cooled to obtain crystals of naphthalene. Assumptions: Equilibrium upon cooling, using phase diagram of Fig. 4.23. Find: Temperature necessary to crystallize 80% of the naphthalene. Flow rates and compositions of crystals and mother liquor. Analyze: The feed is (0.60)(6,000) = 3,600 lb/h of naphthalene and 2,400 lb/h of benzene. If 80% of the naphthalene is crystallized, the crystals flow rate will be (0.8)(3,600) = 2,880 lb/h. This leaves (3,600 - 2,880) = 720 lb/h of naphthalene in the mother liquor, with 2,400 lb/h of benzene. The total flow rate of mother liquor is 720 + 2,400 = 3,120 lb/h. Therefore, the naphthalene concentration in the mother liquor is 720/3,120 x 100% = 23 wt%. In the figure below, A is the feed solution at 50oC. It lies in the homogeneous solution region C is the feed solution in the two-phase region, which separates into naphthalene crystals at D and mother liquor at B, with the 23 wt% solubility. It is seen that a temperature of 3oC is needed to achieve a naphthalene solubility of 23 wt%. Exercise 4.57 Subject: Crystallization from a mixture of benzene and naphthalene. Given: 10,000 kg/h of 90 wt% benzene and 10 wt% naphthalene in liquid state at 30oC to be cooled to 0oC to obtain crystals. Assumptions: Equilibrium upon cooling to 0oC, using phase diagram of Fig. 4.23. Find: Flow rates and compositions of crystals and mother liquor. Analyze: The feed is (0.10)(10,000) = 1,000 kg/h of naphthalene and 9,000 kg/h of benzene. In the figure below: A is the feed solution at 30oC. It lies in the homogeneous solution region. C is the feed solution at 0oC. It now lies in the two-phase region and separates into benzene crystals at B and mother liquor at D. From the diagram, the mother liquor at 0oC contains 15 wt% naphthalene and, therefore, 85 wt% benzene. Because all of the naphthalene is in the mother liquor, the mother liquor flow rate is 1,000/(0.15) = 6,667 kg/h. Therefore, the benzene crystals flow rate is 10,000 - 6,667 = 3,333 kg/h. The % crystallization of benzene is 3,333/9,000 x 100% = 37.0%. Exercise 4.58 Subject: Crystallization of Na2SO4 from an aqueous solution by cooling. Given: 1,000 lb/h of Na2SO4 dissolved in 4,000 lb/h of H2O at 50oC. Assumptions: Equilibrium according to phase diagram of Fig. 4.24. Find: Temperature at which crystallization begins. Temperature to obtain 50% crystallization of the sulfate. Hydrate form of the crystals. Analysis: In the diagram below, the feed of 20 wt% sulfate, at point A, is in the homogeneous solution region at 50oC. Crystallization begins at point B, corresponding to a temperature of 24oC. As the temperature is lowered further, crystals of the decahydrate, Na 2SO 4 10H 2 O form. For 50% crystallization of Na2SO4 , must crystallize 0.5(1,000) = 500 lb/h. For the decahydrate, with molecular weights of 18 for water and 142 for the sulfate, the crystals contain: (10)(18) = 634 lb/h or water of crystallization 142 Therefore, the total flow rate of crystals = 500 + 634 = 1,134 lb/h. This leaves 4,000 - 634 = 3,366 lb/h of water and 1,000 - 500 = 500 lb/h of sulfate in the mother liquor. The total flow rate of mother liquor is 3,366 + 500 = 3,866 lb/h. Thus, the concentration of sulfate in the mother liquor = 500/3,866 x 100% = 12.9 wt% Na2SO4. From the diagram below, the required temperature for this concentration at C in the mother liquor is 18oC. 500 Exercise 4.59 Subject: Dissolving crystals of Na2SO4 with water. Given: 500 kg of Na 2SO 4 10H 2 O crystals and 500 kg of Na2SO4 crystals at 20oC. Assumptions: Equilibrium according to phase diagram of Fig. 4.24. Find: Amount of water to dissolve the crystals at 20oC. Analysis: From Fig. 4.24, the solubility of Na2SO4 in water at 20oC is 15 wt% Na2SO4. Molecular weight of Na2SO4 = 142. Molecular weight of Na 2SO 4 10H 2 O = 322. Therefore, the kg of Na2SO4 in Na 2SO 4 10H 2 O = 500(142/322) = 220 kg. The water in Na 2SO 4 10H 2 O = 500 - 220 = 280 kg. Total Na2SO4 in the crystals = 500 + 220 = 720 kg. Therefore, need a total of 720(85/15) = 4,080 kg water. Additional water needed = 4,080 - 280 = 3,800 kg. Exercise 4.60 Subject: Adsorption of phenol (B) from an aqueous solution at 20oC with activated carbon. Given: One liter of aqueous solution containing 0.01 mol phenol. Freundlich isotherm equation for adsorption of phenol from aqueous solution by activated carbon at 20oC. Assumptions: Attainment of equilibrium. Find: Grams of activated carbon for (a) 75%, (b) 90%, and (c) 98% adsorption of phenol * 0.233 Analysis: At equilibrium, Eq. (1) in Example 4.12 is: q B = 2.16cB (1) * where, q B = mmol phenol adsorbed/g carbon cB = mmol phenol in solution/liter solution For each case, cB = 0.01(1,000 mmol/mol)(1-fraction adsorbed) mmol adsorbed = 0.01(1,000 mmol/mol)(1 liter)(fraction adsorbed) * g activated carbon needed = mmol adsorbed/ q B Using these equations, the results are as follows: * Case g activated cB at equilib, mmol phenol qB , adsorbed carbon mmol/L mmol/g (a) 75% adsorbed 2.5 7.5 2.67 2.81 (b) 90% adsorbed 1.0 9.0 2.16 4.17 (c) 98% adsorbed 0.2 9.8 1.48 6.62 Exercise 4.61 Subject: Adsorption of a colored substance (B) from an oil by clay particles at 25oC. Given: Oil with a color index of 200 units/100 kg oil. Adsorption equilibrium data Assumptions: One adsorption equilibrium contact. Find: (a) Freundlich equation for the adsorption equilibrium data. (b) kg clay to reduce color index to 20 units/100 kg oil for 500 kg of oil * 1/ Analysis: (a) From Eq. (4-30), the Freundlich equation is q B = AcB n , where here, * q B = color units/100 kg clay and cB = color units/100 kg oil By nonlinear regression of the data, A = 0.6733 and (1/n) = 0.5090 * The linearized form of the Freundlich equation is log q B = log A + (1 / n) log cB By linear regression of the data with the linearized form, A = 0.6853, (1/n) = 0.5050 The two results are close and both fit the data quite well. Use the nonlinear regression result: * 0.509 qB = 0.6733cB (1) (b) ( Q = 500 kg oil S = kg clay cBF ) = 200 units/kg oil Need, at equilibrium, cB = 20 units/kg oil * From Eq. (1), qB = 0.6733(20) 0.509 = 3.09 units/kg clay Apply Eq. (4-29 for a material balance on the color units: * q B = 3.09 = - Q Q 500 500 ( cB + cBF ) = - (20) + (200) S S S S (2) Solving Eq. (2), S = 29,100 kg of clay Exercise 4.62 Subject: Absorption of acetone (A) from air by water at 1 atm. Given: y-x data for the distribution of acetone between air and water. Assumptions: Only acetone undergoes mass transfer so that there is no water in the gas phase and no air in the liquid phase. Find: (a) Plots of the equilibrium data as (1) Y-X , (2) p-X' '(in mass units) , and (3) y-x. (b) Compositions of the exiting streams when 20 moles of gas containing 0.015 mole fraction of acetone is brought to equilibrium with 15 moles of water. Analysis: (a) Equilibrium data are given in acetone mole fractions (y and x). To convert to acetone mole ratios (Y and X), acetone partial pressure (p), and acetone mass ratio in the liquid (X'), the following equations are used, with molecular weight of acetone = 58: y x 58 ' YA = A , XA = A , pA = yA P , XA = XA 1 - yA 1 - xA 18 Using these four equations, the following values are obtained from the given data: yA xA YA, mol A/mol air XA, mol A/mol water 0.002004 0.00402 0.00604 0.00807 0.0101 0.01215 pA, atm 0.004 0.008 0.014 0.017 0.019 0.020 ' X A , g A/g water 0.00645 0.01294 0.01945 0.02599 0.0325 0.0391 0.004 0.002 0.00402 0.008 0.004 0.00807 0.014 0.006 0.0142 0.017 0.008 0.0173 0.019 0.010 0.0194 0.020 0.012 0.0204 Plots are given on the next page. (b) Solve by the intersection of the acetone material balance line with the equilibrium curve on the Y-X plot. Feed gas contains (0.015)(20) = 0.30 moles acetone and 20 - 0.3 = 19.7 moles air. The acetone material balance is: 0.30 = 19.7YA + 15 X A or by rearrangement , YA = -0.761 X A + 0.01523 (1) This equation is plotted on the Y-X plot on the next page and gives an intersection at YA = 0.0114 moles acetone/mole air and XA = 0.0051 moles acetone/mole water. The material balance is: Component: Air Water Acetone Gas in, moles 19.7 0.0 0.3 Liquid in, moles 0.0 15.0 0.0 Gas out, moles 19.7 0.0 0.224 Liquid out, moles 0.0 15.0 0.076 Analysis: (a) (continued) Exercise 4.62 (continued) Exercise 4.63 Subject: Separation of air into O2 and N2 by absorption into water followed by desorption. Given: Air containing 79 mol% N2 and 21 mol O2. Pressures ranging from 101.3 kPa (1 atm) to 10,130 kPa (100 atm), and temperatures ranging from 0 to 100oC. Assumptions: Applicability of Henry's law for solubility of H2 and O2 in water. Find: (a) A workable process. (b) Number of batch absorption steps to make 90 mol% pure O2 and the corresponding yield of O2. Analysis: Henry's law constants are given in Fig. 4.27 as (1/H) as function of temperature, where from Eq. (4-32), xi = (1/Hi)yiP with pressure in atm. Absorption is most efficient at high pressures, while desorption is best at low pressures. From Henry's law in this form, Ki = yi/xi = 1/[(1/Hi)P]. Using Fig. 4.27, the following are calculated: P = 100 atm, 25oC: P = 1 atm, 100oC: -1 -1 Ki Ki Component 1/Hi, atm N2-O2 1/Hi, atm N2-O2 -5 -6 Nitrogen 1.1 x 10 910 2.02 7.5 x 10 133,000 1.60 Oxygen 2.2 x 10-5 450 1.2 x 10-5 83,000 (a) A possible continuous scheme is the following, where air is compressed to 100 atm, followed by absorption with water at 25oC. Because the K-value for O2 is lower than that of N2 , the O2 is more easily absorbed. The bottoms liquid is heated to 100oC and flashed at 1 atm to release the dissolved O2. However, because of the large K-value for oxygen in the absorber, the process may require huge amounts of water to obtain a high yield of oxygen. Furthermore, the much higher concentration of nitrogen in the entering air and the fact that the relative volatility of nitrogen with respect to oxygen is not very high will make it very difficult to obtain a sharp split with absorption, as compared to distillation. Thus, the process may not be capable of either high yield or high purity of oxygen. The flash is capable of releasing any dissolved organic. Exercise 4.63 (continued) Analysis: (continued) (b) For batch absorption, the yield and purity of the oxygen, after it flashed off from the water, depends upon the amount of water used. Let the amount of air be 100 kmol (79 kmol N2 and 21 kmol O2). Consider amounts of water equal to 100, 1000, 10000, 50000, and 100000 kmol/h. Add this to the air and flash it at 100 atm and 25oC. Using the isothermal flash equations of Table 4.4 (with the exception of the energy balance, which is not needed) with the K-values of 910 for nitrogen and 450 for oxygen from the table above, the following results are obtained for a single equilibrium absorption, neglecting the mass transfer of water (stripping into the gas phase), i.e. assigning a K-value of 0 to water: Equilibrium conditions: Water, kmol kmol N2 in V kmol O2 in V kmol N2 in L kmol O2 in L Mol fraction O2 in flashed vapor 100 78.92 20.95 0.08 0.05 0.349 1,000 78.15 20.55 0.85 0.45 0.347 10,000 70.33 16.81 8.67 4.19 0.326 50,000 32.95 5.49 46.05 15.51 0.252 100,000 0.00 0.00 79.00 21.00 0.210 These results show that for small amounts of water very little absorption occurs, but the greatest increase in O2 purity occurs. If additional stages were used with small amounts of water to absorb the flashed vapor so as to reach 90 mol% purity of O2 ,the yield would be extremely small. At the highest liquid rate in the table above, all of the air is absorbed and, therefore, no separation between nitrogen and oxygen occurs. If 50,000 kmol of water is used, about 75% of the O2 is absorbed along with about 60% of the N2, with a resulting O2 mole fraction of 0.252 in the flashed vapor. If we start with 30,000 kmol of water and continue at this level using 300 kmol of water per kmol of flashed vapor for each batch absorption stage, the following results are obtained, using the isothermal flash equations for absorption at 100 atm and 25oC, followed by complete release of N2 and O2 at 1 atm and 100oC, but neglecting the vaporization of water: Equilibrium conditions: Stg. Water, kmol N2 kmol O2 kmol N2 kmol O2 y of O2 in flashed kmol in V in V in L in L vapor 1 30,000 52.07 10.27 26.93 10.73 0.285 2 11,300 17.54 5.16 9.38 5.58 0.373 3 4,490 6.02 2.62 3.36 2.96 0.468 4 1,895 2.12 1.35 1.24 1.60 0.563 6 411 0.287 0.383 0.189 0.510 0.729 8 113 0.045 0.120 0.033 0.179 0.844 10 37 0.008 0.042 0.006 0.068 0.913 For these conditions, it takes 10 batch stages of absorption and flashing to obtain a vapor of at least 90 mol% O2. However, the yield of O2 is very low, 0.068/21 x 100% = 0.32%. It is doubtful that any processing conditions exist that can deliver 90 mol% O2 with a high yield by absorption with water, followed by flash desorption. Exercise 4.64 Subject: Absorption of ammonia from nitrogen into water at 200C and 1 atm (760 torr). Given: 14 m3 of gas containing equimolar NH3 (A) and N2 (N)to be contacted in one equilibrium stage with 10 m3 of water (W). Equilibrium data for partial pressure of NH3 over water at 20oC. Assumptions: No mass transfer of water and nitrogen. Find: % absorption of NH3. Analysis: First compute the kmol of water. Density = 1000 kg/m3. M of water = 18.02. Therefore, we have (10)(1000)/18.02 = 555 kmol = L. Next, compute the kmol of gas. M of ammonia = 17.03. Assume the 14 m3 are at 1 atm and 20oC. At 1 atm and 0oC, have 22.4 m3/kmol. At 20oC and 1 atm, have 22.4(293/273) = 24.04 m3/kmol or 14/24.04 = 0.582 kmol. Because NH3 is the only component being transferred use mole ratios, with: X = moles NH3 / mole H2O and Y = moles NH3 / mole N2. Convert equilibrium data to these variables, using: Y = pA/(760 torr - pA) , where pA is in torr. X= grams NH 3 100 grams H 2 O 18.02 1 grams NH 3 = 0.0106 17.03 100 100 grams H 2 O Y 1.6207 0.6450 0.4259 0.2795 0.1765 0.1008 0.0704 0.0435 0.0339 0.0245 0.0201 0.0160 0.0000 X 0.4240 0.3180 0.2650 0.2120 0.1590 0.1060 0.0795 0.0530 0.0424 0.0318 0.0265 0.0212 0.0000 p of g NH3 / NH3 ,torr 100 g H2O 470 40 298 30 227 25 166 20 114 15 69.6 10 50.0 7.5 31.7 5 24.9 4 18.2 3 15 2.5 12 2 0 0 Exercise 4.64 (continued) Analysis: (continued) The kmol of N2 = 0.5(0.582) = 0.291 kmol = G. Material balance for NH3 , where 0 = inlet and 1 = outlet, with Y0 = 1.0 , and X0 = 0.0: (1) LX 0 + GY0 = G = 0.291 = LX 1 + GY1 = 555 X 1 + 0.291Y1 Rearranging Eq. (1), Y1 = 1 - 1907 X 1 (2) Because Y1 at the outlet must be less than the inlet value of 1, X1 must be a very small number. Assume that at small values of X, the equilibrium curve is linear. Using the lowest point at Y = 0.0160 and X = 0.0212, then: Y1 = (0.0160/0.0212)X1 = 0.755X1 (3) Solving Eqs. (2) and (3) simultaneously, Y1 = 3.96 x 10-4 mol NH3/mol N2 X1 = 5.25 x 10-4 mol NH3/mol H2O Therefore, in the final gas, have Y1G = (3.96 x 10-4)(0.291) = 0.000115 kmol NH3 Therefore, NH3 not absorbed = 0.000115/0.291 x 100% = 0.04%. Therefore, % NH3 absorbed = 100 - 0.04 = 99.96. Exercise 4.65 Subject: Desublimation of phthalic anhydride (PA) from a gas by cooling. Given: 8,000 lbmol/h of a gas containing 67 lbmol/h of PA cooled at 770 torr. s Assumptions: Only PA desublimes. Dalton's law. At equilibrium, pPA = PPA . Find: % recovery of PA for temperatures where the vapor pressure of PA, are (a) 0.7 torr, (b) 0.4 torr, and (c) 0.1 torr. Analysis: Let nPA = lbmol/h of PA still in the gas at equilibrium. Then by Dalton's law, s PPA nPA = [(8,000 - 67) + nPA ] (1) P s PPA (7,933) Solving for nPA , with P = 770 torr, nPA = 770 s (2) PPA 1- 770 and the percent recovery = (97 - nPA)/97 x 100% Solving Eqs. (2) and (3), Ps, torr nPA , lbmol/h % Recovery in gas of PA 0.7 7.2 89.3 0.4 4.1 93.9 0.1 1.03 98.5 (3) Exercise 4.66 Subject: Desublimation of anthraquinone (A) from nitrogen (N). Given: Gas containing 10 mol% A and 90 mol% N at 760 torr and 300oC is cooled to 200oC. Assumptions: Given vapor pressure data for solid A can be fitted to Antoine equation Find: % desublimation of A assuming no change in total pressure. Analysis: Using Polymath, the given vapor pressure data are fitted to the following Antoine eq.: PAs = 10 ss A 32 .8653- 49 , 762 .2 1327 .24 + T s A (1) where P is in torr and T is in C. From Eq. (1), P = 1.915 torr at 200oC. Assume 100 moles of initial gas, containing 90 moles N and 10 moles A. Let nA = moles of A still in the gas at equilibrium at 200oC. Then by Dalton's law, where pA = PAs , pA Ps 1.915 ntotal = A 90 + nA = 90 + nA P P 760 1.915 (90) nA = 760 = 0.227 mole Solving for nA , 1.915 1- 770 and the percent recovery = (10 - 0.227)/10 x 100% = 97.73% nA = (2) o (3) Exercise 4.67 Subject: Separation of a gas mixture of propylene (A) and propane (P) by adsorption. Given: T = 25oC, P = 101 kPa, gas containing 0.7 mole A and 1.3 mole P, and S = 0.1 kg silica gel. Equilibrium adsorption data in Fig. 4.30. Find: Equilibrium compositions and amounts of adsorbate and final gas. Analysis: Let y and x be the mole fractions of P in the vapor and adsorbate, respectively. Let G and W be the moles of equilibrium vapor and adsorbate, respectively. Then a propane material balance is: 1.3 = Wx + Gy . A total material balance is: 2.0 = W + G. Combining balances, y= 13 - Wx . 2 -W (1) To solve, assume W, plot the equation on Fig. 4.30a and find intersection with the equilibrium curve to find y and x. Read the mmole adsorbate/g adsorbent = moles adsorbate/kg adsorbent from Fig. 4.30b. Calculate kg adsorbent. If it is not 0.1 kg, repeat with a new guess of W. For W = 0.195 moles, G = 1.805 moles, y = 0.72 - 0.108x, y = 0.67, x = 0.41, S = 0.10 kg. Exercise 4.68 Subject: Separation of a gas mixture of propylene (A) and propane (P) by adsorption. Given: Gas containing 50 mol% A and 50 mol% P, and S = 1,000 lb of less of silica gel/lbmol of feed gas. Equilibrium adsorption data in Fig. 4.30 apply. Find: If one equilibrium stage can produce a vapor with 90 mol% P and an adsorbate of 75 mol% A. If not, what separation can be achieved. Analysis: With a basis of 1 lbmole of feed gas, let y and x be the mole fractions of P in the equilibrium vapor and adsorbate, respectively. Then we desire, y = 0.9 and x = 0.25. Let G and W be the lbmol of equilibrium vapor and adsorbate, respectively. Then a propane material balance is: 0.5 = Wx + Gy . A total material balance is: 1.0 = W + G. Combining balances, y= 0.5 - Wx 1-W (1) From the equilibrium curve in Fig. 4.30a, it is seen that when x = 0.25, y = 0.47, which is far from the desired value of 0.90. Therefore, cannot achieve the desired separation. So, calculate what can be achieved. To solve, assume W, plot the equation on Fig. 4.30a and find intersection with the equilibrium curve to find y and x. Read the mmole adsorbate/g adsorbent = lbmoles adsorbate/1000 lb adsorbent from Fig. 4.30b. However, we see in Fig. 4.30b that the smallest adsorbate amount is 1.57 lbmol adsorbate/1000 lb adsorbent. But we only have a total of 1.0 lbmol of feed gas. Therefore, we can use a lot less adsorbent. Assume a value of W = 0.5 lbmole (50% of the feed gas and equal to the amount of A in the feed gas). Then Eq. (1) becomes: y = 1 - x. From Fig. 4.30a, the intersection with the equilibrium curve is y = 0.635 and x = 0.365. From Fig. 4.30b, for x = 0.365, we have 2.0 lbmoles adsorbate/1000 lb of silica gel. Since we only have 0.5 lbmol of adsorbate, we only need 250 lb of silica gel. The separation achieved is far less than desired. Multiple stages might achieve the desired separation. Other values of W will give other separations. Exercise 4.69 Subject: Crystallization of MgSO4 from an aqueous by evaporation. Given: 1,000 lb of MgSO4 dissolved in 4,000 lb of water at 160oF. 90% of the water is evaporated at 160oF, causing the monohydrate to crystallize. At 160oF, MgSO4 has a solubililty of 36 wt%. Assumptions: Equilibrium. Find: (a) Pounds of water evaporated. (b) Pounds of monohydrate, MgSO 4 H 2 O , crystals produced. (c) Crystallizer pressure. Analysis: (a) Evaporate 0.9(4,000) = 3,600 lb of water Therefore, water left = 4,000 - 3,600 = 400 lb (b) MW of water = 18 and MW of MgSO4 = 120.4 Let W = lb of MgSO4 remaining in solution. Then MgSO4 in the crystals = 1,000 - W. From crystal stoichiometry, water of hydration in the crystals = 18 (1) (1,000 - W ) = 0.15(1,000 - W ) 120.4 Therefore, the water in solution = 400 - 0.15(1,000 - W) = 250 + 0.15W The total solution (water + dissolved sulfate) = 250 + 0.15W + W = 250 + 1.15W From solubility of the sulfate, W 0.36 = (2) 250 + 115W . Solving Eq. (2), W = 153.6 lb MgSO4 dissolved. MgSO4 in the crystals = 1,000 - 153.6 = 846.4 lb Water of crystallization = 18 (846.4) = 126.5 lb 120.4 Total crystals = 846.4 + 126.5 = 972.9 lb of MgSO 4 H 2 O (c) Since the crystallizer pressure depends only on temperature and the composition of the mother liquor in the crystallizer, it is the same as in Example 4.17 or 4.38 psia. Exercise 4.70 Subject: Crystallization of Na2SO4 from an aqueous solution by evaporation. Given: 1,000 kg/h of Na2SO4 dissolved in 4,000 kg/h of water. Water is evaporated at 60oC to crystallization 80% of the Na2SO4. Solubility data in Fig. 4.24. Assumptions: Equilibrium. Raoult's law for the mother liquor. Find: (a) The kg/h of water that must be evaporated. (b) The crystallizer pressure in torr. Analysis: (a) From Fig. 4.24, at 60oC, the stable crystals are Na2SO4 with no water of crystallization. The solubility of Na2SO4 at 60oC is 0.315 wt. Fraction. C = rate of crystal formation = 0.8(1,000) = 800 kg/h Therefore, sulfate dissolved in the mother liquor = 1,000 - 800 = 200 kg/h. From the solubility of the sulfate, water in the mother liquor = Therefore, we must evaporate, 4,000 - 435 = 3,565 kg/h (1 - 0.315) (200) = 435 kg/h 0.315 (b) MW of Na2SO4 = 142.05 Therefore, have 200/142.05 = 1.41 kmol/h of Na2SO4 in the mother liquor. Also, we have 435/18.02 = 24.14 kmol/h of water in the mother liquor. The mole fraction of water in the mother liquor = At 60oC, the vapor pressure of water = 2.89 paia. From Raoult's law, given by Eq. (2-42), the partial pressure of water = s xW PW = 0.945(2.89) = 2.73 psia Assuming that Na2SO4 has no vapor pressure in solution, crystallizer pressure = 2.73 psia. 24.14 = 0.945 24.14 + 1.41 Exercise 4.71 Subject: Bubble, secondary dew, and primary dew points of a three-component mixture assuming immiscibility of water and hydrocarbons. Given: (a) 50 mol% benzene (B), 50 mol% water (W). (b) 50 mol% toluene (T), 50 mol% water (W). (c) 40 mol% benzene (B), 40 mol% toluene (T), and 20 mol% water (W) Assumptions: Ideal hydrocarbon solutions and immiscibility with water. Find: Bubble, secondary dew, and primary dew point pressures at 50oC. Analysis: From the CHEMCAD program, vapor pressures at 50oC are: benzene 5.22 psia, toluene 1.78 psia, water 1.79 psia (a) At the bubble point, both pure liquid phases are present. Therefore, P = sum of the vapor pressures = 5.22 + 1.79 = 7.01 psia. This is also the secondary dew-point pressure. For a binary mixture, the primary dew point occurs when the pressure is increased to the point where the partial pressure of one component equals it vapor pressure. Since we have a 50-50 molar mixture, this occurs for the component with the lowest vapor pressure, water. Thus, water will condense first, with P = pW + pB = 1.79 + 1.79 = 3.58 psia. (b) Calculations for toluene and water are made in an identical manner to part (a). The results are: Bubble-point pressure = 1.79 + 1.78 = 3.57 psia. This is also the secondary dew point. The primary dew point is controlled by the toluene. Thus, the primary dew point pressure = 1.78 +1.78 = 3.56 psia. (c) Assuming Raoult's law from Eq. (2-42) to hold for the hydrocarbon phase, which is equimolar in B and T, and the water phase exerting its full vapor pressure, we have for the bubble point pressure from Eq. (4-34): s P = PW + xB PBs + xT PTs = 1.79 + 0.5(5.22) + 0.5(1.78) = 5.29 psia for the bubble point. For the primary dew point, assume first that the water condenses first. Then, using Dalton's law: p P s 179 . P= W = W = = 8.95 psia yW yW 0.2 But this higher than the bubble point pressure. So, instead, assume that the hydrocarbon phase condenses first. Apply the dew point equation, Eq. (4-13) to just the hydrocarbon phase: zi zi P xi = 1 = = (1) s HCs HCs Ki HCs P i Therefore, solving Eq. (1) for P, 1 1 P= = = 3.32 psia for the primary dew point. zi 0.4 0.4 + 5.22 1.78 HCs K i Analysis: (c) (continued) For this condition, the partial pressure of water is 0.2(3.32) = 0.664 psia (less than 1.79 psia). The secondary dew point is much more difficult to determine. We now know that as the pressure is increased, the hydrocarbons condense first. At the primary dew point, the partial pressure of W is 0.664 psia. W will not condense until its partial pressure increases to its vapor pressure of 1.79 psia. As the pressure is increased above the primary dew point pressure of 3.32 psia, the hydrocarbons will condense, reducing their mole fractions in the vapor phase and causing the mole fraction of water to increase in the vapor phase. This plus the increased total pressure will eventually cause the water partial pressure to reach its vapor pressure and begin condensing. We know the secondary dew point pressure must lie between the bubble point pressure of 5.29 psia and the primary dew point pressure of 3.32 psia. The equations that apply at the secondary dew point, where all pressures are in psia, are: Dalton's law: P = pB + pT + 1.79 = yBP + yTP + 1.79 Raoult's law: (2) Exercise 4.71 (continued) pB = xB PBs = xB (5.22) s T (3) pT = xT P = xT (178) . (4) Substituting Eqs. (3) and (4) into (2), P = 5.22xB +1.78xT + 1.79 (5) But in the hydrocarbon-rich liquid phase, xT = 1 - xB . Therefore, Eq. (5) becomes: (6) P = 5.22xB +1.78(1 - xB) + 1.79 = 3.44 xB + 3.57 Material balance equations for a basis of 100 moles, with 40 moles of B, 40 moles of T, and 20 moles of W: At the secondary dew point, only HCs have condensed. Therefore, the total moles of vapor is: V = 20P/1.79 = 11.17P (7) The moles of liquid is: L = 100 - V (8) A benzene material balance is: 40 = yBV + xBL (9) Substituting Eqs. (7) and (8) into (9), 40 = 11.17 yBP + 100xB -11.17xBP yB = 5.22xB/P (11) (10) We can also write from Eq. (3), Combining Eqs. (6), (10), and (11) to eliminate xB and yB , 3.247P2 -57.61P + 204.29 = 0 (12) Solving for the root between the bubble point and the primary dew point, the secondary dew point pressure = P = 4.90 psia. Exercise 4.72 Subject: Bubble and dew points of a mixture of benzene, toluene, and water. Given: Pressure of 2 atm. (a) 50 mol% benzene and 50 mol% water. (b) 50 mol% toluene and 50 mol% water. (c) 40 mol% benzene, 40 mol% toluene, and 20 mol% water. Assumptions: hydrocarbons and water immiscible. Find: For each mixture, bubble-point, primary dew-point, and secondary dew-point temperatures. Analysis: Use the CHEMCAD simulator with API-SRK for K-values, with the immiscible option and the three-phase flash model. The following results are obtained: Bubble point Secondary dew point Primary dew point T for Case (a), oC 89.4 89.4 100.0 T for Case (b), oC 104.3 104.3 110.3 T for Case (c), oC 93.9 98.0 99.7 Exercise 4.73 Subject: Bubble-point of a mixture of toluene, ethyl benzene, and water. Given: Mixture of 30 mol% toluene, 40 mol% ethylbenzene, and 30 mol% water at 0.5 atm. Assumptions: Hydrocarbons and water are immiscible. Find: Bubble-point temperature and vapor composition. Analysis: : Use the CHEMCAD simulator with API-SRK for K-values, with the immiscible option and the three-phase flash model. The following results are obtained: T = 70.8oC Vapor mole fractions are toluene 0.23, ethyl benzene 0.13, water 0.64 Exercise 4.74 Subject: Bubble, dew, and 50 mol% flash conditions for water-n butanol mixture. Given: Mixture of 60 mol% water (W) and 40 mol% n-butanol (B) at 101 kPa Find: (a) Dew-point temperature and liquid composition. (b) Bubble-point temperature and vapor composition. (c) Temperature and phase compositions for 50 mol% vaporization. Analysis: Use the CHEMCAD simulator with UNIFAC LLE for K-values, with the three-phase flash model. The following results are obtained: (a) 100.1oC with liquid mole fractions of 0.242 for W and 0.758 for B. (b) 93.5oC with vapor mole fractions of 0.773 for W and 0.227 for B. Note that the liquid consists of two phases, with the water-rich phase containing mole fractions of 0.969 for W and 0.031 for B. The less W-rich phase contains mole fractions of 0.564 for W and 0.436 for B. (c) 94.5oC with vapor mole fractions of 0.742 for W and 0.258 for B and a single liquid-phase mole fractions of 0.458 for W and 0.542 for B. The results deviate somewhat from Fig. 4.8. Exercise 4.75 Subject: Isothermal three-phase flash of 6-component mixture. Given: Feed mixture with composition given below. Find: Flow rates and compositions for 25oC and 300 kPa. Analysis: Use the CHEMCAD simulator with UNIFAC LLE for K-values, with the three-phase flash model. The following results are obtained: kmol/h: Component: Feed Vapor Liquid I Liquid II Hydrogen 350 349.90 0.07 0.03 Methanol 107 4.87 14.31 87.82 Water 491 3.32 3.04 484.64 Toluene 107 0.75 106.16 0.09 Ethyl benzene 141 0.37 140.59 0.04 Styrene 350 0.57 349.19 0.24 Total: 1,546 359.78 613.36 572.86 Exercise 5.1 Subject: Three-column interlinked cascade for a four-component system. Given: Configuration for a two-column interlinked cascade for a three-component system Find: Devise the three-column cascade. Analysis: The system is shown in the sketch below. Assume that the feed is comprised of components A, B, C, and D in the order of decreasing volatility. Starting from the feed end at the left-hand side, Column 1 performs a sloppy split to separate the feed into an overhead of A, part of B, and part of C; and a bottoms of part of B, part of C, and D. In the second column, the overhead is A and part of B, the middle sidestream is part of B and part of C, while the bottoms is D and part of C. In the third column, the four products are removed in nearly pure states. Exercise 5.2 Subject: Given: Batch cascades for liquid-liquid extraction Batch process in Fig. 5.19 Find: (a) Cascade diagram for Fig. 5.19. (b) Type of cascade in Part (a). (c) Comparison of process in Fig. 5.19 to a single stage process. (d) Modification to achieve a countercurrent cascade. Analysis: (a) A cascade diagram representing Fig. 5.19 is the following: Exercise 5.2 (continued) (b) The above cascade is a two-dimensional triangular cascade of the crosscurrent type. (c) For a single batch extraction using just the feed and one portion of the solvent, i.e. 100 ml of organic solvent and 100 ml of water containing 100 mg each of A and B, the resulting products are: Extract containing 100 ml solvent, 66.7 mg A, and 33.3 mg B. Raffinate containing 100 ml water, 33.3 mg A, and 66.7 mg B. If all 4 extracts and all 4 raffinates for the Fig. 5.19 are combined, we find: Extract containing 400 ml solvent, 66.7 mg A, and 33.3 mg B. Raffinate containing 400 ml water, 33.3 mg A, and 66.7 mg B. The process in Fig. 5.19 does produce extracts of varying purity, but overall, it is not effective. (d) The cascade above in Part (a) can be modified to give a more countercurrent process by reversing the flows of the solvents as shown in the following diagram. However, it is not possible to carry this out batchwise. Exercise 5.3 Subject: Two-stage membrane cascades for removal of nitrogen from methane by gas permeation. Given: Glassy polymer membrane that is selective for nitrogen. Assumptions: Desired degree of separation can not be achieved with one stage. Find: Two different two-stage cascades. Analysis: The two cascades are shown below. In both cases, the retentate is methanerich and the permeate is nitrogen-rich. In the first cascade, the permeate from Stage 2 is recycled to the feed to obtain a more pure methane retentate. In the second cascade, the retentate from Stage 2 is recycled to the feed to obtain a more pure nitrogen permeate. Exercise 5. 4 Subject: cascades. Leaching of oil from soybeans with benzene using countercurrent-flow Given: Soybean meal containing 2,500 kg/h of oil = L0 and 2,500 kg/h of insoluble solids. Benzene solvent as in Example 4.9. Assumptions: Equilibrium leaching stages where, for each stage, the weight fraction of oil in the overflow equals that in the underflow liquid. No solids in the overflows. Underflows contain 65 wt% solids. Find: Percent extraction of oil if: (a) Two stages are used with 5,000 kg/h of benzene = S. (b) Three stages are used with 5,000 kg/h of benzene = S. and (c) Number of stages to extract 98% of the oil with two times the min. benzene rate. Analysis: Let xi = mass fraction of oil in the underflow liquid from Stage i and yi = mass fraction of oil in the overflow liquid from Stage i. Let Li = underflow liquid leaving Since all Stage i and Vi = overflow liquid leaving Stage i. At any stage, yi = xi. solids entering a stage leave in the underflow from that stage and the total underflow is 65 wt% solids with 2,500 kg/h of solids, Li = (35/65)2,500 = 1,346 kg/h (a) The two-section cascade is shown below. Soybean meal enters Stage 1 at 5,000 kg/h and the benzene solvent enters Stage 2 at 5,000 kg/h. Underflow liquids are L1 = L2 = 1,346 kg/h. Total liquid material balance around Stage 2: S + L1 = V2 + L2 or 5,000 + 1,346 = V2 + 1,346. Therefore, V2 = 5,000 kg/h. Total liquid material balance around Stage 1: V2 + L0 = V1 + L1 or 5,000 + 2,500 = V1 + 1,346. Therefore, V1 = 6,154 kg/h. Oil material balance around Stage 1: L0 +V2y2 = L1x1+ V1y1 or 2,500 + 5,000y2 = 1,346x1 + 6,154y1 (1) Oil material balance around Stage 2: L1x1 = L2x2+ V2y2 or 1,346x1 = 1,346x2 + 5,000y2 (2) From equilibrium at each stage for the leaving liquids, y1 = x1 and y2 = x2 (3, 4) Solving the four linear equations, (1), (2), (3), and (4), we obtain: y1 = 0.3882 x1 = 0.3882 y2 = 0.08234 x2 = 0.08234 Analysis: (continued) Exercise 5. 4 (continued) Oil in the final extract = V1y1 = 6,154(0.3882) = 2,389 kg/h Percent extraction of oil = (2,389/2,500) x 100% = 95.6% This compares to 83.3% for one stage. (b) ) The three-section cascade is shown below. Soybean meal enters Stage 1 at 5,000 kg/h and the benzene solvent enters Stage 3 at 5,000 kg/h. Underflow liquids are L1 = L2 = L3 = 1,346 kg/h. Total liquid material balances around Stages 3, 2, and 1, respectively, similar to those in Part (a), give: V3 = V2 = 5,000 kg/h and V1 = 6,154 kg/h. Oil material balances around Stages 1, 2, and 3, respectively, similar to those in Part (a) give: 2,500 + 5,000y2 = 1,346x1 + 6,154y1 (5) 1,346x1 + 5,000y3 = 1,346x2 + 5,000y2 (6) 1,346x2 = 1,346x3 + 5,000y3 (7) From equilibrium at each stage for the leaving liquids, y2 = x2 y3 = x3 (8, 9, 10) y1 = x1 Solving the six linear equations, (5) through (10), we obtain: y1= x1 = 0.4015 y2 = x2 = 0.1023 y3 = x3 = 0.02169 Oil in the final extract = V1y1 = 6,154(0.4015) = 2,471 kg/h Percent extraction of oil = (2,471/2,500) x 100% = 98.8% This compares to 83.3% for one stage and 95.6% for two stages. (c) From Eq. (5-12), the minimum solvent rate corresponds to zero overflow, with all liquid going to the underflow. Liquid in the underflow = 1,346 kg/h. Can not apply the minimum solvent rate concept here because the liquid in the underflow is not stated as just solvent, but as total liquid including the oil. The total liquid in the underflow is less than the oil in the feed. However, note that for the given solvent rate, somewhat more than the desired 98% extraction is achieved with three stages. Exercise 5.5 Subject: Leaching of sodium carbonate from an insoluble oxide with water in a countercurrent-flow cascade. Given: From Example 5.1, entering flow rates are: Na2CO3 = FB = 1,350 kg/h, insoluble oxide = FA = 2,400 kg/h, and water = S = 4,000 kg/h. Underflow from each stage = 40 wt% on a solute (Na2CO3)-free basis. Assumptions: Equilibrium leaching stages where, for each stage, the weight fraction of carbonate in the overflow equals that in the underflow liquid. No solids in the overflows. Find: Minimum solvent water rate in lb/h. Ratio of solvent rate to minimum solvent rate. Percent recovery of carbonate with five stages and solvent rates from 1.5 to 7.5 times the minimum solvent rate. Analysis: R = mass of solvent/mass of insoluble solids = 40/60 = 0.667. Therefore solvent rate in the underflow = 2,400(0.667) = 1,600 kg/h. From Eq. (5-12), Smin = RFA = 0.667(2,400) = 1,600 kg/h or 3,630 lb/h Therefore, the ratio of solvent to minimum solvent = 2,400/1,600 = 2.5 To determine percent recoveries (extraction) of carbonate for 5 stages and various S/Smin , apply Eq. (5-10) to determine the mass ratio of carbonate to solvent in the final underflow leaving the F 1 last Stage N: (1) XN = B S W N -1 where, from Eq. (5-5), W =S/RFA = S/Smin (2) Combining Eqs. (1) and (2), noting that FB = 1,350 kg/h, FA = 2,400 kg/h, R = 0.667, and N = 5, XN = FB 1 FA R W N or X 5 = 1,350 (2,400)(0.667) 1 S S min 5 = 0.84375 S S min 5 (3) With a solvent rate in the underflow of 1,600 kg/h, the carbonate leaving in the underflow is 1,600 X5 . Therefore, the carbonate leaving in the final overflow (extract) = FB - 1,600 X5 = 1,350 - 1,600 X5 . 1,350 - 1,600 X 5 The percent extraction = 100% (4) 1,350 Analysis: (continued) Solving Eqs. (3) and (4), the following results are obtained: S/Smin 1.5 2.5 3.5 4.5 5.5 6.5 7.5 X5 0.1111 0.00864 0.001607 0.0004573 0.0001676 0.00007272 0.00003556 Exercise 5.5 (continued) % Recovery 86.83 98.98 99.81 99.95 99.98 99.99 99.996 Exercise 5.6 Subject: Recovery of Al2(SO4)3 from bauxite ore by reaction of Al2O3 with sulfuric acid followed by three-stage countercurrent washing with water, with the flowsheet below Given: 40,000 kg/day of bauxite ore containing 50 wt% Al2O3 and 50 wt% inert.Stoichiometric amount of 50 wt% aqueous H2SO4 to convert the Al2O3 to Al2(SO4) by the reaction: Al2O3 + 3 H2SO4 = Al2(SO4)3 + 3H2O (1) S = 240,000 kg/day of water to extract the sulfate in three countercurrent washing stages. Underflow from each washing stage contains 1 kg of water/kg of insoluble inert. Assumptions: Equilibrium washing stages where, for each stage, the weight fraction of sulfate in the overflow equals that in the underflow liquid. No solids in the overflows. Reaction of oxide with sulfuric acid is complete. Find: Flow rates in kg/day of sulfate, water, and inert solid in each product stream. Percent recovery of sulfate. Effect of adding one more stage. Analysis: In Eq. (1), the molecular weights from left to right are 101.96, 98.082, 342.158, and 18.016. Therefore, the yield of sulfate = 0.50(40,000)(342.158)/101.96 = 67,116 kg/day. Water leaving the reactor = water entering with the sulfuric acid + water from the reaction. Sulfuric acid entering the reactor = 0.50(40,000)(3)(98.082)/101.96 = 57,718 kg/day. Same amount of water, 57,718 kg/day, enters the reactor. Analysis: (continued) Exercise 5.6 (continued) Water from the reaction = 0.50(40,000)(3)(18.016)/101.96 = 10,602 kg/day. Total water leaving reactor and entering Stage 1 = 57,718 + 10,602 = 68,320 kg/day = L0 Inert solid leaving reactor = 0.5(40,000) = 20,000 kg/day. For the three-section washing cascade shown above, reactor effluent enters Stage 1 at a liquid rate of 68,320 + 67,116 = 135,436 kg/day and an inert solids rate of 20,000 kg/day. Water solvent enters Stage 3 at a rate, S, of 240,000 kg/day. Let L1 = L2 = L3 = water leaving in each underflow, which equals the inert solids rate or 20,000 kg/day. Let Vi = water rate in overflow from Stage i. Water material balances around Stages 2 and 3 give, respectively, V3 = V2 =240,000 kg/day. The water balance around Stage 1 is: L0 + V2 = L1 + V1 or 68,320 + 240,000 = 20,000 + V1 Solving, V1 = 288,320 kg/day Because the underflow liquid rates are constant in water, rather than water plus dissolved sulfate, it is preferable to work with mass ratios of sulfate to water, instead of sulfate mass fractions. Sulfate material balances around Stages 1, 2, and 3, respectively, in terms of mass ratios, Yi = kg dissolved Al2(SO4)3 / kg H2O in overflow from Stage i Xi = kg dissolved Al2(SO4)3 / kg H2O in underflow from Stage i (2) are: 67,116 + 240,000Y2 = 20,000X1 + 288,320Y1 20,000X1 + 240,000Y3 = 20,000X2 + 240,000Y2 (3) 20,000X2 = 20,000X3 + 240,000Y3 (4) From equilibrium at each stage for the leaving liquids, Y 2 = X2 Y 3 = X3 (5, 6, 7) Y 1 = X1 Solving the six linear equations, (2) through (7), we obtain: Y1= X1 = 0.2327 Y2 = X2 = 0.01927 Y3 = X3 = 0.001482 Sulfate in the final extract = V1Y1 = 288,320(0.2327) = 67,092 kg/day The overall material balance for flow rates in kg/day is: Component Bauxite Acid Reactor Solvent Final ore effluent Extract Al2O3 solid 20,000 Inert solid 20,000 20,000 H2SO4 57,718 Dissolved Al2(SO4)3 67,116 67,092 H2 O 57,718 68,320 240,000 288,320 Total: 40,000 115,436 155,436 240,000 355,412 Percent extraction of sulfate = (67,092/67,116) x 100% = 99.96% Waste 20,000 24 20,000 40,024 Analysis: (continued) Exercise 5.6 (continued) If an additional stage is added, a sulfate material balance for Stage 4 is needed and eight equations in eight unknowns are solved to give Y1= X1 = 0.23277. The percent extraction of sulfate now = (288,320)(0.23277)/67,116 x 100% = 99.996%. For the small additional recovery of sulfate (only 20 kg/day out of 67,116 kg/day), it is doubtful that an additional stage could be justified. Exercise 5.7 Subject: Given: Efficient rinsing of clothes in a washing machine Fixed amount of rinse water. Assumptions: Perfect mixing such that the concentration of soluble dirt in water adhering to the clothes (underflow) is equal to that in the discharged water (overflow). Find: (a) Whether the rinse water should be divided into two rinses. (b) Devise an efficient machine cycle. Analysis: Use the nomenclature of Section 5.2, letting: FA = mass of dry insoluble clothing FB = mass of soluble dirt. S = mass of rinse water Y = mass of soluble dirt/mass of overflow water X = mass of soluble dirt/mass of underflow water R = mass of water/mass of insoluble clothing in the underflow (a) For one rinsing stage, a material balance on soluble dirt gives, FB = XRFA + Y(S - RFA) (1) For equilibrium between overflow and underflow liquids, Y =X (2) Combining Eqs. (1) and (2), X = Y = FB/S (3) Mass of solids still left in liquid adhering to the clothes = XRFA = FBRFB/S Consider two crossflow rinsing stages, as shown in the diagram below, where the rinse water to each stage is 50% of the total rinse water of S/2. A soluble dirt material balance around Stage 1, with X1 = Y1 gives, FB = X1RFA + Y1(S/2 - RFA) = Y1S/2 (4) Solving, Y1 = X1 = 2FB/S (5) Analysis: (a) (continued) Exercise 5.7 (continued) The underflow sent to Stage 2 is X1RFA = 2FBRFA/S A similar material balance around Stage 2, with X2 = Y2 , gives, 2FBRFA/S = Y2(S/2 - RFA) + X2RFA = X2S/2 Solving, Y2 = X2 = 4FBRFA/S2 (6) The mass of solubles still left in the underflow from Stage 2 = X2RFA = 4FB(RFA)2/S2 The ratio of the mass of soluble dirt in the final underflow for two stages to that for one stage is, 2 4 FB RFA 4 RFA S2 = FB RFA S S It is highly likely that the rinsing will be such that the amount of rinse water used, S, will be much greater than the amount of water, RFA, that adheres to the clothing leaving in the underflow. Therefore, S >> 4 RFA and two stages of rinsing is much better than one stage. (b) A possible clothes washing cycle for efficient use of rinse water might be as follows: Put clothes and detergent in tub.. Fill tub with water. Agitate. Spin to small value of R. Fill tub with half of the rinse water. Agitate. Spin to small value of R. Fill tub with other half of the rinse water. Agitate. Spin to small value of R. Remove damp clothes and put into dryer. Exercise 5.8 Subject: Liquid-liquid extraction of acetic acid (B) from water (A) using chloroform (C). Given: Q = 10 liters of aqueous solution with 6 moles of acetic acid per liter. S = 10 liters of chloroform solvent. Extraction at 25oC with a distribution coefficient, cB C '' K DB = = 2.8 (1) cB A where concentrations are in moles/liter. Assumptions: Water is insoluble in chloroform. Chloroform is insoluble in water. Equilibrium is achieved in each stage. No change in volumes of feed and solvent as acid transfers from feed to the solvent. Find: Percent extraction of acid for: (a) One stage. (b) Three stages (crosscurrent) with one-third of solvent to each stage. (c) Three stages (crosscurrent) with 5 liters of solvent to Stage 1, 3 liters to Stage 2, and 2 liters to Stage 3. Analysis: To simplify the nomenclature, let: '' K = K DB y = cB C x = cB A Thus, from Eq. (1) y = Kx = 2.8x (a) For one stage, the acid material balance is, xFQ = y1S + x1Q or (6)(10) = 60 = y110 + x110 Combining Eqs. (2) and (3), y1 = 4.42 mol/L, x1 = 1.58 mol/L % extracted = y1S/ xFQ x 100% = (4.42)(10)/(6)(10) x 100% = 73.7% (b) For three crosscurrent stages, can use Eq. (5-21) in the following form, x3 1 = xF 1+ E / N 3 (2) (3) (4) where from Eq. (5-14), E = KS/Q = (2.8)(10)/10 = 2.8 and N = 3. Therefore, from Eq. (4), Percent extraction = 1 - x3 1 100% = 1 - 100% = 86.2% 3 xF (1 + 2.8 / 3) Exercise 5.8 (continued) (c) Use Eq. (5-13) one stage at a time, where from Eq. (5-14), Ei = KSi /Q = (2.8)Si /10 = 0.28 Si For Stage 1, S1 = 5 L, E1 = 1.4, and x1 = xF[1/(1+E1)] = 6[1/(1+1.4)] = 2.5 mol/L For Stage 2, S2 = 3 L, E2 = 0.84, and x2 = x1[1/(1+E2)] = 2.5[1/(1+0.84)] = 1.36 mol/L For Stage 3, S3 = 2 L, E3 = 0.56, and x3 = x2[1/(1+E3)] = 1.36[1/(1+0.56)] = 0.872 mol/L Percent extraction = (6 - 0.872)/6 x 100% = 85.5%, just slightly lower than in Part (b). Exercise 5.9 Subject: Liquid-liquid extraction of uranyl nitrate (UN) from water (W) with tributyl phosphate (TBP) Given: 100 grams of 20 wt% UN in water. Thus, FW = 80 g . Assumptions: Equilibrium batch contacts. W and TBP are mutually insoluble. Y ' . (1) Distribution coefficient is given in the form of Eq. (5-15): K DUN = UN = 55 X UN Find: Grams of TBP solvent needed to extract 90% of the UN by the following schemes: (a) One stage (b) Two crosscurrent stages with 50% of solvent to each stage. (c) Two countercurrent stages. (d) Infinite number of crosscurrent stages. (e) Infinite number of countercurrent stages. Analysis: Let: ' K = K DUN X (i ) = X UN in aqueous liquid leaving Stage i Y (i ) = YUN in TBP liquid leaving Stage i From Eq. (5-14), the extraction factor = E = KS/FW = 5.5S/80 = 0.0688S For 90% extraction, X(N)/X(F) = 1 - 0.90 = 0.10. (a) For a single stage, using Eq. (5-13), X (1) 1 1 = 010 = . = (F) X 1 + E 1 + 0.0688S Solving Eq. (2), S = 130.8 g of TBP. (b) For two crosscurrent stages with equal solvent additions, where S = total solvent, Eq. X (2) 1 1 (5-21) applies: = 010 = . = (3) 2 2 (F) X 1+ E / 2 1 + 0.0688S / 2 Solving Eq. (3), S = 62.9 g of TBP. (2) Analysis (continued): Exercise 5.9 (continued) (c) For two countercurrent stages, Eq. (5-28) applies: X (2) 1 1 = 010 = . = (F) 2 X 1+ E + E 1 + 0.0688S + 0.0688S Solving Eq. (4), S = 36.9 g of TBP. (d) For an infinite number of crosscurrent stages, Eq. (5-23) applies: (4) 2 X () 1 1 = 010 = . = (F) X exp E exp(0.0688S ) Solving Eq. (5), S = 33.5 g of TBP. (5) (e) For an infinite number of countercurrent stages, if E < 1, Eq. (5-32) applies: X ( ) / X ( F ) = 010 = 1 - E = 1 - 0.0688S . Solving, S = 13.1 g of TBP Exercise 5.10 Subject: Liquid-liquid extraction of uranyl nitrate (UN) from water (W) with tributyl phosphate (TBP) Given: 2 kg of 20 wt% UN in water. Thus, FW =1.6 kg. S = 0.5 kg of TBP. Assumptions: Equilibrium batch contacts. W and TBP are mutually insoluble. Y ' . (1) Distribution coefficient is given in the form of Eq. (5-15): K DUN = UN = 55 X UN Find: Percent recovery of UN for: (a) Two-stage batch extraction. (b) Three batch extractions. (c) Two-stage cocurrent extraction. (d) Three-stage countercurrent extraction. (e) Infinite-stage countercurrent extraction. (f) Infinite-stage crosscurrent extraction. Analysis: Let: ' K = K DUN X (i ) = X UN in aqueous liquid leaving Stage i Y (i ) = YUN in TBP liquid leaving Stage i From Eq. (5-14), the extraction factor = E = KS/FW = 5.5(0.5)/1.6 = 1.72 (a) For a single stage, using Eq. (5-13), X (1) 1 1 = = = 0.368 (F) X 1 + E 1 + 1.72 % extraction = (1 - 0.368) x 100% = 63.2% (b) Three batch extractions is equivalent to a crosscurrent arrangement with equal solvent additions, where S = total solvent. Eq. (5-21) applies: (2) X ( 3) 1 1 = = = 0.257 3 3 (F) X 1+ E / 3 1 + 172 / 3 . % extraction = (1 - 0.257) x 100% = 74.3% (3) (c) Two-stage cocurrent is same as one stage. Therefore, % extraction = 63.2%. Analysis (continued): Exercise 5.10 (continued) (d) For three-stage countercurrent extraction, Eq. (5-28) applies: X ( 3) 1 1 = = = 0.093 (F) 2 3 X 1+ E + E + E 1 + 172 + 1.72 2 + 172 3 . . % extraction = (1 - 0.093) x 100% = 90.7% (e) For infinite-stage countercurrent extraction, because E > 1, Eq. (5-31) applies. Therefore, X ( ) / X ( F ) = 0.0 and % extraction = 100% (f) For infinite-stage crosscurrent extraction, Eq. (5-23) applies: (4) X () 1 1 = = = 0179 . (F) X exp E exp(172) . % extraction = (1 - 0.179) x 100% = 82.1% (5) Exercise 5.11 Subject: Liquid-liquid extraction of dioxane (B) from water (A) with benzene (C). Given: 1,000 kg of 30 wt% B in A. Thus, FA = 700 kg . Assumptions: Equilibrium batch contacts. A and C are mutually insoluble. Distribution coefficient from Example 5.2, in the form of Eq. (5-15) is: Y ' K DB = B = 12 . (1) XB Find: S = kg of C needed to extract 95% of the dioxane by the following schemes: (a) One stage (b) Two crosscurrent stages with 50% of solvent to each stage. (c) Two countercurrent stages. (d) Infinite number of crosscurrent stages. (e) Infinite number of countercurrent stages. Analysis: Let: ' K = K DB X (i ) = X B in aqueous liquid leaving Stage i Y (i ) = YB in benzene liquid leaving Stage i From Eq. (5-14), the extraction factor = E = KS/FA = 1.2S/700 = 0.001714S For 95% extraction, X(N)/X(F) = 1 - 0.95 = 0.05. (a) For a single stage, using Eq. (5-13), X (1) 1 1 = 0.05 = = (F) X 1 + E 1 + 0.001714 S Solving Eq. (2), S = 11,090 kg of benzene. (b) For two crosscurrent stages with equal solvent additions, where S = total solvent, Eq. (5-21) applies: (2) X (2) 1 = 0.05 = (F) X 1+ E / 2 2 = 1 1 + 0.001714 S / 2 2 (3) Solving Eq. (3), S = 4,050 kg of benzene. Analysis (continued): Exercise 5.11 (continued) (c) For two countercurrent stages, Eq. (5-28) applies: X (2) 1 1 = 0.05 = = (F) 2 X 1+ E + E 1 + 0.001714 S + 0.001714 S Solving Eq. (4), S = 2,270 kg of benzene. (d) For an infinite number of crosscurrent stages, Eq. (5-23) applies: (4) 2 X () 1 1 = 0.05 = = (F) X exp E exp(0.001714 S ) Solving Eq. (5), S = 1,748 kg of benzene (5) (e) For an infinite number of countercurrent stages, if E < 1, Eq. (5-32) applies: X ( ) / X ( F ) = 0.05 = 1 - E = 1 - 0.001714 S Solving, S = 554 kg of benzene Exercise 5.12 Subject: Liquid-liquid extraction of benzoic acid (B) from water (A) using chloroform (C). Given: Q = 1,000 L/h of aqueous waste containing 0.05 mol/L of B. S = 500 L/h of chloroform solvent. Extraction with a distribution coefficient, cB C '' K DB = = 4.2 (1) cB A where concentrations are in mol/L. Assumptions: Water is insoluble in chloroform. Chloroform is insoluble in water. Equilibrium is achieved in each stage. No change in volumes of feed and solvent as acid transfers from feed to the solvent. Find: Percent extraction of acid for: (a) One stage. (b) Three crosscurrent stages with one-third of solvent to each stage. (c) Three countercurrent stages. Analysis: To simplify the nomenclature, let: '' K = K DB y = cB C x = cB A Thus, from Eq. (1) y = Kx = 4.2x (2) (a) For one stage, the acid material balance is, xFQ = y1S + x1Q or (0.05)(1,000) = 50 = y1500 + x11,000 Combining Eqs. (2) and (3), y1 = 0.0672 mol/L, x1 = 0.016 mol/L % extracted = y1S/ xFQ x 100% = (0.0672)(500)/(0.05)(1,000) x 100% = 67.2% x3 1 = xF 1+ E / N (3) (b) For three crosscurrent stages, can use Eq. (5-21) in the following form, 3 (4) where from Eq. (5-14), E = KS/Q = (4.2)(500)/1,000 = 2.1 and N = 3. Frm Eq. (4), Percent extraction = 1 - x3 1 100% = 1 - 100% = 79.6% 3 xF (1 + 2.1/ 3) (c) ) For three-stage countercurrent extraction, Eq. (5-28) applies in the form, x3 1 1 = = = 0.058 2 3 xF 1+ E + E + E 1 + 2.1 + 2.12 + 2.13 % extraction = (1 - 0.058) x 100% = 94.2% (5) Exercise 5.13 Subject: Given: Extraction of p-dioxane (B) from water (A) with solvent (C). Extraction factor, E = 0.9 compared to 2.4 in Example 5.2. Assumptions: Water and solvent are mutually insoluble. Equilibrium stages. Find: The effect on % extraction of the number and configuration of stages. Analysis: For a single stage and the cocurrent arrangement with any number of stages, N, Eq. (5-13) applies, where the B subscript has been dropped: X (N) 1 1 = = = 0.526 (F) X 1 + E 1 + 0.9 Therefore, the % extraction = (1 - 0.526) x 100% = 47.4% for any number of cocurrent stages. For a crosscurrent cascade, Eq. (5-21) applies: X (N) 1 = (F) X 1+ E / N N = 1 1 + 0.9 / N N (1) For a countercurrent cascade, Eq. (5-29) applies: X (N) = X (F) 1 N n=0 = n 1 N n=0 (2) n E 0.9 Results from Eqs. (1) and (2) are used to calculate % extraction from: X (N) 100% X (F) Using Eqs. (1) to (3), the following results are obtained: % extraction = 1 - N 2 3 4 5 (3) % extraction, crosscurrent 52.4 54.5 55.6 56.3 % extraction, countercurrent 63.1 70.9 75.6 78.7 Exercise 5.13 (continued) We see that for E < 1, the % extraction increases only slowly with increasing number of stages. It is desirable to have the situation where E > 1. Exercise 5.14 Subject: Multistage absorption of a hydrocarbon gas with a high-molecular-weight oil. Given: Hydrocarbon feed gas and absorbent oil of Example 5.3 in Fig. 5.11. Absorber pressure of 400 psia and average temperature of 97.5oF, with K-values at these conditions. Assumptions: Applicability of Kremser's method. Find: % absorption of components and total gas and % stripping of oil for N = 1, 3, 10, and 30 equilibrium stages. Effect of N on each component. Analysis: The same values of the absorption factors, A, and stripping factors, S, apply. Also Eqs. (5-48) and (5-50) are used to compute the fractions not absorbed, , and not stripped, B, respectively. Then Eq. (5-54) is used to compute the component flow rates in the exit vapor. Then an overall material balance is used to compute the component flow rates in the exit liquid: l N = l0 + N +1 - 1 Let the % absorption of a component in the entering gas = l N / l0 + N +1 100% The % stripping of the oil in the absorbent = 1 oil / l0 oil 100% Using a spreadsheet, the following results are obtained for the % absorption of C1, C2, C3, nC4, and nC5, the total % absorption of these light hydrocarbons, and % stripping of oil. N 1 3 6 10 30 C1 3.01 3.10 3.10 3.10 3.10 C2 11.19 12.58 12.60 12.50 12.60 C3 26.09 34.28 35.26 35.30 35.30 nC4 51.46 77.04 87.94 93.34 98.64 nC5 74.29 93.08 95.24 95.33 95.33 Total Gas 36.26 40.32 40.98 41.15 41.33 oil 0.05 0.05 0.05 0.05 0.05 Additional stages have little effect on the absorption of C1, C2, and C3 because their absorption factors are less than one. Exercise 5.14 (continued) Exercise 5.15 Subject: Multistage absorption of a hydrocarbon gas with a high-molecular-weight oil. Given: Hydrocarbon feed gas and absorbent oil of Example 5.3 in Fig. 5.11. Absorber pressure of 400 psia and average temperature of 97.5oF, with K-values at these conditions. Find: Material balance for absorbent rate of 330 lbmol/h instead of the 165 lbmol/h and 3 stages instead of the 6 in Example 5.3. Compare results with Example 5.3 and discuss effect of trading stages for absorbent flow rate. Analysis: With two times the absorbent rate of Example 5.3, double the absorption factors and halve the stripping factors given in Example 5.3. Then, using the same procedure as in that example, the following results are obtained and compared. Component C1 C2 C3 nC4 nC5 Oil Total N = 3 and L0 = 330 lbmol/h: N = 6 and L0 = 165 lbmol/h: 1 , lbmol/h l3 , lbmol/h 1 , lbmol/h l6 , lbmol/h 150.1 9.9 155.0 5.0 278.1 91.9 323.5 46.5 93.8 146.2 155.4 84.6 1.5 23.6 3.0 22.0 0.3 6.3 0.3 5.5 0.1 328.3 0.1 164.1 523.9 606.2 637.3 327.7 The results show that considerably more absorption takes place when the absorbent rate is doubled and the number of stages is halved. Exercise 5.16 Subject: Multistage absorption of a hydrocarbon gas with a high-molecular-weight oil. Given: Hydrocarbon feed gas and absorbent oil of Example 5.3 in Fig. 5.11. Absorber pressure of 400 psia and average temperature of 97.5oF, with K-values at these conditions. Find: Minimum absorbent rate to reduce the propane flow rate to 155.4 lbmol/h in exit vapor. Analysis: The minimum absorbent rate corresponds to N = . Use Eq. (5-48) to obtain the absorption factor, A = L/KV, needed for propane. Because A for propane < 1 for N = 6, it will be < 1 for N = , since the absorbent rate, L, will be lower. From Eq. (5-48), Fraction of propane not absorbed = 155.4/240 = 0.648 = = ( - 1)/(+1 -1) When A < 1, +1 = 0, therefore, Eq. (5-48) reduces to: A = 1 - = 1 - 0.648 = 0.352 From Eq. (5-38), L = AKV = 0.352(0.584)(800) = 165 lbmol/h = minimum rate This is the same L as in Example 5.3 because for > about 0.5, N above 5 has no effect on A. Exercise 5.17 Subject: Multistage absorption of a hydrocarbon gas with a high-molecular-weight oil. Given: Hydrocarbon feed gas and absorbent oil of Example 5.3 in Fig. 5.11. Absorber pressure of 400 psia and isothermal operation with N = 6 equilibrium stages at: (a) 125oF. (b) 150oF. Assumptions: Applicability of Kremser's method. Find: % absorption of components and % stripping of oil. Analysis: From Fig. 2.8, the following K-values are obtained, where for methane, the oil is assumed to have a MW of 300. The K-values for the absorbent oil are rough guesses. The values are compared to those in Example 5.3. Component C1 C2 C3 nC4 nC5 Oil K-values at 400 psia: T = 97.5F T = 125oF 6.65 8.0 1.64 2.0 0.584 0.73 0.195 0.26 0.0713 0.098 0.0001 0.0002 T = 150oF 8.8 2.4 0.90 0.34 0.135 0.0003 Absorption and stripping factors are obtained with Eqs. (5-38) and (5-51), respectively, using the above K-values with the entering vapor and liquid flow rates. Using the same procedure as in Example 5.3 and letting: The % absorption of a component in the entering gas = l N / l0 + N +1 100% The % stripping of the oil in the absorbent = 1 oil / l0 oil 100% The following results are obtained and compared with those of Example 5.3. % absorption for N = 6 stages at 400 psia: Component T = 97.5F T = 125oF T = 150oF C1 3.10 2.58 2.34 C2 12.60 10.31 8.59 C3 35.26 28.24 22.91 nC4 87.94 74.09 59.32 nC5 95.24 93.08 88.93 Oil 0.05 0.10 0.15 These results show that increasing temperature decreases the % absorption of gaseous components and increases the % stripping of the absorbent. Exercise 5.18 Subject: Multistage absorption of a light-hydrocarbon gas mixture with n-heptane. Given: VN+1 = 1,000,000 lbmol/day of feed gas containing in mole %, C1 94.9, C2 4.2, C3 0.7, nC4 0.1, nC5 0.1. Absorption at 550 psia and -30oF with N = 10 equilibrium stages and K-values below. Assumptions: Applicability of Kremser's method. Find: Absorbent flow rate, L0 , and component distribution for 50% absorption of C2. Analysis: From Eq. (5-48), A C2 = fraction of C 2 not absorbed = 0.50 = Solving and using Eq. (5-38), AC 2 = 0.50 = L0 L0 = KC 2VN +1 (0.36)(1,000,000) AC2 - 1 N AC2+1 - 1 Therefore, L0 = 0.50(0.36)(1,000,000) = 180,000 lbmol/day To compute the distribution of the other components,i, between products V1 and L10 , compute Ai = L0/KiVN+1 and A for i from Eq. (5-48). Then Eq. (5-54) is used to compute the component flow rates in the exit vapor. Then an overall material balance is used to compute the component flow rates in the exit liquid: l N = l0 + N +1 - 1 The results are as follows, using a spreadsheet: Flow rates, lbmol/day: Component K-value A l10 A n+1 = 11 1 + C1 2.85 0.063 0.937 949,000 889,200 59,800 C2 0.36 0.50 0.500 42,000 21,000 21,000 C3 0.066 2.73 0.0000276 7,000 0.2 6,999.8 nC4 0.017 10.59 0 1,000 0 1,000 nC5 0.004 45.0 0 1,000 0 1,000 Exercise 5.19 Subject: Stripping of a hydrocarbon liquid with steam at 50 psia and 300oF. Given: In Fig. 5.8(b), LN+1 = 1,000 kmol/h of 0.03% C1, 0.22% C2, 1.82% C3, 4.47% nC4, 8.59 nC5, and 84.87% nC10 , in mol%. V0 = 1,000 kmol/h of superheated steam. Stripper with N = 3 equilibrium stages. K-value for nC10 = 0.20. Other K-values from Fig. 2.8. Assumptions: Applicability of Kremser's equation. Negligible absorption of steam. Find: Compositions and flow rates of exiting stripped liquid and rich gas. Dew point. Analysis: For a stripper, Eq. (5-55) applies, which for no hydrocarbons in the entering steam simplifies to: l1 = lN+1 S (1) where, from Eq. (5-50), S = S -1 S -1 = 4 N +1 S -1 S -1 (4) (2) (3) where, from Eq. (5-53), By overall material balance, S = KV/L = KV0/LN+1 = K(1,000)/(1,000) = K N = lN+1 + 0 - l1 The results from applying the above equations with a spreadsheet are as follows: Flow rates, kmol/h: Component Ki = S S lN+1 = l in li = l out N = out C1 60 0 0.3 0.0 0.3 C2 28 0 2.2 0.0 2.2 C3 14 0 18.2 0.0 18.2 nC4 6.5 0.003 44.7 0.1 44.6 nC5 3.5 0.0168 85.9 1.4 84.5 nC10 0.20 0.8013 848.7 680.0 168.7 Steam 1,000.0 Total 1,000.0 681.5 1,318.5 Now check the primary dew point. Partial pressure of steam in the exit vapor = yi P = (1,000)/(1,318.5) x 50 = 37.9 psia. Vapor pressure of water at 300oF = 67 psia. Therefore, steam will not condense. Now check the possibility of the HCs condensing. yi 1 i No condensation if, for the HCs, from Eq. (4-13), = < 10 . V HCs Ki HCs Ki Using the above table, this sum is: 1 0.3 2.2 18.2 44.6 84.5 168.7 + + + + + = 0.0 + 0.0 + 0.001 + 0.005 + 0.018 + 0.640 = 0.664 1,318.5 60 28 14 6.5 3.5 0.2 Therefore, no condensation occurs. CHEMCAD computes a primary dew point of 287oF. Exercise 5.20 Subject: Given: Successive flash distillation as in Fig. 5.12 Two multiple flash arrangements, one with no recycle and one with recycle. Assumptions: Equilibrium at each flash stage. Find: Cooling requirements if heaters in Stages 2 and 3 are removed, and if so whether system will produce the same products. Analysis: Nothing is gained by totally condensing the vapor leaving each stage, followed by partial vaporization in the stage because equilibrium compositions depend only on the temperature and pressure. Therefore, remove a heater at a stage (2 or 3) and reduce the condenser duty on the vapor feed to that stage by the amount of the heater. Product compositions will be the same. In Fig. 12(a), new duty of condenser feeding stage 2 = 734 - 487 = 247 MBH. New duty of condenser feeding stage 3 = 483 - 376 = 107 MBH. Save two heat exchangers and reduce the size of the two condensers. Save total heating and cooling duties of 247 + 107 = 364 MBH. In Fig. 12(b), new duty of condenser feeding stage 2 = 883 - 665 = 218 MBH. New duty of condenser feeding stage 3 = 646 - 503 = 143 MBH. Save two heat exchangers and reduce the size of the two condensers. Save total heating and cooling duties of 218 + 143 = 361 MBH. Exercise 5.21 Subject: Effect of reflux rate on key-component distribution in multicomponent distillation by the group method. Given: Results in Example 5.4 for distillation at 400 psia of a superheated vapor feed at 105oF and containing, lbmol/h of 160 C1, 370 C2, 240 C3, 25 nC4, and 5 nC5 in a column with a partial condenser, 5 equilibrium rectification stages, equilibrium feed stage, 5 equilibrium stripping stages, and a partial reboiler, for an external reflux rate of 1,000 lbmol/h. Assumptions: No changes in stage temperatures when reflux rate is changed. Find: Molar ratio of C3 flow rate in the distillate to that in the bottoms, with a plot, for reflux rates of (a) 1,500, (b) 2,000, and (c) 2,500 lbmol/h. Analysis: Using the Edmister group method, the ratio of d to b for any component is given by a rearrangement of Eq. (5-66), d 1 = b AF From Eq. (5-65), From Eq. (5-59), From just below Eq. (5-64), S B AX + 1 / SX AC SE + 1 / AE (1) (2) (3) (4) AF = LF / K FVF AC = LC / KCVC = LC / KC D S B = K BVB / LB = K BVB / B where values of K are assumed to be the same as for Example 5.4 because of the assumption of no change in stage temperatures. Thus, AF , AC , and SB change only by changes in the reflux rate, which changes stagewise vapor and liquid rates as given below. From Eqs. (5-48) and (5-50), A -1 A N +1 - 1 S -1 S = N +1 S -1 A = (5) (6) where A = L/KV and S = KV/L. Analysis: (continued) In Eq. (1), Exercise 5.21 (continued) AX and SX pertain to the stripping (exhauster)section, AE and SE pertain to the rectifying (enricher) section, Because the temperatures change over the stripping and rectifying sections, average absorption and stripping factors are used. For all three cases, the new flow rates needed in the absorption and stripping factors are determined from the assumption of constant molar overflow, with the following results: Flow rates, lbmol/h: L0 = 1,000 (a) L0 = 1,500 Vapor Liquid Vapor Liquid 530 1,000 530 1,500 1,530 1,000 2,030 1,500 1,530 1,000 2,030 1,500 730 1,000 1,230 1,500 730 270 1,230 270 (b) L0 = 2,000 (c) L0 = 2,500 Vapor Liquid Vapor Liquid 530 2,000 530 2,500 2,530 2,000 3,030 2,500 2,530 2,000 3,030 2,500 1,730 2,000 2,230 2,500 1,730 270 2,230 270 Section Condenser Enricher Feed Exhauster Reboiler Assume that as with Example 5.4 and at higher reflux rates, components C1, C2, and nC5 will not distribute to any extent. Therefore, make distribution calculations only for the component of interest, propane. Using the above flow rates for Cases (a), (b), and (c) , the absorption and stripping factors for C3 with L0 = 1,000 lbmol/h become as follows, where, for example, AC for Case (a) is given from Eq. (3) by: AC Case (a) = L0 = DKC L0 D L0 AC D Example 5.4 = 1,500 = 15.72 1,000 530 10.48(530) (c) L0 = 2,500 26.20 2.17 0.461 1.397 1.416 0.706 8.49 Factor AC AE SE = 1/AE AF AX SX = 1/AX SB L0 = 1,000 10.48 1.72 0.581 1.108 1.73 0.577 2.78 (a) L0 = 1,500 15.72 1.94 0.515 1.250 1.54 0.649 4.68 (b) L0 = 2,000 20.96 2.08 0.481 1.340 1.46 0.685 6.59 Analysis: (continued) Exercise 5.21 (continued) Using Eqs. (5) and (6), the following factors are determined for the enricher and exhauster from the above values of AE , SE , AX , and SX : Factor AE AX SE SX L0 = 1,000 0.0289 0.028 0.435 0.439 (a) L0 = 1,500 0.0180 0.0438 0.494 0.379 (b) L0 = 2,000 0.0135 0.0530 0.526 0.351 (c) L0 = 2,500 0.0113 0.0589 0.544 0.336 Using Eq. (1) with the above factors, the following d/b ratios for C3 are obtained: Case Example5.4 (a) (b) (c) d/b for C3 0.0114 0.00522 0.00322 0.00237 As the reflux rate increases, the separation is enhanced. Exercise 5.22 Subject: Effect of numbers of rectification and stripping equilibrium stages on keycomponent distribution in multicomponent distillation by the group method. Given: Results in Example 5.4 for distillation at 400 psia of a superheated vapor feed at 105oF and containing, lbmol/h of 160 C1, 370 C2, 240 C3, 25 nC4, and 5 nC5 in a column with a partial condenser, 5 equilibrium rectification stages, equilibrium feed stage, 5 equilibrium stripping stages, and a partial reboiler, for an external reflux rate of 1,000 lbmol/h. Assumptions: No changes in absorption and stripping factors when numbers of equilibrium stages are changed. Find: Molar ratio of C3 flow rate in the distillate to that in the bottoms, with a plot, for (a) 10 rectification and 10 stripping stages, and (b) 15 rectification and 15 stripping stages. Analysis: Using the Edmister group method, the ratio of d to b for any component is given by a rearrangement of Eq. (5-66), d 1 = b AF From Eqs. (5-48) and (5-50), S B AX + 1 / SX AC SE + 1 / AE (2) (3) (1) A -1 A N +1 - 1 S -1 S = N +1 S -1 A = where A = L/KV and S = KV/L. In Eq. (1), AX and SX pertain to the stripping (exhauster)section, AE and SE pertain to the rectifying (enricher) section. Assume that all absorption and stripping factors are unchanged from the values given for Example 5.4. Recalculate the factors from Eqs. (2) and (3) for Case (a), where N = 10 for each section and Case (b), where N = 15 for each section. Analysis: (continued) Exercise 5.22 (continued) N =5 0.0289 0.028 0.435 0.439 N = 10 0.00185 0.00176 0.420 0.424 N = 15 0.000123 0.000113 0.419 0.423 Factor AE AX SE SX Using Eq. (1) with the above phi factors and the absorption and stripping factors from Example 5.4, as given below, the following d/b ratios for C3 are obtained: Factor Example 5.4 AC 10.48 AE 1.72 SE = 1/AE 0.581 AF 1.108 AX 1.73 SX = 1/AX 0.577 SB 2.78 Case d/b for C3 Example 5.4 0.0114 (a) 0.000733 (b) 0.0000487 As the number of stages increases, the separation is greatly enhanced. Exercise 5.23 Subject: Light hydrocarbon distillation by the Edmister group method. Given: Column feed at 225oF and 250 psia, corresponding to approximately 23 mol% vapor, with component flow rates below. Distillate rate is 230 kmol/h to give a separation between C3 and nC4. Column operates at 250 psia with a partial condenser and a reflux ratio of 5.0. Column contains 14 theoretical stages, with feed to stage 8, counting up from the bottom. Find: Compositions of the distillate and bottoms. Analysis: Using the DISTL (Edmister method) model of ASPEN PLUS with the ChaoSeader method for K-values, the results, with a computed 25.4 mol% vaporization, are: kmol/h: Component Feed Distillate Bottoms C2 30 29.999 0.001 C3 200 192.663 7.337 nC4 370 7.325 362.674 nC5 350 0.012 349.988 nC6 50 0 50.000 Total 1,000 229.999 770.000 Exercise 5.24 Subject: Light hydrocarbon distillation by the Edmister group method. Given: Bubble-point column feed at 120 psia, with component flow rates below. Distillate rate is 45 kmol/h to give a separation between nC4 and iC5. Column operates at 120 psia with a partial condenser and a boilup of 200 kmol/h Column contains 20 theoretical stages, with feed to stage 10, counting up from the bottom. Find: Compositions of the distillate and bottoms. Analysis: Using the DISTL (Edmister method) model of ASPEN PLUS with the ChaoSeader method for K-values, and a reflux rate of 200 - 45 = 155 kmol/h (reflux ratio = 155/45), gives: kmol/h: Component Feed Distillate Bottoms C3 5 4.996 0.004 iC4 15 14.987 0.013 nC4 25 24.587 0.413 iC5 20 0.354 19.646 nC5 35 0.077 34.923 Total 100 45.001 54.999 Exercise 5.25 Subject: Degrees of freedom analyses for a partial reboiler and a total condenser. Given: A partial reboiler with exiting vapor and liquid in equilibrium. A total condenser. Find: Number of variables, number of independent equations, and degrees of freedom. Analysis: Partial reboiler: Have 3 streams with C + 3 variables each, plus the heat transfer rate. Therefore, NV = 3(C + 3) + 1 = 3C + 10 The independent relationships are: C Component material balances 1 Enthalpy balance 1 Equality of temperatures of exiting vapor and liquid 1 Equality of pressures of exiting vapor and liquid C Component phase equilibrium between exiting vapor and liquid 3 Mole fraction summations for 3 streams Therefore, NE = 2C + 6 From Eq. (5-67), ND = NV - NE = (3C + 10) - (2C + 6) = C + 4 Total condenser: Have 2 streams with C + 3 variables each, plus the heat transfer rate. Therefore, NV = 2(C + 3) + 1 = 2C + 7 The independent relationships are: C Component material balances 1 Enthalpy balance 2 Mole fraction summations for 2 streams Therefore, NE = C + 3 From Eq. (5-67), ND = NV - NE = (2C + 7) - (C + 3) = C + 4 Exercise 5.26 Subject: Given: Degrees of freedom analyses for a stream mixer and a stream divider. A steam mixer and a stream divider, each with heat transfer. Find: Number of variables, number of independent equations, and degrees of freedom. Analysis: Stream mixer with two inlet streams: Have 3 streams with C + 3 variables each, plus the heat transfer rate. Therefore, NV = 3(C + 3) + 1 = 3C + 10 The independent relationships are: C Component material balances 1 Enthalpy balance 3 Mole fraction summations for 3 streams Therefore, NE =C + 4 From Eq. (5-67), ND = NV - NE = (3C + 10) - (C + 4) = 2C + 6 Stream divider with two exiting streams: Have 3 streams with C + 3 variables each, plus the heat transfer rate. Therefore, NV = 3(C + 3) + 1 = 3C + 10 The independent relationships are: 1 Total material balance Component mole fraction equalities for exiting streams relative to the feed 2(C - 1) 1 Enthalpy balance 1 Equality of temperatures of exiting streams 1 Equality of pressures of exiting streams 3 Mole fraction summations for 3 streams Therefore, NE = 2C + 5 From Eq. (5-67), ND = NV - NE = (3C + 10) - (2C + 5) = C + 5 Exercise 5.27 Subject: Degrees of freedom analysis for the specifications of a distillation column. Given: Feed of 10 mol% benzoic acid and 90 mol% maleic anhydride. Column has a total condenser and a partial reboiler. Operation is at 100 torr with a reflux ratio of 1.2 times minimum. Desired products are a distillate of 99.5 mol% maleic anhydride and a bottoms of 0.5 mol% anhydride. Find: Whether problem is completely specified. Analysis: From operation (b) in Table 5.4, ND = 2N + C + 9 = 2N + 11 The given specifications are equivalent to: Pressure at each stage, including partial reboiler Pressure at reflux divider outlet Pressure at total condenser outlet Heat transfer rate for each stage, except reboiler Heat transfer rate for divider Feed mole fractions (only count C - 1) Reflux rate Mole fraction of component in distillate Mole fraction of component in bottoms Total Need 5 more specifications, which might be: Feed rate Feed temperature Feed pressure Optimal feed location Saturated liquid leaving condenser N 1 1 N-1 1 1 1 1 1 ______ 2N + 6 Exercise 5.28 Subject: Verification of numbers of degrees of freedom for 3 operations, and effect of adding another feed. Given: Degrees of freedom for operations (b), (c), and (g) in Table 5.4. Find: Verify values of ND and determine effect of adding another feed. Analysis: Operation (b), distillation with total condenser and partial reboiler: From Eq. (5-70), ND NR = 9 and NA = 0 Therefore, Eq. (1) becomes, N D unit = unit = ND all elements e - NR C + 2 + N A (1) ND all elements e -9 C+2 (2) Let N1 = number of rectifying stages and N2 = number of stripping stages. Using the values of ND for the elements in Table 5.3: C+4 Total condenser Reflux divider C+5 Rectifying section 2N 1 + 2C + 5 Feed stage with heat transfer 3C + 8 Stripping section 2N 2 + 2C + 5 C+4 Partial reboiler N D e = 2(N1 + N2) + 10C + 31 all elements (3) Substituting Eq. (3) into (2), N D unit = 2 N 1 + N 2 + 10C + 31 - 9 C + 2 = 2 N1 + N 2 + C + 13 (4) However, in Operation (b) of Table 5.4, the feed stage and the partial reboiler are counted in N. Therefore, N1 + N2 = N - 1 - 1 = N - 2. Substituting this into Eq. (4), gives ( N D )unit = 2( N - 2) + C + 13 = 2 N + C - 9 which verifies the result given in Table 5.4. (5) Exercise 5.28 (continued) If a second feed is added, we add C + 2 degrees of freedom for the feed variables and for feed location. Then, ( N D )2 feeds = ( N D )1 feed + C + 3 = 2 N + 2C + 12 Operation (c), distillation with a partial condenser and partial reboiler: With a partial condenser instead of a total condenser, there are only 8 redundant streams and no reflux divider. Therefore, Eq. (4) becomes: ( N D )unit = 2 N + C + 9 + (C + 2) - (C + 5) = 2 N + C + 6 which verifies the result given in Table 5.4. With two feeds, (6) ( N D )2 feeds = ( N D )1 feed + C + 3 = 2 N + 2C + 9 Operation (g), distillation with a partial reboiler and a mixed condenser that produces both vapor distillate and liquid distillate: With a mixed condenser, we can modify Operation (c), by adding a reflux divider and one redundant stream. Therefore, Eq. (6) becomes: ( N D )unit = 2 N + C + 6 + (C + 5) - (C + 2) = 2 N + C + 9 which verifies the result given in Table 5.4. With two feeds, ( N D ) 2 feeds = ( N D )1 feed + C + 3 = 2 N + 2C + 12 Exercise 5.29 Subject: Verification of numbers of degrees of freedom for 2 operations, and effect of adding a vapor sidestream.. Given: Degrees of freedom for operations (e) and (f) in Table 5.4. Find: Verify values of ND and determine effect of adding a vapor stream Analysis: Operation (e), reboiled absorber: From Eq. (5-70), ND NR = 6 and NA = 0 Therefore, Eq. (1) becomes, unit = ND all elements e - NR C + 2 + N A (1) ND unit = ND all elements e - 6(C + 2) (2) Let N1 = number of rectifying stages and N2 = number of stripping stages. Using the values of ND for the elements in Table 5.3: Rectifying section 2N 1 + 2C + 5 Feed stage with heat transfer 3C + 8 Stripping section 2N 2 + 2C + 5 Partial reboiler C+4 N D e = 2(N1 + N2) + 8C + 22 all elements (3) Substituting Eq. (3) into (2), N D unit = 2 N 1 + N 2 + 8C + 22 - 6 C + 2 = 2 N1 + N 2 + 2C + 10 (4) However, in Operation (e) of Table 5.4, the feed stage and the partial reboiler are counted in N. Therefore, N1 + N2 = N - 1 - 1 = N - 2. Substituting this into Eq. (4), gives ( N D )unit = 2( N - 2) + 2C + 10 = 2 N + 2C + 6 which verifies the result given in Table 5.4. (5 Exercise 5.29 (continued) Analysis: (continued) If a vapor sidestream is added, we add 2 degrees of freedom, one for the location and one for the sidestream flow rate. Then, ( N D ) unit with sidestream = 2 N + 2C + 6 + (2) = 2 N + 2C + 8 Operation (f), reboiled stripper: Using the values of ND for the elements in Table 5.3: Section of N - 1 stages Partial reboiler 2(N - 1) + 2C + 5 C+4 N D e = 2N +3C + 7 (6) all elements Substituting Eq. (3) into (2), with NR = 2 and NA = 0, ( N D )unit = 2 N + 3C + 7 - 2 ( C + 2 ) = 2 N + C + 3 If a vapor sidestream is added, we add 2 degrees of freedom, then, (7) ( N D )unit with sidestream = 2 N + C + 3 + (2) = 2 N + C + 5 Exercise 5.30 Subject: Verification of numbers of degrees of freedom for one operation, and effect of adding a liquid sidestream.. Given: Degrees of freedom for operation (h) in Table 5.4. Find: Verify value of ND and determine effect of adding a liquid sidestream Analysis: Operation (h), extractive distillation: From Eq. (5-70), N D unit = ND e - NR C + 2 + N A all elements (1) NR = 13 and NA = 0 Therefore, Eq. (1) becomes, N D unit = ND all elements e - 13(C + 2) (2) Let N1 = number of rectifying stages above the MSA feed, N2 = number of intermediate stages between the MSA feed and the feed, N3 = number of stripping stages. Using the values of ND for the elements in Table 5.3: Total condenser C+4 Reflux divider C+5 Rectifying section 2N 1 + 2C + 5 MSA Feed stage with heat transfer 3C + 8 Intermediate section 2N 2 + 2C + 5 Feed stage with heat transfer 3C + 8 Stripping section 2N 2 + 2C + 5 C+4 Partial reboiler N D e = 2(N1 + N2 + N3 ) + 15C + 44 (3) all elements Substituting Eq. (3) into (2), N D unit = 2 N 1 + N 2 + N 3 + 15C + 44 - 13 C + 2 = 2 N 1 + N 2 + N 3 + 2C + 18 (4) However, in Operation (h) of Table 5.4, the two feed stages and the partial reboiler are counted in N. Therefore, N1 + N2 + N3 = N - 3. Substituting this into Eq. (4), gives (5) ( N D )unit = 2( N - 3) + 2C + 18 = 2 N + 2C + 12 which verifies the result in Table 5.4. If a liquid sidestream is added, we add 2 degrees of freedom, then, ( N D )unit with sidestream = 2 N + 2C + 12 + (2) = 2 N + 2C + 14 Exercise 5.31 Subject: Alternative design variables for distillation Given: Distillation design variables in Table 5.4 Find: Conditions under which the alternative variables below could be specified. Analysis: (a) Condenser heat duty might be substituted for the reflux ratio if the maximum heat duty of the condenser in an existing tower were known. (b) Stage temperature might be substituted for the recovery of a light or heavy component if there was a limit on the temperature of the cooling or heating medium in the condenser or reboiler. (c) Intermediate stage vapor rate might be substituted for reflux ratio for an existing tower diameter with a limiting vapor velocity to prevent flooding. (d) Reboiler heat load might be substituted for reflux ratio in the case of an existing reboiler area to prevent transition to film boiling. Exercise 5.32 Subject: Difference in degrees of freedom between a distillation column with a total condenser and one with a partial condenser. Given: A difference of 3 provided that only a vapor distillate leaves the partial condenser. Find: The difference in degrees of freedom Analysis: From Eq. (5-70), N D unit = ND all elements e - NR C + 2 + N A (1) For a total condenser plus reflux divider, NR = 1 and NA = 0 Therefore, Eq. (1) becomes, N D unit = N D e - (C + 2) all elements (2) Using the values of ND for the elements in Table 5.3: Total condenser C+4 Reflux divider C+5 Therefore, N D unit = (C + 4) + (C + 5) - (C + 2) = C + 7 For a partial condenser, NR = 0 and NA = 0, and using Table 5.3, N D unit = (C + 4) Therefore, the difference = (C + 7) - (C + 4) = 3 degrees of freedom Exercise 5.33 Subject: Replacement of a partial reboiler with live steam injection. Given: Operation (b) in Table 5.4 for distillation with a total condenser and partial reboiler. A column without the reboiler, but with steam injection and design variables listed below. Assumptions: Injection of steam will not cause the formation of a second liquid phase. Find: (a) The number of degrees of freedom = number of design variables. (b) The additional design specifications needed. Analysis: (a) Table 5.4, Operation (b), ND = 2N + C + 9. If the partial reboiler is eliminated, then from Eq. (5-70), N D unit = ND e - NR C + 2 + N A all elements (1) NR = 7 and NA = 0 Therefore, Eq. (1) becomes, N D unit = ND all elements e - 7 ( C + 2) (2) Let N1 = number of rectifying stages N2 = number of stripping stages Using the values of ND for the elements in Table 5.1: C+4 Total condenser Reflux divider C+5 Rectifying section 2N 1 + 2C + 5 Feed stage with heat transfer 3C + 8 Stripping section 2N 2 + 2C + 5 N D e = 2(N1 + N2 ) + 9C + 27 all elements (3) Substituting Eq. (3) into (2), N D unit = 2 N 1 + N 2 + 9C + 27 - 7 C + 2 = 2 N 1 + N 2 + 2C + 13 (4) However, if the feed stage is counted in the total N stages, then N1 + N2 = N - 1. Substituting this into Eq. (4), gives (5) ( N D )unit = 2( N - 1) + 2C + 13 = 2 N + 2C + 11 (b) The specifications are: Pressures at all stages + condenser + divider = N + 3 Heat transfer rate for all stages + divider = N + 1 Feed conditions = C + 2 Alcohol concentration in distillate = 1 Stripping steam composition = C - 1 This totals 2N + 2C + 6. Therefore, 5 additional specifications are needed. Might also specify: Temperature and pressure of stripping steam, saturated liquid reflux, alcohol concentration in bottoms, and optimal feed-stage location. Exercise 5.34 Subject: Degrees of freedom and specs for distillation column with a boilup divider. Given: Column flow diagram in Fig. 5.26 and specifications listed below. Find: (a) Number of degrees of freedom. (b) The reason why a feed rate cannot be specified. Analysis: (a) From Eq. (5-70), N D unit = NR =10 and NA = 0 Therefore, Eq. (1) becomes, N D unit = ND all elements e - NR C + 2 + N A (1) ND all elements e - 10 C + 2 (2) Let N1 = number of rectifying stages and N2 = number of stripping stages. Using the values of ND for the elements in Table 5.3: Total condenser C+4 C+5 Reflux divider Rectifying section 2N 1 + 2C + 5 Feed stage with heat transfer 3C + 8 Stripping section 2N 2 + 2C + 5 Boilup divider C+5 C+4 Total reboiler N D e = 2(N1 + N2) + 11C + 36 (3) all elements Substituting Eq. (3) into (2), N D unit = 2 N 1 + N 2 + 11C + 36 - 10 C + 2 = 2 N1 + N 2 + C + 16 (4) However, the feed stage is counted in N. Therefore, N1 + N2 = N - 1. Substituting this into Eq. (4), gives (5) ( N D )unit = 2( N - 1) + C + 16 = 2 N + C + 14 (b) The specified variables are: Number of stages and feed stage 2 Pressures of all stages, condenser, reboiler, and 2 dividers N + 4 Adiabatic for all stages and 2 dividers N+2 Feed composition , temperature, and pressure C+1 2 Reflux rate set by L/V = 1.2 and vapor rate by velocity Condensate temperature set by water temperature Concentrations of A in distillate and C in bottoms Total 1 2 2N + C + 14 Thus, all of degrees of freedom are utilized and the column feed rate can not be specified. Exercise 5.35 Subject: Degrees of freedom and specifications for a mixed-feed, triple-effect evaporation system to concentrate an aqueous solution of an organic acid. Given: Flow diagram of system in Fig. 5.27, and design specifications listed below Assumptions: Organic acid is non-volatile. No heat losses from evaporator effects. Find: Number of degrees of freedom. Any additional specifications needed. Analysis: The system in Fig. 5.27 consists of 3 evaporator effects, a total condenser, and a pump. Number of interconnecting streams = NR = 6, and NA = 0. First determine the number of degrees of freedom for one evaporator effect. Let: V = vapor mass rate leaving an effect. L = aqueous solution mass rate leaving an effect. F = aqueous solution mass feed rate to an effect. S (= D) = heating steam mass rate to and from an effect. wF = weight fraction of organic acid in feed. wL = weight fraction of organic acid in aqueous solution leaving effect. TF and PF be temperature and pressure of feed. T and P be temperature and pressure of leaving vapor and liquid. TS and PS be temperature and pressure of steam to effect. TC and PC be temperature and pressure of condensate from effect. Therefore, NV = 14. The relationships among the variables are: Total material balance: F = V + L Acid material balance: FwF = LwL Equilibrium temperature: T = f{P, wL}, i.e. a modified Raoult's law Energy balance: LHL + VHV - FHF = S(HS - HC) Therefore, NE = 4 From Eq. (5-67), ND = NV - NE = 14 - 4 = 10. A typical specification for a single evaporator effect might be F, wF, TF, PF, wL, P, TS, PS, TC , and PC . For the total condenser, the variables are: (Into) V, T, P; (out of) V, PT , TT ; Q Therefore, NV = 6. The relationships among the variables are: Energy balance: Q = V(HV in - HV out) Analysis: (continued) Exercise 5.35 (continued) Therefore, NE =1 and ND = 6 - 1 = 5. Typical specs might be V, T, P, PT , TT For the pump, the variables are L, wL, TL, PL, PP with no equations and 5 degrees of freedom. Now analyze the system. From Eq. (5-70), ND system = all units ND u - interconnecting redundant variables ND u (1) = 3(10) + 5 + 5 = 40 all units The number of interconnecting redundant variables are 4 (L, wL, T, P) for each of 3 liquid streams and 3 (V, T, P) for each of 3 vapor streams, or 4(3) + 3(3) = 21 Therefore, from Eq. (1), ( N D )system = 40 - 21 = 19 The given specifications are as follows, where effects are numbered 1, 2, 3 from left to right in Fig. 5.23: wF1 , w L3 , PS1 , TS1 . Therefore, additional specifications needed = 19 - 4 = 15, which might be: Feed rate to Effect 1 = F1 Temperature and pressure of feed to Effect 1 = TF1 , PF1 Pressure in each effect = P1, P2, P3 Temperature and pressure leaving condenser = TT, PT Pressure leaving the pump = PP Pressure of condensed steam leaving each effect = pressure of steam entering the effect Temperature of condensed steam leaving each effect = saturation for the exit pressure Exercise 5.36 Subject: Degrees of freedom and additional specifications for a reboiled stripper. Given: Reboiled stripper with specifications in Fig. 5.28. Find: (a) (b) (c) (d) Number of variables Number of relationships. Number of degrees of freedom. Additional specifications, if any. Analysis: (a) and (b) For an N -1 cascade and a partial reboiler from Table 5.3 NV NE Element N -1 cascade 7(N-1)+2C(N-1)+2C+7 5(N-1)+2C(N-1)+2 Partial reboiler 3C+10 2C+6 Total 7N+2NC+3C+10 5N+2NC+3 From Eq. (5-68), From Eq. (5-69), NV NE unit = = NV all elements e - NR C + 3 + N A - NR (1) (2) unit NE all elements e Since NR = 2 and NA = 0, Eqs. (1) and (2) become, ( NV )unit = (7N +2NC +3C +10) - 2 ( C + 3) = 7 N + 2 NC + C + 4 ( N E )unit = 5N +2NC +3 - 2 = 5N +2NC + 1 (c) From Eq. (5-67), ND = NV - NE = (7N+2NC+C+4) - (5N+2NC+1) = 2N + C + 3 (d) Specifications are: All feed conditions Number of stages Total C+2 1 C+3 Therefore, can make 2N additional specifications, which might be: All stage and reboiler pressures N Adiabatic stages N -1 Have one specification left. Make it the bottoms rate, distillate rate, or boilup rate or ratio. Exercise 5.37 Subject: Degrees of freedom and specifications for thermally coupled distillation. Given: Distillation system in Fig. 5.29. Find: (a) (b) (c) (d) Number of variables Number of relationships. Number of degrees of freedom. Reasonable set of specifications. Analysis: Treat the system as one of two units with NR = 4 interconnecting streams and NA = 0. (a) and (b) The feed unit is like an absorber with an additional intermediate feed. Therefore, from Table 5.3 for an N-stage cascade, NV = (7 N + 2 NC + 2C + 7) + (C + 3) for extra feed +1 for extra feed location = 7 N + 2 NC + 3C + 11 N E = (5 N + 2 NC + 2) + 1 for extra feed = 5 N + 2 NC + 3 The second unit is a distillation column like that in Example 5.5, except with an additional feed and two more sidestreams that are located at the two feed stages. Therefore, from Example 5.5, using M for the number of stages, including the reboiler, NV = (7 M + 2 MC + 5C + 20) + 3(C + 3) for extra feed and 2 extra sidestreams +1 for extra feed location = 7 M + 2 MC + 8C + 30 N E = (5 M + 2 MC + 4C + 9) + 2(C + 2) for 2 extra sidestreams + 1 for extra feed = 5 M + 2 MC + 6C + 14 From Eq. (5-68), NV system = NV u - N R C + 3 + N A (1) all units From Eq. (5-69), NE system = all units NE u - NR (2) Therefore, since NR = 4 and NA = 0, Eqs. (1) and (2) become, ( NV )system = (7 N + 2 NC + 3C + 11)+ ( 7 M + 2 MC + 8C + 30 ) - 4 ( C + 3) ( N E )system = 7( N + M ) + 2C ( N + M ) + 7C + 29 = (5N +2NC +3)+(5M +2MC +6C +14) - 4 = 5(N + M )+2C ( N + M ) + 6C + 13 Analysis: (d) Exercise 5.37 (continued) (c) From Eq. (5-71), ( N D )system = ( NV )system - ( N E )system = [ 7( N + M ) + 2C ( N + M ) + 7C + 29] - [5( N + M ) + 2C ( N + M ) + 6C + 13] = 2( N + M ) + C + 16 A reasonable set of specifications is as follows: Pressure at each stage, reboiler, condenser, reflux divider Heat transfer at each stage and, divider Complete feed specification Number of stages, N Number of stages, M Feed stage location Locations of 3 sidestreams Flow rates of 3 sidestreams Reflux ratio Distillate rate Saturated liquid reflux or reflux temperature if subcooled Total N+M+2 N+M C+2 1 1 1 3 3 1 1 1 2(N + M) + C + 16 Exercise 5.38 Subject: Degrees of freedom and specs for a column with a liquid sidestream. Given: Column in Fig. 5.30. Find: (a) Number of degrees of freedom. (b) Reasonable set of design variables. Analysis: (a) The column in Fig. 5.30 has a partial condenser. A similar column with total condenser is analyzed in Example 5.5, giving ND = 2N + C + 11. In Table 5.2, in operations (b) and (c), it is seen that a column with a partial condenser has 3 less degrees of freedom than a column with a total condenser. Therefore, here, ND = (2N + C + 11) - 3 = 2N + C + 8. (b) A reasonable set of specifications is: Pressures of stages, reboiler, condenser N+1 Heat transfer for stages N -1 Feed completely specified C+2 Feed stage and sidestream stage locations 2 Distillate (volatile impurities) rate 1 Reflux ratio 1 Number of stages 1 Sidestream rate 1 Total 2N + C + 8 Exercise 5.39 Subject: Degrees of freedom analysis for distillation with a sidecut reboiled stripper. Given: Distillation system in Fig. 5.31. Find: (a) (b) (c) (d) Number of variables. Number of equations. Number of degrees of freedom. Reasonable set of design variables. Analysis: The system consists of (1) distillation column with 2 feeds and a sidestream, (2) reboiled stripper, (3) cooler, and (4) valve, with NR = 4 and NA = 0. (a) and (b) From Eq. (5-68), NV system = NV u - N R C + 3 + N A (1) all units From Eq. (5-69), NE system = all units NE u - NR (2) For the distillation column, NV and NE can be obtained by adding a second feed to the results given in Example 5.5. This adds C + 3 to NV and 1 to NE. Therefore, NV = (7N+2NC+5C+20) + (C+3) = 7N+2NC+6C+23 and NE = (5N+2NC+4C+9) + 1 = 5N+2NC+4C+10 For the reboiled stripper, the determination of NV and NE is given in Exercise 5.36. For the cooler, use the total condenser in Table 5.1. For the valve, use the total condenser in Table 5.1, less one variable for no heat transfer. Unit (or element) NV NE Distillation column 7N+2NC+6C+23 5N+2NC+4C+10 Reboiled stripper 7M+2MC+C+4 5M+2MC+1 Cooler 2C+7 C+3 Valve (no heat transfer) 2C+6 C+3 Total 7(N+M)+2(N+M)C+11C+40 5(N+M)+2(N+M)C+6C+17 Using Eqs. (1) and (2), ( NV )system = 7(N +M )+2(N +M )C +11C +40 - 4 ( C + 3) ( N E )system = 7(N +M )+2(N +M )C + 7C + 28 = 5(N +M )+2(N +M )C +6C +17 - 4 = 5(N +M )+2(N +M )C + 6C + 13 (c) From Eq. (5-71), ( N D )system = ( NV )system - ( N E )system = [ 7(N +M )+2(N +M )C + 7C + 28] - [5(N +M )+2(N +M )C + 6C + 13] = 2( N + M ) + C + 15 Exercise 5.39 (continued) Analysis:. (d) A reasonable set of specifications is: Pressures of stages, 2 reboilers, condenser Pressures of cooler, reflux divider, valve Heat transfer for stages and divider Feed completely specified Feed stage and sidestream stage locations Distillate rate Reflux ratio Number of stages for N and M Sidestream rate Boilup ratio in stripper Saturated liquid reflux Exit temperature from cooler Total N+M+1 3 N +M-1 C+2 2 1 1 2 1 1 1 1 2(N + M) + C + 15 Exercise 5.40 Subject: Allowable specifications for system of extractive distillation and distillation Given: System in Fig. 5.32 Find: If given specifications are sufficient. If not, give additional specifications. Analysis: System consists of a distillation column with 2 feeds (Case h in Table 5.4), a distillation column (Case b in Table 5.4), a cooler (use total condenser in Table 5.3), and a mixer (Table 5.4). From these sources, the degrees of freedom areas follows: Unit ND Extractive distillation 2N1+2C+12 Distillation 2N2+C+9 Cooler C+4 Mixer 2C+6 Total 2(N1+N2)+6C+31 From Eq. (5-70), ND unit = ND all elements e - NR C + 2 + N A (1) NR = 4 and NA = 0 Therefore, Eq. (1) becomes, N D unit = 2 N 1 + N 2 + 6C + 31 - 4 C + 2 = 2 N 1 + N 2 + 2C + 23 A reasonable set of specifications is: Pressures of stages, 2 reboilers, 2 condensers N1 + N2 + 2 Pressures of cooler, 2 reflux dividers, mixers 4 Heat transfer for stages, 2 dividers, mixer N 1 + M2 + 1 Feed completely specified C+2 Makeup specified except flow rate C+1 Two feed stage locations 2 Distillate rate 1 Two reflux rates 2 Number of stages for N1 and N2 2 Phenol recycle rate 1 Total 2(N1 + N2) + 2C + 18 Therefore, 5 more specifications are needed. They might be the following: Saturated liquid reflux from 2 condensers 2 Feed stage location for extractive distillation 1 Temperature leaving cooler 1 Makeup phenol rate 1 Exercise 5.41 Subject: Specifications for a distillation column with sidestream Given: Distillation column in Fig. 5.33 with specifications. Find: If given specifications are sufficient. If not, give additional specifications. Analysis: For a distillation column with a total condenser and sidestream, Example 5.5 applies, giving ND = 2N + C +11. The following specifications are given: Pressures of stages, reboiler, condenser, divider Heat transfer for stages and divider Feed completely specified Feed stage and sidestream stage locations Distillate purity Sidestream rate and purity Number of stages Total N+2 N C+2 2 1 2 1 2N + C + 10 Therefore, one more specification is needed. It might be saturated liquid leaving condenser or temperature of reflux if subcooled. Exercise 6.1 Subject: Given: Simultaneous absorption and stripping. Fig. 6.1 in which simultaneous absorption and stripping occur. Find: Whether absorption and stripping is more important. Stripping of water = 22 kmol/h Absorption of acetone = 10.35 kmol/h Therefore, more stripping than absorption occurs. However, the operation is primarily absorption because a very large percentage, 99.5%, of the acetone is absorbed. Only a very small fraction of the water is stripped. Analysis: Exercise 6.2 Subject: Given: Column packings since 1950. History of the development of column packings. Find: Advantages of new packings. Advances in packing design and fabrication. Need for structured packings. Analysis: The newer packings provide more surface area for mass transfer, a higher flow capacity, and a lower pressure drop. They provide more effective through flow. Many are plastics made from molds or thin metal strips that can be inexpensively fabricated into intricate shapes. Structured packings largely eliminate the problems of channeling and at the same time give improvement in efficiency, capacity, and pressure drop. Exercise 6.3 Subject: Advantages of bubble-cap trays Given: Characteristics and performance of bubble-cap trays Find: Characteristics that give bubble-cap trays a very high turndown ratio. Analysis: Unlike sieve and valve trays, bubble-cap trays do not allow liquid to weep. Therefore, bubble-cap trays can be operated at very low liquid flow rates. Also, bubble caps force the vapor to flow out sideways, rather than vertically up, thus allowing a relatively high vapor rate. Exercise 6.4 Subject: Given: Selection of alternative absorbent for Example 6.3 Flow rate, density, and MW of three potential absorbents listed below. Find: Select the best absorbent with reasons why. Are any of the absorbents unacceptable? Analysis: In Example 6.3, the rich gas contains C1 to nC6 hydrocarbons, with mostly C3. Object of absorber is to absorb most of the nC4. The absorbent is an oil of 250 MW and 21oAPI, at a flow rate, L, of 368 lbmol/h. An alternative absorbent must also be of higher molecular weight than nC4 and must have a flow rate of at least 368 lbmol/h. If its molecular weight is too low, a significant amount of it will be stripped. This can be judged by its K-value and stripping factor, S=KV/L. The entering gas rate, V, is 946 lbmol/h. Column operating conditions are about 94 psia and a maximum temperature of 126oF. Pertinent properties and factors for the three potential absorbents are: Absorbent C5 s light oil medium oil gpm 115 36 215 , lb/gal 5.24 6.0 6.2 lb/h 36,200 13,000 80,000 MW 72 130 180 lbmol/h 500 100 440 K-value 0.9 0.005 0.0005 S=KV/L 2.3 0.013 0.0013 The light oil can not be used because its flow rate of 100 lbmol/h is much lower than the necessary 368 lbmol/h. The C5s can not be used because their stripping factor is very high. The only possible alternative is the medium oil. Exercise 6.5 Subject: Stripping of VOCs from water effluents by air and water Given: Packed tower for stripping Find: Advantages and disadvantages of air over steam Advantages: Air is available anywhere. Air is inexpensive. Disadvantages: Air can form a flammable or explosive mixture with the VOC. With steam, the exit gas can be condensed and the VOC recovered as a liquid. Analysis: Exercise 6.6 Subject: Preferred operating conditions for absorbers and strippers. Find: Best conditions of temperature and pressure, and the trade-off between number of stages and flow rate of separating agent, using equations. Analysis: Absorbers: For high performance, want a large absorption factor, A = L/KV Therefore, want a small K-value. Assume that: iL Pi s Ki = P s Pi = vapor pressure, which increases with increasing temperature Therefore, operate at a high pressure and a low temperature. Strippers: For high performance, want a large stripping factor, S = KV/L Therefore, want a large K-value. Therefore, operate at low pressure and a high temperature. For the tradeoff between number of equilibrium stages, N, and flow rate of mass separating agent, L or V. Consider absorption. The fraction of a component absorbed is given from a modification of Eq. (5-48), A -1 A N +1 - A 1 - A = 1 - N +1 = N +1 A -1 A -1 Thus, a large fraction absorbed can be achieved with either N or A = L/KV. The tradeoff is most clearly shown in Fig. 5.9. For low fractions absorbed (i.e. high A), N has little effect and (1 A) is approximately equal to A. But for high fractions absorbed, the larger the value of N, the smaller the required value of A, and, thus, the smaller the required value of the flow rate of the liquid separating agent, L. The tradeoff for stripping is similar. The fraction of a component stripped is given by a modification of Eq. (5-51), S -1 S N +1 - S 1 - S = 1 - N +1 = N +1 S -1 S -1 Thus, a large fraction stripped can be achieved with either N or S = KV/L. The tradeoff is most clearly shown in Fig. 5.9. For low fractions stripped (i.e. high S), N has little effect and (1 - S) is approximately equal to S. But for high fractions absorbed, the larger the value of N, the smaller the required value of S, and, thus, the smaller the required value of the flow rate of the vapor separating agent, V. xercise 6.7 Subject: Absorption of CO2 from air at 25oC by 5-N aqueous triethanolamine Given: Feed gas containing 10 mol% CO2 and 90 mol% air. Absorbent of 5N aqueous triethanolamine containing 0.04 moles of CO2 per mole of amine solution. Column with 6 equilibrium stages. Exit liquid to contain 78.4% of the CO2 in the feed gas. Therefore, exit gas contains 21.6% of the CO2 in the entering gas. Equilibrium data for CO2 at 25oC in terms of mole ratios. Assumptions: Negligible absorption of air and stripping of amine and water. Find: (a) Moles of amine solution required per mole of feed gas. (b) Exit gas composition. Analysis: Use the nomenclature and type of plot shown in Fig. 6.11(a). Therefore, for CO2, X0 = 0.04 mol CO2/mol amine solution YN+1 = Y7 = 10/90 = 0.1111 mol CO2/mol air Y1 = 0.216(10)/90 = 0.024 mol CO2/mol air (b) Therefore, the exit gas composition is 0.024 mol CO2/mol air or 0.024/(1 + 0.024) x 100% = 2.34 mol% CO2 and 97.66 mol% air. (a) A plot of the equilibrium data as Y vs. X is given below. The operating point (X0, Y1) at the top of the column is included. A straight operating line through this point is found by trial and error to give 6 equilibrium stages, when using Y7 = 0.1111. The resulting XN = X6 = 0.085. From Eq. (6-3), the slope of the operating line = L'/V' = (0.1111 - 0.024)/(0.085 - 0.04) = 1.936 mol triethanolamine solution/mol air. The feed gas contains 9 mol air/10 mol feed gas. Therefore, mols of amine solution/mol feed gas = 1.936(0.9) = 1.74. See plot on next page. Exercise 6.7 (continued) Exercise 6.8 Subject: Absorption of acetone from air by water at 20oC and 101 kPa (760 torr) in a valve-tray column. Given: 100 kmol/h of feed gas containing 85 mol% air and 15 mol% acetone. Pure water is the absorbent. Overall tray efficiency is 50%. Absorb 95% of the acetone. Equilibrium p-x data for acetone are given as listed below. Assumptions: Negligible absorption of air and stripping of water. Find: (a) Minimum ratio, L'/V' of moles of water/mole of air. (b) Number of equilibrium stages for L'/V' = 1.25 times minimum. (c) Concentration of acetone in the exit water. Analysis: Use the nomenclature and type of plot shown in Fig. 6.11(a). Then, the operating line will be straight. For acetone, X0 = 0.0 mol acetone/mol entering water YN+1 = 0.15/0.85 = 0.1765 mol acetone/mol air in entering gas Flow rate of acetone in exit gas = (1 - 0.95)(15) = 0.75 kmol/h. With 85 kmol/h of air, Y1 = 0.75/85 = 0.00882 mol acetone/mol air Convert the p-x equilibrium data to mole ratio, Y-X data, using y = p/P, Y = y/(1- y), X = x/(1-x) p, torr x y X Y 30.0 0.033 0.0395 0.0341 0.0411 62.8 0.072 0.0826 0.0776 0.0901 85.4 0.117 0.1124 0.1325 0.1266 103.0 0.171 0.1355 0.2063 0.1568 (a) With the type of curvature in the Y-X equilibrium curve, shown below, the minimum absorbent rate is determined by a straight operating line that passes through the point (Y1 , X0 ) and is drawn tangent to the equilibrium curve, as shown. From Eq. (6-3), the slope of the operating line = L'/V' = 1.06 mol water/mol of air on an acetone-free basis = minimum ratio. (b) For 1.25 times minimum, L'/V' = 1.25(1.06) = 1.325. Now a straight operating line that passes through the point (Y1 = 0.00882 and X0 = 0.0) is drawn as shown below, where equilibrium stages are stepped off, as in Figs. 6.11(a) and 6.12, until the steps reach or exceed the required value of YN+1 = 0.1765. As seen, in the figure below, the determined number of equilibrium stages is approximately 9. (c) The concentration of acetone in the exit liquid, XN is the point on the operating line for Part (b), where YN+1 = 0.1765. From Eq. (6-3), for the operating line for an absorber, Y -Y 0.1765 - 0.00882 X N = N +1 1 + X 0 = + 0.0 = 0.127 mol acetone/mol water L 1.325 V or an exit mole fraction of acetone in the water of 0.127/(1.127) = 0.113 Analysis: (continued) Exercise 6.8 (continued) Exercise 6.9 Subject: Absorption-stripping system for benzene (B) removal from gas by oil and subsequent stripping of benzene from oil by steam. Given: Plate column absorber and stripper. Feed gas containing 0.06 mol B/mol B-free gas. Absorbent oil containing 0.01 mol B/mol B-free oil. Liquid leaving absorber contains 0.19 mol B/mol B-free oil. 90% of B in the gas stream is absorbed. Liquid leaving stripper contains 0.01 mol B/mol oil. Flow ratio of B-free oil-to-B-free steam = 2.0. Molecular weights are 200, 78, and 32 for oil, benzene, and gas, respectively. Equilibrium X-Y data given for absorber (25oC) and stripper (110oC). Assumptions: Gas is not soluble in oil and oil is not volatile. Find: (a) Molar flow rate ratio of B-free oil to B-free gas in the absorber. (b) Number of equilibrium stages in the absorber. (c) Minimum steam flow rate in the stripper per mol of B-free oil. Analysis: A flow diagram of the absorber-stripper system is shown in Fig. 5.10(a), except that there is no makeup absorbent oil because the oil is assumed to be non-volatile. (a) Consider just the absorber. Take a basis of V ' = 1 mol/h of B-free gas entering the absorber. Because YN+1 = 0.06 mol B/mol B-free gas, we have 0.06 mol/h of B in entering gas. With 90% absorption, 0.06(0.9) = 0.054 mol/h of B absorbed, leaving 0.006 mol/h of B in the exiting gas. In the entering absorbent oil, X0 = 0.01 mol B/mol B-free oil. In the liquid leaving the absorber, XN = 0.19 mol B/mol B-free oil. A material balance around the absorber on the benzene gives: YN +1V + X 0 L = YV + X N L Therefore, 1 0.06(1.0)+0.01L = 0.006 + 0.19 L (1) Solving Eq. (1), L' = 0.30 mol/h of B-free oil, giving L'/V'' = 0.30 mol B-free oil to B-free gas. (b) Use a graphical method to determine the number of equilibrium stages in the absorber, using Y-X coordinates to give a straight operating line. Data for the equilibrium curve of the absorber are given and the operating line is obtained from terminal values of Y and X. The results are given in the plot below. As seen, between 9 and 10 equilibrium stages are needed. (c) The equilibrium curve for the stripper is shown in the plot below. With the type of curvature in the Y-X equilibrium curve the minimum stripping gas rate is determined by a straight operating line that passes through the operating point (Y0 = 0 , X1 = 0.01) at the bottom of the column, as shown in Fig. 6.11(b) and is drawn tangent to the equilibrium curve, as shown. From Eq. (6-5), the slope of the operating line = L'/V' = 3.0 mol B/mol of steam on a B-free basis. Therefore, (V'')min/L' = 0.33. This is compared to the given operating value of L'/V' = 2.0 or V'/L' = 0.5. Analysis: (continued) (a) Plot of results: Exercise 6.9 (continued) (b) Plot of results: Exercise 6.10 Subject: Stripping of benzene (B) from straw oil by steam at 1 atm (101.3 kPa). Given: Oil enters stripper with 8 mol% benzene. 75% of B is stripped. Pure steam enters. Steam exits with 3 mol% B. Sieve-plate column. Henry's law holds with a B partial pressure of 5.07 kPa when benzene mole fraction in the oil is 10 mol%. Assumptions: Straw oil is not volatile and steam does not condense. Dalton's law. Find: (a) Number of equilibrium stages required. (b) Moles of steam required per 100 moles of benzene-oil feed to stripper. (c) Number of equilibrium stages needed if 85% of B is stripped with same amount of steam as in Part (b). Analysis: With reference to the stripper in Fig. 6.11(b), XN+1 = 8/92 = 0.087 mol B/mol B-free oil Y0 = 0.0 mol B/mol B-free steam YN = 3/97 = 0.0309 mol B/mol B-free steam (b) For 100 moles of benzene-oil feed mixture, have 8 moles of benzene. Amount of benzene stripped = 0.75(8) = 6 moles. Because the exit gas contains 0.0309 mol B/mol B-free steam, the steam rate = 6/0.0309 = 194 moles. Therefore, moles of steam/100 moles of benzene-oil feed = 194 Also, X1 = (8 - 6)/92 = 0.0217 mol B/mol B-free oil (a) Assume Henry's law is given by Eq. (4-32), pB = HxB.. Because pB = 5.07 kPa when xB = 0.10, H = 5.07/0.1 = 50.7 kPa. Convert this to an a Y-X equilibrium equation. For a pressure of 101.3 kPa, assuming Dalton's law, yB = pB/P = 50.7xB/101.3 = 0.5xB . From Eq. (6-1), dropping the benzene subscript, B, and noting that KB = 0.5, KB = 0.5 = Y / (1 + Y ) X / (1 + X ) Solving, Y = 0.5 X 1 + 0.5 X (1) The Y-X plot includes the equilibrium curve given by Eq. (1), and a straight operating line that connects the column terminal points (Y0 = 0.0, X1 = 0.0217) and (YN = 0.0309, XN+1 = 0.087). The corresponding slope of the operating line is (0.0309 - 0.0)/(0.087 - 0.0217) = L'/G' = 0.474. From the plot below, stepping off stages as in Fig. 6.11(b), required N = approximately 3 stages. (c) For 85% stripping of B, amount of B stripped = 0.85(8) = 6.8 moles B. Not stripped is1.2 moles B. Therefore, for the same B-oil and steam feeds, X1 = 1.2/92 = 0.01304 mol B/mol oil and YN =6.8/194 = 0.03505 mol B/mol steam. The Y-X plot for this case is also given below, where the equilibrium curve is the same as in Part (a), and a straight operating line connects the column terminal points (Y0 = 0.0, X1 = 0.01304) and (YN = 0.03505, XN+1 = 0.087). The corresponding slope of the operating line is (0.03505 - 0.0)/(0.087 - 0.01304) = L'/G' = 0.474 (same as in Part (a)). From the plot below, stepping off stages as in Fig. 6.11(b), required N = between 5 and 6 stages. Exercise 6.10 (continued) Analysis: (continued) (a) Plot of results: (b) Plot of results: Exercise 6.11 Subject: Stripping of VOCs from water by air in a trayed tower to produce drinking water. Given: 1,500 gpm of groundwater containing ppm amounts of DCA, TCE, and TCA given below. Stripping at 1 atm and 25oC to reduce the ppm amounts to the very low levels below. K-values of VOCs given below. Assumptions: No stripping of water and no absorption of air. System is dilute with respect to the VOCs. Find: Minimum air flow rate in scfm (60oF and 1 atm). Number of equilibrium stages for 2 times the minimum air flow. Composition in ppm for each VOC in the resulting drinking water. Analysis: Flow rate of water = L = (1,500 gpm)(8.33 lb/gal)/18.02 lb/lbmol = 693 lbmol/min % stripping of a VOC = (inlet ppm - outlet ppm)/inlet ppm) VOC K-value Inlet ppm Outlet ppm % stripping DCA 60 85 0.005 99.994 TCE 650 120 0.005 99.996 TCA 275 145 0.200 99.862 Because of its high % stripping and a K-value that is much lower than for the other two VOCs, the stripper is likely to be controlled by DCA. So base the calculations on DCA and then check to see that the % stripping for TCE and TCA exceed the above requirements. Because of the dilute conditions, use Kremser's method. 693 L From Eq. (6-12), Vmin = (fraction stripped) = (0.99994) = 11.55 lbmol./min K 60 or minimum air flow rate = 11.55 (379 scf/lbmol) = 4,377 scfm at 60oF and 1 atm For 2 times the minimum value, V = 2(11.55) = 23.1 lbmol/min. K V 60(23.1) The stripping factor, given by Eq. (6-16), is for DCA, S DCA = DCA = = 2.0 L 693 S N +1 - S 2 N +1 - 2 From Eq. (6-14), Fraction stripped = 0.99994 = N +1 = (1) S - 1 2 N +1 - 1 Solving Eq. (1), N = 13 stages. Now, for V = 11.55 lbmol/min and N = 13 stages, compute the fraction stripped for TCE and TCA from Eq. (6-14) with the stripping factor from Eq. (6-16). The results are: VOC S Fraction stripped DCA 2.0 0.99994 TCE 21.6 1.00000 TCA 9.17 1.00000 The drinking water contains 0.005 ppm of DCA and essentially zero ppm of TCE and TCA. Exercise 6.12 Subject: Stripping of a solution of SO2, butadienes, and butadiene sulfone with nitrogen. Given: 120 lbmol/h of liquid containing in lbmol/h,10.0 SO2, 8.0 1,3-butadiene (B3), 2.0 1,2butadiene (B2), and 100.0 butadiene sulfone (BS). Liquid effluent to contain < 0.05 mol% SO2 and < 0.5 mol% (B3 + B2). Stripping agent is pure nitrogen. Stripping at 70oC and 30 psia. Kvalues are 6.95 for SO2, 3.01 for B2, 4.53 for B3, and 0.016 BS. Assumptions: Negligible stripping of BS because of its low K-value. No absorption of N2. Find: Flow rate of nitrogen. Number of equilibrium stages. Analysis: Assuming no stripping of BS, the exiting liquid will contain 0.05 mol% SO2, 0.50 mol% (B3 + B2), and 99.45 mol% BS, or in lbmol/h, 100.0 BS, 0.0503 SO2, and 0.503 (B3 + B2). Therefore, % stripping of SO2 = (10.0 - 0.0503)/10.0 = 0.99497 and % stripping of (B3 + B2) = (8.0 + 2.0 - 0.503)/(8.0 + 2.0) = 0.9497. The minimum stripping agent flow rate is given by Eq. (6-12). Assume that because the Kvalue of B2 is the lowest of SO2, B2, and B3, that B2 controls. Further assume that essentially all of the B3 is stripped so that the 0.5 mol% of (B2 + B3) is entirely B2. The solute-free liquid rate = flow rate of BS = L' = 100 lbmol/h and the fraction of B2 stripped = (2.0 - 0.503)/2.0 = 0.748 L' 100 ' (fraction stripped) = (0.748) = 24.9 lbmol/h Vmin = KSO2 3.01 Now use this value to compute the values of fraction stripped of SO2 and B3. KV ' 6.95(24.9) = > 1.0 100 L' KV ' 4.53(24.9) Fraction stripped of B3 = ' = > 1.0 L 100 Fraction stripped of SO 2 = Therefore, the assumption that B2 controls is correct at the minimum stripping gas rate. Assume an operational nitrogen flow rate of 1.5 times minimum. Therefore V' =1.5(24.9) = 37.4 lbmol./h. The stripping factor, given by Eq. (6-16), is for B2, S B2 = K B2V ' 3.01(37.4) = = 1.13 L' 100 S N +1 - S 113 N +1 - 113 . . = N +1 N +1 S -1 113 - 1 . Using Eq. (6-14), Fraction B2 stripped = 0.748 = Solving, N = 2.4 equilibrium stages. Analysis: (continued) Exercise 6.12 (continued) S N +1 - S 1.7 2.4 +1 - 17 . = = 0.863 N +1 2 . 4 +1 -1 S -1 17 . S N +1 - S 2.612.4 +1 - 2.61 = = 0.936 S N +1 - 1 2.612.4 +1 - 1 Fraction B3 stripped = Fraction SO 2 stripped = These fractions do not meet specifications. Therefore, the number of stages and/or the nitrogen flow rate must be increased. First try increasing the number of stages. S = KV/L Fraction stripped: N=5 N=7 N = 10 SO2 2.61 0.995 0.9993 0.99996 B2 1.13 0.880 0.922 0.954 B3 1.7 0.970 0.990 0.998 For N = 5 stages, have in the exit liquid in lbmol/h, 0.05 SO2, 0.24 B2, and 0.24 B3. These values correspond to a total of 100.53 lbmol/h of exit liquid. This results in 0.05 mol% SO2 and 0.48% (B2 + B3). Therefore, N = 5 stages is sufficient with a feed gas rate, V' , of 37.4 lbmol/h. Other combinations of N and V' could be used.. Exercise 6.13 Subject: Absorption of a light hydrocarbon gas mixture by n-decane as a function of pressure and number of equilibrium stages at 90oF. Given: Light hydrocarbon gas containing in lbmol/h, 1,660 C1, 168 C2, 96 C3, 52 nC4, and 24 nC5 for a total of 2,000 lbmol/h = V. Absorbent of L = 500 lbmol/h of nC10. K-value of nC10 = 0.0011. Find: Component flow rates in the lean gas and rich oil for: (a) N = 6 and P = 75 psia. (b) N = 3 and P = 150 psia. (c) N = 6 and P = 150 psia. Analysis: Use the Kremser method with Eqs. (5-48) for fraction not absorbed, A, (5-50 for fraction not stripped, A ,(5-38) for absorption factor, A, (5-51) for stripping factor, S , (5-45) and (5-53) for component flow rates in the exit gas, 1 , and (5-44) and (5-52) for component flow rates in the exit liquid, lN, since no light hydrocarbons enter with the absorbent. Use Fig. 2.8 or other source for K-values of C1 through nC5. For nC10 , assume the K-value is inversely proportional to pressure. (a) N = 6 and P = 75 psia: K-value 2.9 6.5 1.95 0.61 0.19 0.0011 A 0.0086 0.0385 0.1282 0.4098 1.3160 S A 0.9914 0.9615 0.8718 0.5916 0.0541 S lbmol/h 1,646.0 161.5 83.7 30.8 1.3 2.2 1,925.5 Component C1 C2 C3 nC4 nC5 nC10 Total (b) 1 0.0044 0.9956 l6 lbmol/h 14.0 6.5 12.3 21.2 22.7 497.8 574.5 N = 3 and P = 150 psia: K-value 14.0 3.5 1.05 0.33 0.105 0.00055 A 0.0179 0.0714 0.2381 0.7576 2.3810 S A 0.9821 0.9284 0.7644 0.3615 0.0443 S lbmol/h 1,630.0 156.0 73.4 18.8 1.1 1.1 1,880.4 Component C1 C2 C3 nC4 nC5 nC10 Total 1 0.0022 0.9978 l3 lbmol/h 30.0 12.0 22.6 33.2 22.9 498.9 619.6 Exercise 6.13 (continued) Analysis: (continued) (c) N = 6 and P = 150 psia: K-value 14.0 3.5 1.05 0.33 0.105 0.00055 A 0.0179 0.0714 0.2381 0.7576 2.3810 S A 0.9821 0.9284 0.7619 0.2829 0.0032 S lbmol/h 1,630.0 156.0 73.1 14.7 0.1 1.1 1,875.0 Component C1 C2 C3 nC4 nC5 nC10 Total 1 0.0022 0.9978 l6 lbmol/h 30.0 12.0 22.9 37.3 23.9 498.9 625.0 Exercise 6.14 Subject: Absorption of a light hydrocarbon gas by nC10. Given: 1,000 kmol/h of rich gas at 70oF with, in kmol/h, 250 C1, 150 C2, 200 C3, 200 nC4, and 150 nC5. 500 kmol/h of nC10 at 90oF. Absorber operates at 4 atm. K-value of nC10 at 80oF and 4 atm = 0.0014. Assumptions: Using the Kremser method, use V = 1,000 kmol/h and L = 500 kmol/h, with an average temperature of 80oF for K-values from Fig. 2.8 or other source. Find: Percent absorption of each component for 4, 10, and 30 equilibrium stages. Analysis: The K-values and absorption factors, from Eq. (6-15), A = L/KV, are: Component C1 C2 C3 nC4 nC5 K-value 38 7.6 2.25 0.64 0.195 A 0.0132 0.0658 0.2222 0.7813 2.564 For each component in the feed gas, the percent absorption is given by Eq. (6-13) x 100%, A N +1 - A Percent absorbed = N +1 100% A -1 (1) Using Eq. (1) with the values of A from the above table, the results are: Component C1 C2 C3 nC4 nC5 Percent absorbed: for N = 4 for 1.32 6.58 22.03 65.14 96.30 N = 10 1.32 6.58 22.22 76.58 99.995 for N = 30 1.32 6.58 22.22 78.12 100.00 The number of stages affects only the heavier hydrocarbons that are absorbed to the greatest extent. Exercise 6.15 Subject: Back calculation of tray efficiency from performance data on absorption of propane. Given: Large commercial absorber with Na = 16 actual plates. From Example 6.3, average total liquid rate, L, is 446.7 lbmol/h, average total gas rate, V, is 866.5 lbmol/h. Propane in gas feed = 213.8 lbmol/h. Propane in exit liquid = 43 lbmol/h. P = 93.7 psia. Average T = 111oF. Find: Tray efficiency from performance data. Estimated tray efficiency from DrickamerBradford and O'Connell correlations. Analysis: Fraction of propane absorbed = 43/213.8 = 0.20 From Fig. 2.8, for propane, K-value = 2.0 From Eq. (6-15), A = L/KV = 446.7/(2.0)(866.5) = 0.258 From Eq. (6-13), A N +1 - A 0.258 N +1 - 0.258 Fraction absorbed = 0.20 = N +1 = A -1 0.258 N +1 - 1 Solving Eq. (1), N = Nt = 0.93 From Eq. (6-1), Eo = Nt /Na = 0.93/16 = 0.058 or 5.8% From Drickamer-Bradford correlation, Eq. (6-22), with L = 1.4 cP from Example 6.3, Eo , % = 19.2 - 57.8 log L = 19.2 - 57.8 log(1.4) = 10.8 % From O'Connell correlation, Eq. (6-23), with ML = 250, L = 57.9 lb/ft3 from Example 6.3, log E o = 1587 - 0.199 log C - 0.0896 log C . where, C = 2 (1) (2) KM L L (2.0)(250)(1.4) = = 12.1 L 57.9 Therefore, from Eq. (2), log E o = 1587 - 0.199 log 12.1 - 0.0896 log 12.1 = 1277 . . which gives Eo = 18.9% The propane tray efficiency is lower than predicted. 2 Exercise 6.16 Subject: Oil absorption of a gas to produce 95% hydrogen Given: 800,000 scfm (0oC, 1 atm) of fuel gas containing 72.5% H2, 25% CH4, and 2.5% C2H6. Absorbent of nC8. Absorber operates at 400 psia and 100oF. Exiting gas is to contain at least 80% of the entering H2 at a H2 purity of 95 mol%. Assumptions: Ideal gas law. Find: (a) (b) (c) (d) (e) (f) (g) Minimum absorbent rate in gpm. Actual absorbent rate at 1.5 times minimum. Number of theoretical stages. Stage efficiency of each component from O'Connell correlation. Number of actual trays needed. Composition of exit gas, accounting for stripping of octane. Annual cost of lost octane. Analysis: At 0oC and 1 atm, have 359 scf/lbmol. Therefore, entering V = 800,000/359 = 2,230 lbmol/min, with 0.725(2,230) = 1,616 lbmol/min H2, 558 lbmol/min CH4, 56 lbmol/min C2H6 . From Fig. 2.8, at 400 psia, 100oF, K-values are: 30 for H2 , 7 for CH4 , and 1.75 for C2H6 . (a) Assume the key component is CH4. If we neglect stripping of octane and assume complete absorption of C2H6, with 20% absorption of H2, then exit gas will be 0.8(1,616) = 1,293 lbmol/min H2 and the exit liquid will contain 1,616 - 1,293 = 323 lbmol/min H2. For a purity of 95 mol% H2, then will have 5 mol% CH4 or 1,293(5/95) = 68 lbmol/min of CH4 in the exit gas and 558 - 68 = 490 lbmol/min of CH4 in exit liquid. Can not use Eq. (6-11) to compute the minimum absorbent rate because we do not have a dilute system. Instead, for infinite stages, assume the exiting rich oil is in equilibrium with the inlet gas. Then, for CH4, K = 7 = y/x. But y for the inlet gas = 0.25. Therefore, x in the exiting oil = 0.25/7= 0.0357. Therefore, with 490 lbmol/min of CH4 in this oil, the total exiting oil = 490/0.0357 = 13,700 lbmol/min. Therefore, the entering oil is 13,700 - 323 - 490 - 56 = 12,831 lbmol/min of nC8 = Lmin. For nC8 , MW = 114 and from Perry's Handbook, SG = 0.703. Therefore, liquid density = 0.703(8.33) = 5.86 lb/gal. Therefore, Lmin = 12,831(114)/5.86 = 250,000 gpm. (b) For 1.5 times minimum, inlet absorbent rate = 19,200 lbmol/min = 375,000 gpm. (c) Calculate the number of equilibrium stages from the Kremser equation based on CH4 absorption, inlet vapor and liquid flow rates. From Eq. (6-15), ACH 4 = L / KV = 19,200 / (7)(2,230) = 123 . From Eq. (6-13), for CH4, with a fraction absorbed = 490/558 = 0.878, fraction of CH4 absorbed = 0.878 = N ACH+41 - ACH 4 N ACH+41 - 1 123 N +1 - 123 . . = N +1 1.23 - 1 (1) Analysis: (c) (continued) Exercise 6.16 (continued) Solving Eq. (1), N =4.1 stages. Check absorption of H2. Absorption factor = A = 19,200/(30)(2,230) =0.287 N AH 2+1 - AH 2 0.287 4.1+1 - 0.287 From Eq. (6-13, fraction of H2 absorbed = = = 0.286 N 0.287 4.1+1 - 1 AH 2+1 - 1 This exceeds the minimum 20% specified. Actually, this problem is overspecified. The absorber can be designed for 95 mol% purity of H2 in the exit gas or 20% absorption of H2. In this problem, we can not achieve both specifications, unless the plate efficiency for H2 is much less than that for CH4. Check this next. (d) From the O'Connell correlation, Eq. (6-23), with ML = 114, L = 44 lb/ft3, and L= 0.47 cP from Perry's Handbook, log E o = 1587 - 0.199 log C - 0.0896 log C . where, for H2, C = Therefore, from Eq. (2), log E o = 1587 - 0.199 log 36.5 - 0.0896 log 36.5 = 1057 . . which gives Eo = 11.4 % for H2. Similarly, Eo = 21.1 % for CH4, and Eo = 32.5 % for C2H6. (e) From Eq. (6-21), using CH4, Na = Nt /Eo = 4.1/0.211 = 19.4 or say 20 trays. (f) For the stripping of nC8, from Fig. 2.8 by extrapolation, K = 0.003. From Eq. (2), Eo = 0.35. Therefore, if 20 trays, have 20(0.35) = 7 equilibrium stages. From Eq. (6-16), S = KV/L =0.003(2,230)/19,200 = 0.00035. From Eq. (6-14) for 7 stages, S N +1 - S 0.000357 +1 - 0.00035 fraction stripped = N +1 = = 0.00035 0.000357 +1 - 1 S -1 Therefore, 0.00035(19,200) = 7 lbmol/h of nC8 leaves in the exit gas. For H2 , with 20 trays and Eo = 0.114, we have (0.114)(20) = 2.3 equilibrium stages. From Eq. N AH 2+1 - AH 2 0.287 2.3+1 - 0.287 (6-13), with A = 0.287, fraction H2 absorbed = = = 0.275 N AH 2+1 - 1 0.287 2.3+1 - 1 Therefore, we can not meet the specification of less than 20% absorbed. For C2H6, with 20 trays and Eo = 0.325, we have (0.325)(20) = 6.5 equilibrium stages. From Eq. (6-13), with A = (19,200)/(1.75)(2,230) = 4.92, A N +1 - A 4.92 6.5+1 - 4.92 fraction C2H6 absorbed = = = 0.999975 A N +1 - 1 4.92 6.5+1 - 1 Composition of exit gas in lbmol/min is 1,172 H2 , 68 CH4 , and 7 nC8 . H2 purity = 94% (g) Cost of lost oil = 7(114)(60)(7,900)($1.00)/(5.86 lb/gal) = 64.6 million $/year. 2 2 (2) KM L L (30)(114)(0.47) = = 36.5 L 44 Exercise 6.17 Subject: Scale-up of absorber using Oldershaw column efficiency. Given: Absorption operation of Examples 6.1 and 6.4 with a column diameter of 3 ft. New column with 11.5 ft diameter. New tray with Oldershaw-column efficiency of 55% (EOV = 0.55). Liquid flow path length from Fig. 6.16. Value of u/DE = 6 ft-1. From Example 6.4, = 0.68. Assumptions: Straight operating line and equilibrium line as stated in Example 6.4 Find: Efficiencies EMV and Eo . Analysis: Consider three cases: (a) complete mixing, (b) plug flow, and (c) partial mixing. (a) From Eq. (6-31), EMV = EOV = 0.55. From Eq. (6-37), log [1 + EMV ( - 1) ] log [1 + 0.55(0.68 - 1) ] = Eo = = 0.50 log log 0.68 (b) From Eq. (6-32), 1 1 EMV = eEOV - 1 = ( e(0.68)(0.55) - 1) = 0.67 0.68 ( ) From Eq. (6-37), log [1 + EMV ( - 1) ] log [1 + 0.67(0.68 - 1) ] = Eo = = 0.63 log log 0.68 (c) To use Fig. 6.16, we need an estimate of the liquid flow rate. For the 3 ft diameter column of Examples 6.1 and 6.4, liquid rate = 151.5 kmol/h of water. Assume the liquid rate is proportional to the column cross-sectional area. Thus, for the 11 ft diameter, the liquid rate = 151.5(11/3)2 = 2,037 kmol/h = 162 gpm. From Fig. 6.16, a single-pass tower is adequate. Assume liquid flow path = ZL = 7 ft. Eq. (6-34) is used to compute EMV, which requires the Peclet number, given by Eq. (6-36), Z u N Pe = L = 7(6) = 42 DE N From Eq. (6-35), = Pe 2 4 EOV 1+ N Pe 1/ 2 -1 = 42 2 1+ 4(0.68)(0.55) 42 1/ 2 - 1 = 0.37 From Eq. (6-34), EMV/EOV = 1.20. Therefore, EMV = 1.20(0.55) = 0.66 From Eq. (6-37), log [1 + EMV ( - 1) ] log [1 + 0.66(0.68 - 1)] Eo = = 0.62 = log log 0.68 From these results, we see that the column is operating closely to plug flow for the liquid. Exercise 6.18 Subject: Estimation of column diameter based on conditions at bottom tray of a reboiled stripper. Given: Vapor and liquid conditions at bottom tray given in Fig. 6.48. Valve trays with 24-inch spacing. Find: Column diameter for 80% of flooding. Analysis: Use entrainment flooding correlation of Fig. 6.24, where the abscissa is, LM L G FLV = VM G L From Fig. 6.48, V = 546.2 lbmol/h and L = 621.3 lbmol/h 1/ 2 (1) C i =1 Average molecular weights of the gas and liquid are computed from M = zi Mi , Component M y My x Mx Ethane 30.07 0.000006 0.0 0.000010 0.0 Propane 44.10 0.004817 0.2 0.001448 0.1 n-Butane 58.12 0.602573 35.0 0.391389 22.7 n-Pentane 72.15 0.325874 23.5 0.430599 31.1 n-Hexane 86.18 0.066730 5.8 0.176563 15.2 Total 64.5 69.1 Therefore, MV = 64.5 and ML = 69.1 Gas volumetric flow rate = QV = 6.192 ft3/s. Therefore, V = VMV /QV = (546.2)(64.5)/(6.192)(3,600) = 1.58 lb/ft3 Liquid volumetric flow rate = 171.1 gpm = (171.1)(60)/7.48 = 1,372 ft3/h Therefore, L = LML /QL = (621.3)(69.1)/1,372 = 31.3 lb/ft3 1/ 2 621.3(69.1) 158 . From Eq. (1), FLV = = 0.274 546.2(64.5) 31.3 Assume 24-inch tray spacing. From Fig. 6.24, CF = 0.24 ft/s From relations below Eq. (6-44), since FLV is between 0.1 and 1.0, F - 0.1 0.274 - 01 . Ad = 0.1 + LV = 01 + . = 0.119 A 9 9 From Perry's Handbook, surface tension = = 7 dyne/cm From relation below Eq. (6-42), FST = (/20)0.2 = (7/20)0.2 = 0.8. FF = 1.0 , FHA = 1.0 From Eq. (6-42), C = FST FF FHA CF = (0.8)(1.0)(1.0)(0.24) = 0.192 ft/s From Eq. (6-40), U f = C ( L - V ) / V 1/ 2 = 0.192 ( 31.3 - 1.58 ) /1.58 1/ 2 = 0.833 ft/s Exercise 6.18 (continued) From Eq. (6-44), 4VM V DT = fU f (1 - Ad / A ) V 1/ 2 4(546.2 / 3, 600)(64.5) = 0.80(0.833)(3.14)(1 - 0.119)(1.58) 1/ 2 = 3. 56 ft Exercise 6.19 Subject: Estimation of column diameter based on conditions at the top tray of an absorber. Given: Vapor and liquid conditions at the top tray. Valve trays with 24-inch spacing. Find: Flooding velocity and column diameter for 85% of flooding (f = 0.85) Analysis: Use entrainment flooding correlation of Fig. 6.24, where the abscissa is, FLV LM L V = VM V L 1/ 2 889(109) 1.924 = 530(26.6) 41.1 1/ 2 = 1.487 (1) For 24-inch tray spacing, from Fig. 6.24, CF = 0.082 ft/s From relations below Eq. (6-44), since FLV is < 0.1, Ad/A = 0.1 From relation below Eq. (6-42), FST = (/20)0.2 = (18.4/20)0.2 =0.98. FF = 0.75 , FHA = 1.0 From Eq. (6-42), C = FST FF FHA CF = (0.98)(0.75)(1.0)(0.082) = 0.0603 ft/s From Eq. (6-40), U f = C ( L - V ) / V 1/ 2 = 0.0603 ( 41.1 - 1.924 ) /1.924 1/ 2 1/ 2 = 0.272 ft/s 1/ 2 4VM V From Eq. (6-44), DT = fU f (1 - Ad / A ) V 4(530 / 3, 600)(26.6) = 0.85(0.272)(3.14)(1 - 0.1)(1.924) = 3. 53 ft Exercise 6.20 Subject: Hydraulic calculations for bottom tray of a sieve-tray column Given: Absorber of Exercise 6.16 with an absorbent flow rate of 40,000 gpm of nC8. Sieve trays on 24-inch spacing , with a weir height of 2.5 inches and 1/4-inch holes. Foaming factor = 0.80 and fraction flooding = 0.70. Find: (a) (b) (c) (d) (e) (f) Column diameter at bottom. Vapor pressure drop/tray. Whether weeping will occur. Entrainment rate. Fractional decrease in Murphree efficiency due to entrainment. Froth height in downcomer. Analysis: From Exercise 6.16, V = 800,000/359 = 2,230 lbmol/min of entering gas. Average MW of gas = 0.725(2.016) + 0.25(16.04) + 0.025(30.05) = 6.22 Assuming ideal gas law for gas at 400 psia and 100oF, V = PM/RT = (400)(6.22)/(10.73)(560) = 0.414 lb/ft3 For the liquid, neglect absorbed components. L = 43.9 lb/ft3 = 5.86 lb/gal MW of nC8 = 114. Therefore, L = 40,000(5.86)/114 = 2,056 lbmol/min (a) Use entrainment flooding correlation of Fig. 6.24, where the abscissa is, FLV LM L V = VM V L 1/ 2 = 2, 056(114) 0.414 2, 230(6.22) 43.9 1/ 2 = 1.64 (1) From Fig. 6.24, for 24-inch tray spacing, CF = 0.08. Because FLV > 1, Ad /A = 0.2. FHA = 1.0, FF = 0.8, and since = 20 dynes/cm, FST = 1.0. From Eq. (6-24), C = FSTFFFHACF = (1)(0.8)(1)(0.08) = 0.064 ft/s From Eq. (6-40), U f = C L - V / V 1/ 2 = 0.064 43.9 - 0.414 / 0.414 1/ 2 1/ 2 = 0.66 ft / s 1/ 2 4VM V From Eq. (6-44), DT = fU f (1 - Ad / A ) V 4(2, 230 / 60)(6.22) = 0.70(0.66)(3.14)(1 - 0.2)(0.414) = 44 ft This is a very large diameter. Would probably use two parallel columns of 31 ft each, but continue with 44 ft diameter. (b) Vapor pressure drop per tray is given by Eq. (6-49), ht = hd + hl + h (2) From the continuity equation, m = uA , hole velocity for 10% hole area = uo = (2,230/60)(6.22)/(3.14/4)(44)2(0.414)(0.1) = 3.68 ft/s and superficial velocity = 0.368 ft/s u 2 V 3.682 0.414 From Eq. (6-50), hd = 0186 o2 . = 0186 . = 0.045 in. of nC8 Co L 0.732 43.9 Analysis: (continued) Exercise 6.20 (continued) Active bubbling area for Ad /A = 0.2 is Aa = A - 2Ad = 0.6 A. So, Ua = 0.368/0.6 = 0.61 ft/s From Eq. (6-53), Ks = U a V L - V 1/ 2 0.414 = 0.61 43.9 - 0.414 1/ 2 = 0.06 ft/s From Eq. (6-52), e = exp(-4.257Ks0.91) = exp[-4.257(0.06)0.91] = 0.72 From Eq. (6-54), C = 0.362 + 0.317 exp( -35hw ) = 0.362 + 0.317 exp[ -3.5(2.5)] = 0.362 . Take weir length, Lw = 0.73DT = 0.73(44) = 32.1 ft = 385 in. qL From Eq. (6-51), hl = e hw + C Lw e 2/3 40,000 = 0.72 2.5 + 0.362 385 0.72 2/3 = 9.0 in. of nC8 From Eq. (6-55), with maximum bubble size of 1/4 inch = 0.00635 m, h = 6 / g L DBmax = 6(20 / 1000) / (9.8)(703)(0.00635) = 0.00274 m = 0.11 in. nC8 From Eq. (2), ht = hd + hl + h = 0.045 + 9.0 + 0.11 = 9.2 in. nC8 = 0.23 psi/tray. Excessive! (c) Apply criterion of Eq. (6-64). hd + h = 0.045 + 0.11 = 0.155 in. < hl = 9.0 in. Therefore, weeping will occur. (d) From Fig. 6.28, since FLV = 1.64, entrainment will be very low. (e) Because entrainment is very low, EMV will not decrease. (f) From Eqs. (6-70) and (6-72), hdf =(ht + hl + hda) / 2 (3) Area of downcomer opening = Ada = Lwha Take ha = 2 in., Ada = (32.1)(2/12) = 5.4 ft2 qL From Eq. (6-71), hda = 0.03 100 Ada 2 40, 000 = 0.03 100(5.4) 2 =164 in. Very excessive Exercise 6.21 Subject: Column performance for 40% of flooding. Given: Data in Examples 6.5, 6.6, and 6.7. Find: (a) Column diameter in Example 6.5 for f = 0.4. (b) Vapor pressure drop in Example 6.6 for f = 0.4. (c) Murphree vapor-point efficiency in Example 6.7 for f = 0.4 Analysis: (a) Example 6.5: In this example, a value of f = 0.80 was used, giving DT = 2.65 ft From Eq. (6-44), by ratioing values of f, 0.80 DT = 2.65 0.40 1/ 2 = 3.75 ft = 1.15 m (b) Example 6.6: In this example, a tower diameter of 1 m gives a vapor pressure drop = 0.093 psi/tray, with a vapor hole velocity of 47.9 ft/s, a weir length of 0.73 m, an active area vapor velocity of 5.99 ft/s, and Ks = 0.265 ft/s. Vapor hole velocity varies inversely with the square of the column diameter. Therefore, uo = 14.6(1/1.15)2 = 11.0 m/s = 36 ft/s From Eq. (6-50), hd is directly proportional to hole velocity squared. Therefore, hd = 1.56(36/47.9)2 = 0.88 in. of liquid Weir length is proportional to column diameter. So, Lw = 0.73(1.15/1) = 0.84 m = 0.33 in. Vapor velocity based on active area varies inversely with the square of column diameter. Therefore, Ua = 5.99(1/1.15)2 = 4.53 ft/s. From Eq. (6-53), Ks is proportional to Ua. Thus, Ks = 0.265(4.53/5.99) = 0.20 ft/s From Eq. (6-52), e = exp(-4.257Ks0.91) = exp[-4.257(0.20)0.91] = 0.37 The value of C in Eq. (6-51) remains at 0.362. From Eq. (6-51), qL hl = e hw + C Lw e 2/3 12.9 = 0.37 2 + 0.362 33 0.37 2/3 = 0.88 in. h = 0.36 in. (no change from Example 6.6) From Eq. (6-49), ht = hd + hl + h = 0,88 + 0.88 + 0.36 = 2.12 in. Tray pressure drop = htL = 2.12(0.0356) = 0.076 psi/tray (c) Example 6.7: In this example, a tower diameter of 1 m gives a EOV = 0.77. Must redo all calculations. DT = 1.15 m, A = 1.038 m2 = 10,380 cm2, Aa = 0.8(1.038) = 0.83 m2 = 8,300 cm2 Lw = 33 in. = 0.84 m, e = 0.37, hl = 0.88 in. = 2.24 cm Ua = 4.53 ft/s = 137 cm/s, Uf = 10.2 ft/s, f = Ua/Uf = 4.53/10.2 = 0.44 Analysis: (c) (continued) Exercise 6.21 (continued) qL = 812 cm3/s F = UaV0..5 = 1.37(1.92)0.5 = 1.90 (kg/m)0.5/s , From Eq. (6-64), t L = From Eq. (6-65), t G = hl Aa 2.24(10,380) = = 28.6 s qL 812 1 - e hl (1 - 0.37)2.24 = = 0.028 s eU a 0.37(137) 0 From Eq. (6-67), k L a = 78.8 DL.5 ( F + 0.425) = 78.8(181 10-5 )(190 + 0.425) = 0.78 s-1 . . From Eq. (6-66), kG a = 0 1,030 DV .5 f - 0.842 f 2 h 0.5 l = 1,030(7.86 10 -2 ) 0.5 0.44 - 0.842 0.44 2.24 0.5 2 = 53.4 s-1 From Eq. (6-63), N L = k L at L = 0.78(28.6) = 22.3 From Eq. (6-62), N G = k G at G = 53.4(0.28) = 15 . From Example 6.7, KV/L = 0.662 1 1 1 From Eq. (6-61), N OG = = = = 144 . 1 KV / L 1 0.662 0.667 + 0.030 + + NG NL 15 22.3 . Mass transfer is controlled by the vapor phase. From a rearrangement of Eq. (6-52), EOV = 1 - exp(- N OG ) = 1 - exp(-1.44) = 0.76 or 76%. Exercise 6.22 Subject: Sizing, hydraulics, and mass transfer for an acetone absorber Given: Acetone absorber of Fig. 6.1, using sieve trays with 10% hole area, 3/16-inch holes, and an 18-inch tray spacing. Foaming factor = 0.85. Find: (a) Column diameter for 75% of flooding. (b) Vapor pressure drop per tray. (c) Number of transfer units. (d) Number of overall transfer units. (e) Controlling resistance to mass transfer. (f) Murphree point vapor efficiency If 30 trays are adequate. Analysis: (a) Base column diameter on the top tray, where pressure is the lowest and gas flow rate is the highest. Use data in Fig. 6.1. L = 1,943 kmol/h, ML = 18, L = 1,000 kg/m3 V = 6.9 + 144.3 + 536.0 + 22.0 + 0.1 = 709.3 kmol/h, MV = 28.8, P = 90 kPa, T = 25oC From the ideal gas law, V = PM/RT = (90)(28.8)/(8.314)(298) = 1.05 kg/m3 Use entrainment flooding correlation of Fig. 6.24, where the abscissa is, LM L V 1,943(18) 105 . FLV = = = 0.056 VMV L 709.2(28.8) 1,000 From Fig. 6.24, for 18-inch tray spacing, CF = 0.28 ft/s. Because FLV < 1, Ad /A = 0.1. FHA = 1.0, FF = 0.85, and since = 70 dynes/cm, FST = (70/20)0.2 = 1.285 From Eq. (6-24), C = FSTFFFHACF = (1.285)(0.85)(1)(0.28) = 0.306 ft/s From Eq. (6-40), U f = C L - V / V From Eq. (6-44), 4VM V DT = fU f (1 - Ad / A ) V 1/ 2 1/ 2 1/ 2 1/ 2 = 0.306 1,000 - 105 / 105 . . 1/ 2 = 9.44 ft / s 4(709.2 / 3, 600)(28.8) = 0.75(9.44 / 3.28)(3.14)(1 - 0.1)(1.05) 1/ 2 = 1.89 m = 6.2 ft (b) Vapor pressure drop per tray is given by Eq. (6-49), ht = hd + hl + h (1) From the continuity equation, m = uA , hole velocity for 10% hole area = uo = (709.2/3,600)(28.8)/(3.14/4)(1.9)2(1.05)(0.1) = 19.0 m/s =62.3 ft/s Superficial velocity = (0.1)(19.0) = 1.90 m/s = 6.23 ft/s u 2 V 62.32 105 . From Eq. (6-50), hd = 0186 o2 . = 0186 . = 142 in. of liquid . 2 Co L 0.73 1,000 Active bubbling area for Ad /A = 0.1 is Aa = A - 2Ad = 0.8 A. So, Ua = 6.23/0.8 = 7.8 ft/s From Eq. (6-53), Ks = U a V L - V 1/ 2 1.05 = 7.8 1,000 - 105 . 1/ 2 = 0.253 ft/s From Eq. (6-52), e = exp(-4.257Ks0.91) = exp[-4.257(0.253)0.91] = 0.30 Assume a 2-inch weir height = hw Analysis: (b) (continued) . . From Eq. (6-54), C = 0.362 + 0.317 exp( -35hw ) = 0.362 + 0.317 exp[ -35(2.0)] = 0.362 Take weir length, Lw = 0.73DT = 0.73(6.2) = 4.52 ft = 54.3 in. Volumetric liquid rate = qL = m/L = (1,943/60)(18)/(1,000)(0.003785 m3/gal) = 154 gpm qL From Eq. (6-51), hl = e hw + C Lw e 2/3 Exercise 6.22 (continued) 154 = 0.30 2 + 0.362 54.3 0.30 2/3 = 1.09 in. of liquid From Eq. (6-55), with maximum bubble size of 3/16 inch = 0.00476 m, h = 6 / g L DBmax = 6(70 / 1000) / (9.8)(1,000)(0.00476) = 0.009 m = 0.35 in. of liquid From Eq. (2), ht = hd + hl + h =1.42 + 1.09 + 0.35 = 2.86 in. liquid = 0.103 psi/tray = 0.7 kPa/tray (c) DT = 1.89 m, A = 2.804 m2 = 28,040 cm2, Aa = 0.8(2.804) = 2.24 m2 = 22,400 cm2 e = 0.30, hl = 1.09 in. = 2.77 cm Ua = 7.8 ft/s = 238 cm/s, Uf =9.44 ft/s, f = Ua/Uf = 7.8/9.44 = 0.826 F = UaV0..5 = 2.38(1.05)0.5 = 2.44 (kg/m)0.5/s , qL = 154 gpm = 9,716 cm3/s hA 2.77(22,400) From Eq. (6-64), t L = l a = = 6.39 s qL 9,716 From Eq. (6-65), t G = 1 - e hl (1 - 0.30)2.77 = = 0.027 s eU a 0.30(238) From the Wilke-Chang Eq. (3-39), DL = 1.12 x 10-5 cm2/s 0 . From Eq. (6-67), k L a = 78.8 DL.5 ( F + 0.425) = 78.8(112 10-5 )(2.44 + 0.425) = 0.756 s-1 From Perry's Handbook, with a temperature correction based on Eq. (3-36), DV = 0.127 cm2/s From Eq. (6-66), kG a = 0 1,030 DV .5 f - 0.842 f 2 hl0.5 = 1,030(0127) 0.5 0.826 - 0.842 0.826 . 2.77 0.5 2 = 55.4 s-1 From Eq. (6-63), N L = k L at L = 0.756(6.39) = 4.83 From Eq. (6-62), N G = kG atG = 55.4(0.027) = 1.5 (d) From p. 271, for acetone, A = L/KV = 1.38. Therefore, KV/L = 1/1.38 = 0.725 1 1 1 From Eq. (6-61), N OG = = = = 122 . 1 KV / L 1 0.725 0.667 + 0150 . + + NG NL 15 4.83 . (e) Mass transfer is controlled by the vapor phase. (f) From a rearrangement of Eq. (6-56), E OV = 1 - exp( - N OG ) = 1 - exp( -122) = 0.705 or 70 % . Now, the separation requires 10 equilibrium stages. Assume the liquid on a tray is well mixed. Then, from Eq. (6-31), EMV = EOV = 0.70. From below Eq. (6-33), take =KV/L=0.725. log 1 + E MV ( - 1) log 1 + 0.70(0.725 - 1) From Eq. (6-37), Eo = = = 0.66 (worst case) log log 0.725 Therefore, at most, we need from Eq. (6-21), Na =Nt /Eo = 10/0.66 = 15. Therefore, 30 okay. Exercise 6.23 Subject: Design of a column to strip VOCs from water by air at 15 psia and 70oF. Given: Conditions in Example 6.2 except that entering flow rates are twice as much. Assumptions: Dilute system such that changes in vapor and liquid rates in the column are negligible. Sieve trays on 24-inch spacing, with 10 % hole area of 3/16-inch holes, FF = 0.9, FHA = 1.0 and f = 0.80. Find: (a) (b) (c) (d) (e) Number of equilibrium stages required. Column diameter for sieve trays. Vapor pressure drop per tray. Murphree vapor-point efficiency from Chan-Fair method. Number of actual trays. Analysis: (a) Need 3 equilibrium stages, as determined in Example 6.2 and which is unchanged when flow rates are doubled. (b) liquid rate = 2(500) = 1,000 gpm or 1,000(60)(8.33)/18.02 = 27,800 lbmol/h = L vapor rate = 2(3,400) = 6,800 scfm or 6,800(60)/379 = 1,077 lbmol/h = V L = 62.4 lb/ft3. From ideal gas law, V = PM/RT = (15)(29)/(10.73)(530) = 0.0765 lb/ft3 Use entrainment flooding correlation of Fig. 6.24, where the abscissa is, FLV LM L V = VMV L 1/ 2 27,800(18) 0.0765 = 1,077(29) 62.4 1/ 2 = 0.56 From Fig. 6.24, for 24-inch tray spacing, CF = 0.18 ft/s. Because 0.1 < FLV > 1, from below Eq. (6-44), Ad /A = 0.1 + (FLV - 0.1)/9 = 0.15 FHA = 1.0, FF = 0.9, and since = 80 dynes/cm, FST = (80/20)0.2 = 1.32 From Eq. (6-24), C = FSTFFFHACF = (1.32)(0.9)(1)(0.18) = 0.21 ft/s From Eq. (6-40), U f = C L - V / V From Eq. (6-44), for f = 0.80, 4VM V DT = fU f (1 - Ad / A ) V 1/ 2 1/ 2 = 0.21 62.4 - 0.0765 / 0.0765 1/ 2 = 6.0 ft / s 4(1, 077 / 3, 600)(29) = 0.80(6.0)(3.14)(1 - 0.15)(0.0765) 1/ 2 = 6.0 ft (1) (c) Vapor pressure drop per tray is given by Eq. (6-49), ht = hd + hl + h From the continuity equation, m = uA . Hole velocity for 10% hole area = uo = (1,077/3,600)(29)/(3.14/4)(6)2(0.0765)(0.1) = 40 ft/s Superficial velocity = (0.1)(40) = 4 ft/s Analysis: (c) (continued) From Eq. (6-50), hd = 0186 . Exercise 6.23 (continued) 2 uo V 402 0.0765 = 0186 . = 0.68 in. of liquid 2 2 Co L 0.73 62.4 Active bubbling area for Ad /A = 0.15 is Aa = A - 2Ad = 0.7 A. Therefore, Ua = 4/0.7 = 5.7 ft/s From Eq. (6-53), Ks = U a V L - V 1/ 2 0.0765 = 5.7 62.4 - 0.0765 1/ 2 = 0.20 ft/s From Eq. (6-52), e = exp(-4.257Ks0.91) = exp[-4.257(0.20)0.91] = 0.37 Assume a 2-inch weir height = hw . . From Eq. (6-54), C = 0.362 + 0.317 exp( -35hw ) = 0.362 + 0.317 exp[ -35(2.0)] = 0.362 Take weir length, Lw = 0.73DT = 0.73(6)(12) = 53 in. qL From Eq. (6-51), hl = e hw + C Lw e 2/3 1, 000 = 0.37 2 + 0.362 ( 53) 0.37 2/3 = 2. 58 in. of liquid From Eq. (6-55), with maximum bubble size of 3/16 inch = 0.00476 m, h = 6 / g L DBmax = 6(80 / 1,000) / (9.8)(1,000)(0.00476) = 0.0103 m = 0.40 in. of liquid From Eq. (2), ht = hd + hl + h =0.68 + 2.58 + 0.40 = 3.66 in. liquid = 0.13 psi/tray (d) DT = 6 ft, A = 28.3 ft2, Aa = 0.7(28.3) = 19.8 ft2 = 18,390 cm2 e = 0.37, hl = 2.58 in. = 6.55 cm, G = 0.0765 lb/ft3 = 1.23 kg/m3 Ua = 5.7 ft/s = 1.74 m/s , Uf =6.0 ft/s, f = Ua/Uf = 5.7/6.0 = 0.95 F = UaV0..5 =1.74(1.23)0.5 = 1.93 (kg/m)0.5/s , qL = 1,000 gpm = 63,100 cm3/s hA 6.55(18,390) From Eq. (6-64), t L = l a = = 19 s . qL 63,100 1 - e hl (1 - 0.37)6.55 = = 0.064 s eU a 0.37(174) From the Wilke-Chang Eq. (3-39), DL = 0.96 x 10-5 cm2/s 0 From Eq. (6-67), k L a = 78.8 DL.5 ( F + 0.425) = 78.8(0.96 10-5 )(193 + 0.425) = 0.58 s-1 . From Perry's Handbook, with T and P corrections based on Eq. (3-36), DV = 0.086 cm2/s From Eq. (6-65), t G = From Eq. (6-66), kG a = 0.5 1, 030 DV ( f - 0.842 f 2 ) h 0.5 l = 1, 030(0.086)0.5 0.95 - 0.842 ( 0.95 ) 6.55 0.5 2 = 2. 56 s-1 From Eq. (6-63), N L = k L at L = 0.58(1.9) = 1.1 Analysis: (continued) Exercise 6.23 (continued) From Eq. (6-62), N G = kG atG = 2.56(0.064) = 0.16 (d) From Example 6.2, for benzene, S = KV/L = 9.89 1 1 1 From Eq. (6-61), N OG = = = = 0.066 1 KV / L 1 9.89 6.25 + 8.99 + + NG NL 016 11 . . From a rearrangement of Eq. (6-52), EOV = 1 - exp(- N OG ) = 1 - exp(-0.066) = 0.064 or 6.4% (e) From Example 6.2, the separation requires 3 equilibrium stages. If we assume the liquid on a tray is well mixed. Then, from Eq. (6-31), EMV = EOV = 0.064. From Eq. (6-21), Na =Nt /Eo = 3/0.064 = 47 trays. If we assume plug flow of liquid on a tray, then, if we take, below Eq. (6.33), =KV/L=9.89, from Eq. (6-37), log 1 + E MV ( - 1) log 1 + 0.064(9.89 - 1) = = 0197 . log log 9.89 Then we need from Eq. (6-21), Na =Nt /Eo = 3/0.197 = 15 trays. Sufficient information is not given to establish the partial mixing prediction. Therefore, the number of trays required ranges widely from 15 to 47. Eo = Exercise 6.24 Subject: Absorption of SO2 from air into water in an existing packed column. Given: Feed gas flow rate of 0.062 kmol/s containing 1.6 mol% SO2. Absorbent is 2.2 kmol/s of pure water. Packed column is 1.5 m2 in cross sectional area and packed with No. 2 plastic super Intalox saddles to a 3.5-m height. Exit gas contains an SO2 mole fraction of 0.004. Operating pressure is 1 atm. At operating temperature, equilibrium curve for SO2 is y = Kx = 40x Assumptions: No stripping of water. No absorption of air. Find: (a) (b) (c) (d) L/Lmin NOG and Nt HOG and HETP KGa Analysis: Compute material balance. SO2-free inlet air rate = 0.062(1-0.016) = 0.061 kmol/s SO2 inlet rate in feed gas = 0.062(0.016) = 0.001 kmol/s = V' SO2 outlet rate in gas = 0.061(0.004/0.996) = 0.00025 kmol/s SO2 rate in outlet water = 0.001 - 0.00025 = 0.00075 kmol/s Fraction absorbed = 0.00075/0.001 = 0.75 or 75% (a) From Eq. (6-11), L'min = V'K(fraction absorbed) = 0.061(40)(0.75) = 1.83 kmol/s Therefore, L/L'min = 2.2/1.83 = 1.20 0.887 N +1 - 0.887 (b) Take A = L/KV = 2.2/[(40)(0.062)] = 0.887, From Eq. (6-13), 0.75 = 0.887 N +1 - 1 Solving, Nt = 4. For NOG, use Eq. (6-89) with yin = 0.016, yout = 0.004, xin =0.0, ln N OG = 0.887 - 1 0.016 1 + 0.887 0.004 0.887 = 3.78 (0.887 - 1) / 0.887 (c) Given height of packing = 3.5 m = lT From Eq. (6-73), HETP = lT /Nt = 3.5/4 = 0.875 m From Eq. (6-89), HOG = lT /NOG = 3.5/3.78 = 0.926 m (d) From Table 6.7, HOG = V K G aPS 3 = 0.045 kmol/s-m -atm Therefore, K G a = V H OG PS = 0.062 (0.926)(1)(1.5) Exercise 6.25 Subject: Operating data for absorption of SO2 from air into water in a packed column. Given: Column operates at 1 atm (760 torr) and 20oC. Solute-free water enters at 1,000 lb/h. Mole ratio of water to air is 25. Liquid leaves with 0.6 lb SO2/100 lb of solute-free water. Partial pressure of SO2 in exit gas is 23 torr (0.0303 atm). Equilibrium data are given as partial pressures of SO2 in air as a function of lb SO2 dissolved/ 100 lb H2O. Assumptions: No stripping of water. No absorption of air. Density of liquid taken as water. Find: (a) % of SO2 absorbed. (b) Concentration of SO2 in the liquid at the gas-liquid interface in lbmol/ft3 at a point where the bulk liquid concentration is 0.001 lbmol SO2/lbmol of water and: kL = 1.3 ft/h kp = 0.195 lbmol/h-ft2-atm Analysis: (a) Inlet water rate = 1,000/18.02 = 55.5 lbmol/h Inlet air rate = water rate/25 = 55.5/25 = 2.22 lbmol/h Partial pressure of air in exit gas = 760 - 23 = 737 torr By partial pressure ratio, SO2 flow rate in exit gas = 2.22(23/737) = 0.0693 lbmol/h Molecular weight of SO2 = 64.06 SO2 flow rate in exit liquid = 0.6(1,000/(100)(64.06) = 0.0937 lbmol/h By material balance, SO2 flow rate in entering air = 0.0693 + 0.0937 = 0.1630 lbmol/h Partial pressure of SO2 in entering gas = 0.1630/(0.1630 + 2.22) = 0.0690 atm % of SO2 absorbed = 0.0937/0.1630 x 100% = 57.5% (b) At the point, the rate of mass transfer for SO2 as a flux across the gas-liquid interface, can be written by the two-film theory as, kp(pb - pi) = kL(ci - cb) (1) As shown in Fig. 6.31, the equilibrium interface composition can be determined in terms of the ratio of mass-transfer coefficients. However, here, instead of compositions in mole fractions, the gas composition is in partial pressure and the liquid is in concentration. From Eq. (1), kL 1.3 p - pi = = 6.67 = b k p 0.195 ci - cb (2) At the point (column height location), bulk liquid concentration = 0.001 lbmol SO2/lbmol H2O and the flow rate is 0.001(55.5) = 0.0555 lbmol/h for SO2 in the liquid phase. The flow of SO2 in the gas phase at that location is obtained by a material balance around the top of the column: SO2 in entering liquid + SO2 in gas at the point = SO2 in liquid at the point + SO2 in exit gas Therefore, 0 + SO2 flow rate in gas at the point = 0.0693 + 0.0555 Analysis: (b) (continued) Exercise 6.25 (continued) SO2 flow rate in gas at the point = 0.1248 lbmol/h. Therefore, the partial pressure of SO2 in the bulk gas at the point = 0.1248/(0.1248 + 2.22) x 1 atm = 0.0532 atm = pb . The density of water is 62.4 lb/ft3 or 62.4/18.02 = 3.46 lbmol/ft3. Therefore, the concentration of SO2 in the bulk liquid at the point = 0.001(3.46) = 0.00346 lbmol/ft3 = cb . We now need an algebraic equilibrium relationship between pi and ci . Convert the given equilibrium data to partial pressures and concentrations in the vicinity of the values at the point: The SO2 concentration in the liquid is obtained from, c, lbmol SO 2 lb SO 2 = 3 100 lb H 2 O ft ci , lbmol SO2 / ft3 0.00292 0.00487 0.00682 0.00974 62.4 64.06 1 100 pi of SO2 , atm 0.01855 0.03421 0.05132 0.07763 (3) (4) lb SO2 / 100 lb H2O 0.30 0.50 0.70 1.00 Partial pressure SO2 , torr 14.1 26.0 39.0 59.0 These equilibrium data fit the curve, pi = 5.04536ci + 511.783ci2 - 21714.4ci3 From Eq. (2), Solving Eqs. (3) and (4), 0.0532 - pi = 6.67(ci - 0.00346) ci = 0.00550 lbmol/ft3 and pi = 0.0396 atm The same result can be obtained by constructing a plot of SO2 partial pressure versus SO2 concentration in the liquid, similar to Fig. 6.31. First, the operating line is drawn as a straight line connecting the column end points (0.0303, 0.0) and (0.0690, 0.00584), as (p , c). Then the equilibrium curve is drawn, using Eq. (3). Then the point (pb , cb) is marked on the operating line. A straight line is extended from this point, with a slope of (-kL/kp) = -6.67, to the point of intersection on the equilibrium line, giving the same result as above Exercise 6.26 Subject: Stripping of benzene from wastewater with air in a packed column. Given: Column operation at 2 atm (1,520 torr) and 25oC. Wastewater enters at 600 gpm containing 10 ppm by weight of benzene. Suggested air rate is 1,000 scfm (60oF and 1 atm). Exit water to contain just 0.005 ppm benzene. Vapor pressure of benzene = 95.2 torr at 25oC. Solubility of benzene in water = 0.180 g/100 g water at 25oC. Packing is 2-inch polypropylene Flexirings. Mass transfer coefficients are: kLa = 0.067 s-1 and kGa = 0.80 s-1 (both driving forces are in concentration units). Assumptions: No stripping of water. No absorption of air. Find: (a) rate. (b) (c) (d) (e) Minimum air stripping rate. Is it less than suggested? If not, use 1.4 times minimum Stripping factor NOG KGa in units of s-1 and mol/m3-s-kPa, and which phase controls mass transfer Volume of packing in m3 Analysis: Water flow rate = 600(8.33)(60)/18.02 = 16,640 lbmol/h For benzene, xin = (10/1,000,000)(18.02/78) = 2.31 x 10-6 xout = (0.005/1,000,000)(18.02/78) = 1.16 x 10-9 Fraction of benzene stripped = 1 - (2.31 x 10-6/1.16 x 10-9) = 0.9995 (a) System is very dilute with respect to benzene, therefore use Kremser equation to determine the minimum gas rate. From Eq. (6-12), Vmin = L (fraction stripped)/ K (1) For the K-value, use Eq. (4) in Table 2.3, a modified Raoult's law, K = LPs/P (2) To obtain the liquid-phase activity coefficient, use the given benzene solubility data. For liquidliquid equilibrium with respect to benzene, which is distributed between a water-rich phase (2) and a benzene-rich phase (1), using Eq. (2-30), ( 2 ) = x (1) / x ( 2 ) (1) . But if the benzene-rich phase is nearly pure benzene, (1) = 1 and x(1) = 1. Therefore, (2) = 1/x(2) 018(18.02) . From the given benzene solubility in water, x = = 4.2 10 -4 100(78) 1 Therefore from Eq. (3), L = = 2,380 4.2 10 - 4 (3) From Eq. (2), K = (2,380)(95.2)/1,520 = 149 From Eq. (1), Vmin = 16,640(0.9995)/150 = 111.6 lbmol/h At 60oF and 1 atm, there are 379 scf/lbmol. Therefore, Vmin = 111(379)/60 = 705 scfm. This is less than the 1,000 scfm suggested by the expert. Therefore, use the value suggested by the expert, rather than revise it. Analysis: (continued) Exercise 6.26 (continued) (b) Operating V = 111.6(1,000/705) = 158.3 lbmol/h From Eq. (5-51), S = KV/L = 149(158.3)/16,640 = 1.417 (a good value) (c) For NOG, use Eq. (6-93), with A = 1/S = 1/1.417 = 0.7057 Benzene stripped = Linxin(fraction stripped) = 16,640(2.31 x 10-6)(0.9995) = 0.0384 lbmol/h Therefore, yout = 0.0384/158.3 = 0.0002427. Also, xin = 2.31 x 10-6 and yin = 0.0. Therefore, ln N OG = = 14.2 (0.7057 - 1) / 0.7057 (d) Note that both kL and kG are given for concentration driving forces. Therefore, we write the rate of mass transfer of benzene as: r = kG acG ( yb - yi ) = k L acL ( xi - xb ) = K G acG ( yb - y* ) (4) 0.7057 - 1 0.7057 0.0 - 149(2.31 10-6 ) 1 + -6 0.0002427 - 149(2.31 10 ) 0.7057 where cG and cL are total gas and liquid concentrations, respectively, and y* is the vapor mole fraction in equilibrium with the bulk liquid mole fraction, as given by y* = Kxb . Solving Eq. (4) for the driving vapor and liquid phase driving forces, Adding Eqs. (5) and (6) to eliminate yi and solving for r and equating to the last term in Eq. (4), r= yb - y* 1 K + kG acG k L acL 1 KcG 1 + kG a k L acL = yb - y * 1 K G acG (7) Solving Eq. (7) for KGa, KG a = (8) The total gas concentration is obtained from the ideal gas law, Analysis: (d) continued) xercise 6.26 (continued) cG = P 2(101.3) = = 81.8 mol/m3 -3 RT 8.314 10 (298) L 106 g / m3 = = 55,500 mol/m3 ML 18.02 1 1 KG a = = = 0.22 s-1 1 150(81.8) 1.25 + 3.29 + 0.80 0.067(55,500) This is with a concentration driving force. cL = The liquid phase controls with a relative resistance of 3.29 compared to 1.25 for the gas phase. To obtain KGa with a partial pressure driving force, we write the rate of mass transfer as, r = K G acG ( yb - y* ) = ( K G ) p aP ( yb - y* ) (9) where KGa is with the concentration driving force = 0.22 s-1 and (KG)pa is with the partial pressure driving force. Solving Eq. (9), and applying the ideal gas law in the form, cG=P/RT, (K G ) p a = K G a cG K G a 0.22 = = = 0.089 mol/kPa-m3-s P RT (8.314 10-3 )(298) (e) From Table 6.7, with a partial pressure driving force, H OG = G (K G ) p aPS However, the cross sectional area of the tower, S, is not known. Therefore, can not compute HOG From Eq. (6-89), lT = HOGNOG . Therefore, can not compute the height because HOG is unknown. However, we can compute the packed volume. Packed volume = lTS = NOGHOGS = N OG G (149)(454) = (14.2) = 14.8 m3 (K G ) p aP 0.089(3, 600)(2)(101.3) Exercise 6.27 Subject: Absorption of GeCl4 from air into dilute caustic solution in an existing packed column. Given: Column operates at 25oC (77oF) and 1 atm. Gas enters at 23,850 kg/day containing 288 kg/day of GeCl4 and 540 kg/day of Cl2. Dissolved GeCl4 and Cl2 react with the caustic so that neither has a vapor pressure. Packed tower is 2-ft diameter with 10 ft of 1/2-inch ceramic Raschig ring packing of given characteristics. Liquid rate is to give 75% of flooding. Equation is given for estimating Kya. Assumptions: No stripping of water. No absorption of air. Find: (a) Entering dilute caustic flow rate. (b) Required packed height. Which controls, GeCl4 or Cl2? Is 10-ft height adequate? (c) % absorption of GeCl4 and of Cl2 for 10 ft of packing. If necessary, select an alternative packing. Analysis: (a) Determine entering dilute caustic flow rate from the 75% of flooding specification, using the flooding curve of Fig. 6.36(a), with the correction factors of Figs. 6.36(b,c). The air rate in the entering gas = 23,850 - 288 - 540 = 23,022 kg/day. MW of Cl2 = 71. MW of GeCl4 = 214.6. MW of air = 29. 23,022 288 540 Molar gas rate = V = + + = 803 kmol/day or 73.7 lbmol/h or 2189 lb/h 29 214.6 71 Average molecular weight of gas = 23,850/803 = 29.7 Tower cross sectional area for 2-ft diameter = 3.14(2)2/4 = 3.14 ft2 The continuity equation for flow through the tower based on the gas superficial velocity is, m = uoS. Therefore, the superficial velocity is, uo = m/S (1) 3 From the ideal gas law, = PM/RT = (1)(29.7)/(0.7302)(460+77) = 0.076 lb/ft Therefore, from Eq. (1), uo = (2189)/(3.14)(0.076) = 9,170 ft/h or 2.55 ft/s. From Eq. (6-102), noting that for the dilute caustic solution at 25oF, f{L} = 1 and f{L} = 0.98 for a liquid viscosity of 0.95 cP at 25oC, Y= 2 uo FP V g H 2O f { L }f { L }= (2.55) 2 (580) 0.076 (1.0)(0.98)=0.14 32.2 62.4 Therefore, Yflood = 0.14/(0.75)2 = 0.25 V L 1/ 2 (uo)flood = uo/0.75 =2.55/0.75 = 3.40 ft/s. LM L From Fig. 6.36, VM V LM L = X (GM G ) L G 1/ 2 = 0.009 62.4 = 0.009(2,191) 0.076 1/ 2 = 565 lb/h or 0.071 kg/s of entering liquid Analysis: (b) Assume that GeCl4 is the controlling species, with mass transfer controlled in the gas phase. From Table 6.7, column height = lT = HGNG where, HG = V/kyaS (1) and, N G = (1 - y ) LM dy (1 - y )( y - yI ) (2) Exercise 6.27 (continued) At the interface, yI = 0 because it is given that dissolved GeCl4 has no vapor pressure In entering gas, y = (288/214.6)/803 = 0.00162. Therefore (1 - y)bottom = (1 - 0.00162) = 0.9984 In exiting gas, y = 0.01 of y in entering gas = 0.0000162. Therefore (1 - y)top = 1.0000 Therefore, (1 - y)LM = 0.9992. Also, on the average, (1 - y) is approximately 0.9992. Therefore, Eq. (2) for NG simplifies to, NG = dy 0.00162 = ln = 4.61 y 0.0000162 0.0000162 -0.36 0.00162 V DPV Equation given for ky is, k y = 1.195 S (1 - o ) ( NSc ) -2 / 3 (3) V = 803/24(3,600) = 0.0093 kmol/s S = 3.14 ft2 or 0.292 m2 V' = (V/S)MG = (0.0093/0.292)(29.7) = 0.944 kg/s-m2 DP = 0.01774 m At 25oC, from Perry's Handbook, for air, = 0.018 cP or 1.8 x 10-5 kg/m-s = - hL = 0.63 hL = 0.03591(L')0.331 (4) L' =(LML) /S= 0.071/0.292 = 0.243 kg/s-m2 From Eq. (4), hL = 0.0359(0.243)0.331 = 0.0225 Therefore, o = 0.63 - 0.0225 = 0.608 NSc = /DGeCl4 (5) 3 3 From above, of the gas = 0.076 lb/ft = 1.21 kg/m DGeCl4 = 6 x 10-6 m2/s From Eq. (5), NSc = (1.8 x 10-5)/(1.21)(6 x 10-6) = 2.48 0.0093 (0.01774)(0.944) From Eq. (3), k y = 1195 . 0.292 (18 10 -5 ) 1 - 0.608 . Equation given for interfacial area is, Equation given for exponent, -0.36 2.48 -2 / 3 = 126 10 -3 kmol/s-m2 . 14.69(808 V / 1/ 2 ) n a= 0.111 ( L ) (6) n = 0.01114 L' + 0.148 = 0.01114(0.243) + 0.148 = 0.151 0.151 Therefore, from Eq. (6), a = (808)(0.944) 14.69 (121)1/ 2 . 0.243 0.111 = 46.2 m2/m3 Exercise 6.27 (continued) Analysis: (b) (continued) Kya = kya = (1.26 x 10-3)(46.2) = 0.058 kmol/s-m3 From Eq. (1), HG = HOG = (V/S)/kya = (0.0093/0.292)/0.058 = 0.55 m or 1.8 ft Packed height = HGNG = 0.55(4.61) = 2.53 m or 8.3 ft. Therefore, a 10 foot height is sufficient. Now check the assumption that GeCl4 controls. Because gas is dilute in Cl2 and 99% of it is absorbed, NG is the same as for GeCl4. Note that in Eq. (3), k y is proportional to ( N Sc ) -2 / 3 or Di2 / 3 . Since the given gas diffusivity of Cl2 is about twice that of GeCl4, the mass transfer coefficient for Cl2 is higher and the corresponding HG is lower. Therefore, the assumption that GeCl4 controls is correct. (c) If a packed height of 10 ft rather than the computed 8.3 ft is used, then for GeCl4 , NG = lT /HG = 10/1.8 = 5.55 rather than 4.61. y From above, N G = ln in = 5.55 yout y Solving, in = 258 Therefore, yout = 0.00162/258 = 0.0000063 yout % absorption of GeCl4 = (1 - 0.0000063/0.00162) = 0.996 or 99.6% compared to 99% specified For Cl2, from above, HG = HG of GeCl4 (DGeCl4/DCl2)2/3 = 1.8(0.000006/0.000013)2/3 = 1.08 Therefore, NG = 10/1.08 = 9.26 = = ln yin yout Solving, yin = 10,500 yout % absorption of Cl2 = (1 - 1/10,500) = 0.9999 or 99.99% compared to 99% specified Exercise 6.28 Subject: Packed tower diameter and pressure drop for conditions of Exercise 6.26. Given: Suggested tower diameter of 0.80 m and pressure drop of 500 N/m2-m of packed height (0.612 in. H2O/ft). Find: (a) (b) (c) (d) Fraction of flooding if FP = 24 ft2/ft3. Pressure drop at flooding. Pressure drop at operating conditions using GPDC chart. Pressure drop at operating conditions using Billet-Schultes correlation. Analysis: (a) Using data from Exercise 6.26 and Fig. 6.36, LML = 600(8.33) = 5,000 lb/min VMV = (1,000/379)(29) = 76.5 lb/min From ideal gas law, V = PM/RT = (2)(29)/(0.7302)(460+77) = 0.148 lb/ft3 L =62.3 lb/ft3 LM L V 5, 000 0.148 From Fig. 6.36, X = = = 3.19 VM V L 76.5 62.3 From Fig. 6.36, at flooding, Y = 0.0053 For the actual operation, by the continuity equation, m = uoS Area for flow = S = D2/4 = (3.14)(0.8/0.3048)2/4 = 5.41 ft2 Superficial velocity = uo = m/S = (76.5/60)/(5.41)(0.148) = 1.59 ft/s From Eq. (6-102), noting that for the dilute aqueous solution at 25oF, f{L} = 1 and f{L} = 0.98 for a liquid viscosity of 0.95 cP at 25oC, 2 uo FP V Y= g H 2O 1/ 2 1/ 2 (1.59) 2 (24) 0.148 (1.0)(0.98)=0.0044 f { L }f { L }= 32.2 62.3 1/ 2 1/ 2 u Y 0.0044 Fraction of flooding = o = = uflood Yflood 0.0053 This is too high. Should increase tower diameter. = 0.91 or 91% (b) From Eq. (6-104), Pflood = 0.115 FP0.7 = 0.115(24)0.7 = 1.06 in. H2O/ft of packed height (c) From Fig. 6.36, P at design = 1.50 in. H2O/ ft of packed height (seems high) (d) From Table 6.8, for 2-inch (50-mm) plastic Pall rings, the packing characteristics are: a = 111.1 m2/m3 or 33.9 ft2/ft3 , = 0.919 m3/m3 , Ch = 0.593 , Cp = 0.698, Cs = 2.816 The pressure drop per unit height of packed bed is given by Eq. (6-115), P = Po - hL 3/ 2 exp 1/ 2 13300 N 3/ 2 ( FrL ) a (1) Exercise 6.28 (continued) Analysis: (d) (continued) where, hL is given by (6-97) and is in m2/m3 and N FrL is given by (6-99). From Eq. (6-110), Po a u2 1 = o 3 V V lT 2 KW (2) 1- 1 - 0.919 =6 = 0.0143 ft a 33.9 1 2 1 DP 2 1 0.0143 From Eq. (6-111), = 1+ = 1+ = 1045 . KW 3 1 - DT 3 1 - 0.919 2.62 Therefore, KW = 0.957 . From Eq. (6-112), DP = 6 From Perry's Handbook, for air at 25oC = 0.018 cP, and uo = uV = 1.59 ft/s From Eq. (6-114), N ReV = uV DP V (1.59)(0.0143)(0.148)(0.957) KW = = 3, 287 (1 - ) V (1 - 0.919 ) [(0.018)(0.000672)] 64 1.8 64 1.8 + 0.08 = 0.698 + = 0.671 3, 287 3, 287 0.08 N ReV N ReV From Eq. (6-113), o = C p From Eq. (2), with the introduction of gc because of the use of American Engineering units, Po a u2 1 33.9 1.592 (0.148) 1 = o 3 V V = 0.671 = 0.164 lbf/ft3 or 0.00114 psi/ft 3 lT 2 g c KW 0.919 2(32.2) 0.957 From Eq. (6-97), hL = 12 N FrL N Re L 1/ 3 ah a 2/3 (3) Superficial liquid velocity = uL = m/SL = (5,000/60)/(5.41)(62.4) = 0.247 ft/s From Perry's Handbook, liquid viscosity = 0.95 cP 2 u L a (0.247) 2 (33.9) From Eq. (6-99), N FrL = = = 0.064 g 32.2 u (0.247)(62.4) From Eq. (6-98), N Re L = L L = = 712 a L (33.9)(0.95)(0.000672) a 0. 0 From Eq. (6-101), h = 0.85Ch N Re25 N Fr.1 = 0.85(0.593)(712) 0.25 (0.247) 0.1 = 2.27 L L a Exercise 6.28 (continued) Analysis: (d) (continued) 0.064 From Eq. (3), hL = 12 712 From Eq. (1), P = Po - hL 3/ 2 1/ 3 2.27 2 / 3 = 0.177 m3/m3 or ft3/ft3 exp 1/ 2 13300 N 3/ 2 ( FrL ) a P 0.919 = ( 0.00114 ) lT 0.919 - 0.103 3/ 2 exp 13300 1/ 2 0.064 ) = 0.0241 psi/ft or 0.669 in. H2O/ft 3/ 2 ( 111.1 This is only 10% higher than the suggested value of 0.612 in. H2O/ft Exercise 6.29 Subject: Mass-transfer coefficients for conditions of Exercise 6.26. Given: Suggested mass-transfer coefficients in Exercise 6.26. Find: Mass-transfer coefficients kLa and kGa from Billet-Schultes correlation. Analysis: Coefficient kLa is obtained from a rearrangement of the left-most part of Eq. (6-131), uL kL a = (1) H L ( aPh / a ) with HL from Eq. (6-132) and (aPh /a) from (6-136). From Table 6.8, CL = 1.239 , = 0.919, and a = 111.1 m2/m3. From the results of Exercise 6.28, come the following parameters and properties: hL = 0.177 ft2/ft3, uL = 0.247 ft/s, L = 62.4 lb/ft3, L = 0.95 cP Take the surface tension, = 0.005 lbf/ft or 0.005(32.2) = 0.161 lb/s2 To compute (aPh /a), we need for the liquid phase, the Reynolds, Weber, and Froude numbers from equations (6-138), (6-139), and (6-140), respectively, based on the packing hydraulic diameter from (6-137). d h = packing hydraulic diameter = 4 From Exercise 6.28, Reynolds number = N ReL ,h = 0.919 =4 = 0.0331 m = 0.108 ft a 111.1 uL d h L (0.247)(0.108)(62.4) = = 2, 607 [(0.95)(0.000672)] L 2 Weber number = N WeL ,h ( 0.247 ) (62.4)(0.108) = 2.55 u 2 d = L L h = 0.161 ( 0.247 ) = 0.0175 u2 = L = gd h 32.2(0.108) 2 Froude number = N FrL , h From (6-136), aPh -1/ 2 = 1.5 ( ad h ) N Re L ,h a ( ) (N -0.2 -1/ 2 We L , h ) (N ) 0.75 FrL , h -0.45 = 1.5 [ (111.1)(0.0331) ] ( 2607 ) ( 2.55 ) ( 0.0175 ) -0.2 0.75 -0.45 = 2.02 Exercise 6.29 (continued) Analysis: (continued) From Exercise 6.23, correcting for temperature, DL = 9.6 x 10-6(298/294) = 9.73 x 10-6 cm2/s or 1.05 x 10-8 ft2/s From (6-132), using American Engineering Units, 1 1 HL = CL 12 1/ 6 4hL DL au L 1/ 6 1/ 2 uL a a a ph 1/ 2 1 1 = 1.239 12 4(0.177)(0.919) (1.05 10-8 )(33.9)(0.247) 0.247 1 = 5.23 ft 33.9 2.02 From (1), k L a = uL 0.247 = = 0.0234 s-1 H L ( aPh / a ) 5.23 ( 2.02 ) This is 35% of the suggested value of 0.067 s-1 Coefficient kGa is obtained from the gas-phase analog of (1), kG a = uV H G ( aPh / a ) (2) with HG obtained from (6-133), using dimensionless groups from (6-134) and (6-135). From Exercise 6.28, uV = 1.59 ft/s, V = 0.148 lb/ft3, V = 0.018 cP = 1.21 x 10-5 lb/ft-s From Eq. (6-134), N ReV = uV V (1.59)(0.148) = = 574 (33.9)(1.21 10-5 ) aV From Exercise 6.23, the diffusivity of benzene in air = DV = 0.086 cm2/s at 70oF and 15 psia. Using Eq. (3-36) to correct for temperature and pressure, DV = 0.086 15 (2)(14.7) 77 + 460 70 + 460 1.75 = 0.045 cm2/s = 4.84 x 10-5 ft2/s From Eq. (6-135), NScV = V 1.21 10-5 = = 1.69 V DV (0.148)(4.84 10 -5 ) Exercise 6.29 (continued) Analysis: (continued) From Eq. (6-133) 1 1/ 2 4 HG = ( - hL ) CV a4 1/ 2 1/ 2 (N ) (N ) -3/ 4 -1/ 3 ReV ScV uV DV 1.59 a aPh = 1 4(0.919) ( 0.919 - 0.177 )1/ 2 0.368 33.94 From (2), ( 574 )-3 / 4 (1.69 )-1/ 3 4.84 10-5 1 = 0.455 ft 2.02 kG a = 1.59 = 1.73 s-1 0.455 ( 2.02 ) This is approximately twice the suggested value of 0.80 s-1. Exercise 6.30 Subject: Absorption of NH3 from air into water in a packed tower. Given: Column operation at 68oF and 1 atm. Inlet gas mass velocity = 240 lb/h-ft2. Inlet water mass velocity = 2,400 lb/h-ft2. Tower packed with 1.5-inch ceramic Berl saddles. Henry's law applies for NH3 with p (atm) = 2.7x. Assumptions: No stripping of water. No absorption of air. Dilute system. Find: (a) Packed height for 90% absorption of NH3. (b) Minimum water mass velocity for 98% absorption of NH3. (c) KGa, pressure drop, maximum liquid rate, KLa, packed height, column diameter, HOG, and NOG for 1.5-inch ceramic Hiflow rings. Analysis: (a) Assuming Dalton's law, the equilibrium equation for 1 atm is, p = yP = 2.7x atm or with P = 1 atm, y = Kx = 2.7x Because the system is dilute, Eq. (6-89) can be used to compute NOG . Then the packed height is given by Eq. (6-85), lT = NOGHOG (2) The absorption factor from Eq. (6-15) is, A = L/KV = (2,400/18)/2.7(240/29) = 6.0 Entering and exiting ammonia mole fractions are: yin = 0.02, yout = 0.002, xin = 0.0 The ammonia mole fraction in the exiting liquid is obtained by an overall ammonia balance: 240 2,400 240 2,400 yin + xin = yout + xout 29 18 29 18 Solving Eq. (3), xout = 0.00112. From Eq. (6-89), N OG = ln ( A - 1) / A ( yin - Kxin ) / ( yout - Kxin ) + (1 / A) ( A - 1) / A = ln (6 - 1) / 6 0.02 / 0.002 + (1 / 6) (6 - 1) / 6 = 2.57 (1) (3) To obtain an estimate of HOG, the correlations of Billet and Schultes in Chapter 6 or literature data can be used. The latter is found on page 16-38 of Perry's Chemical Engineers' Handbook, 6th edition or on page 663 of "Equilibrium-Stage Separation Operations in Chemical Engineering" by Henley and Seader, as shown below. Analysis: (a) (continued) Exercise 6.30 (continued) Air-Ammonia-Water system with 1.5-inch Berl Saddles, AE units From the above figure, HOG = 0.85 ft. From Eq. (2), lT = 2.57(0.85) = 2.2 ft (b) From Eq. (6-11), (L/S)min = (V/S) K (fraction absorbed) = 240(2.7)(0.98) = 635 lb/h-ft2 (c) Consider the differences between 1.5-inch ceramic Berl saddles (Saddles) and 1.5-inch ceramic Hiflow rings (Rings). In Table 6.8, data are given for the rings, but not for the saddles of the 1.5-inch size. To make the comparison, take the near-1 inch size. Then, the following parameters apply: Packings Saddles Rings FP 110 a 260 261.2 0.68 0.779 Ch 0.62 1.167 Cp 0.628 CL 1.246 1.744 CV 0.387 0.465 Analysis: (c) (continued) Exercise 6.30 (continued) Note that for the NH3-air-water system, the above plot shows that HOG is approximately equal to HG. Therefore, the liquid-phase resistance is small and KL a need not be considered. Consider KGa and HOG: In the above table, values of a are essentially the same. From Eqs. (6-136) and (6-140), a / a for Rings 0.779 = = 1.07 aPh/a is proportional to /a. Therefore, Ph aPh / a for Saddles 0.68 From Eq. (6-133), if we ignore holdup, hL, in the term ( - hL), and differences in a, then, 1 4 a H G is proportional to CV aPh 1/2 1/ 2 1 (0.779) 4 H G for Rings Therefore, = 0.465 (1.07 ) = 1.53 1 H G for Saddles 4 (0.680) 0.387 This ratio should be about the same for HOG . K G a for Rings 1 From Table 6.7, KG is inversely proportional to HG. Therefore, = = 0.653 K G a for Saddles 1.53 Consider pressure drop: The pressure drop is given by Eq. (6-106). If we ignore holdup, hL, in the term ( - hL), and differences in a, and combine Eqs. (6-99) and (6-110) to (6-115), then, Cp P is approximately proportional to 3 However, Table 6.8 does not contain Cp for Berl saddles, so no comparison can be made. Consider Column NOG and height: The value of NOG is independent on packing material. Therefore, the height is shorter for the saddle packing because the HOG is smaller. Consider column diameter: The column diameter is related to the flooding velocity, as given in Fig. 6.36(a). For a given value X, the value of Y is fixed and the product (uo)2FP is a constant. The packing factor for the ceramic Hiflow rings is not given in Table 6.8. However, if a value is extrapolated from data for larger rings, the packing factor for the Rings is much less than for the Saddles. Therefore, the flooding velocity for the Rings will be higher and the column diameter smaller. Consider maximum liquid rate: For a fixed gas velocity, the value of Y in Fig. 6.36(a) is smaller for the Rings because the packing factor is less. Therefore, as shown in the same figure, the value of X is greater for flooding, which gives a larger maximum liquid rate. Exercise 6.31 Subject: Absorption of CO2 from air into a dilute caustic solution with a packed column. Given: 5,000 ft3/min at 60oF and 1 atm, containing 3 mol% CO2. Recovery of 97% of CO2. Equilibrium curve is Y = KX = 1.75X (i.e. mole ratios) at column operating conditions. Assumptions: No stripping of water. No absorption of air. Caustic solution has properties of water. Initial estimate of column diameter is 30 inches. Try 2-inch Intalox saddles packing. Find: (a) (b) (c) (d) (e) (f) (g) Minimum caustic solution-to-air molar flow-rate ratio. Maximum possible concentration of CO2 in caustic solution. Nt at L/V = 1.4 times minimum. Caustic solution rate. Pressure drop per ft of column packed height. NOG Packed height for KGa = 2.5 lbmol/h-ft3-atm. Analysis: Gas flow rate = 5,000(60)/379 = 792 lbmol/h. Therefore, air in entering gas = V' = 0.97(792) = 768 lbmol/h The CO2 in the entering gas = 0.03(792) = 24 lbmol/h The CO2 in the exiting gas = 0.03(24) = 0.72 lbmol/h The CO2 absorbed in the leaving liquid = 24 - 0.72 = 23.28 lbmol/h The mole ratios are Yin = 24/768 = 0.03125, Yout = 0.72/768 = 0.0009375, Xin = 0.0 (a) From Eq. (6-11), L'min = V'K(fraction absorbed) = 768(1.75)(0.97) = 1,304 lbmol/h (b) The maximum possible CO2 concentration in the caustic solution occurs at infinite stages with the minimum liquid rate. The leaving liquid is in equilibrium with the entering gas. Therefore, Xout = Yin/K = 0.03125/1.75 = 0.0179 mol CO2 / mol caustic solution. (c) Let L/V = L'/V' . Then, L = 1.4 (1,304) = 1,826 lbmol/h and L/G = 1,826/768 = 2.38. For this liquid flow rate, a material balance for CO2 is, 768(0.03125 - 0.0009375) = 23.28 = 1,826(Xout - 0.0). Solving, Xout = 0.01275 In the Y-X plot on the next page, the equilibrium line, Y = 1.75X, and the straight operating line, passing through the column Y-X end points {Yin = 0.03125, Xout = 0.01275} and {Yout = 0.0009375, Xin = 0.0} are shown, with the equilibrium stages stepped off, giving Nt = 7.3. (d) From Part (c), caustic rate = 1,826 lbmol/h. (e) Assume the use of 2-inch of ceramic Intalox saddles. From Table 6.8, FP = 40 ft2/ft3. Use Fig. 6.36(a). Assume for the dilute caustic solution, f{L} = 1.0 and f{L} = 1.0. Take MV = 29, ML = 18, L = 62.4 lb/ft3, V = 29/379 = 0.0765 lb/ft3 Using conditions at the bottom of the column, Exercise 6.31 (conditions) Analysis: (d) and (e) (continued) LM L V X= VM V L 1/ 2 (1,849)(18) 0.0765 = (792)(29) 62.4 1/ 2 = 0.051 From Fig. 6.36(a), at flooding, Y = 0.18 Now compute Y for the suggested column diameter of 30 inches = 2.5 ft. From the continuity equation, uo = m/SV = [(792)(29)/3600]/[3.14(2.5)2/4](0.0765) = 17 ft/s From Table 6.8, the packing factor, FP = 40 ft2/ft3. 2 uo FP V 17 2 (40) 0.0765 Y= = = 0.440 , which is much greater than 0.18 g H 2O 32.2 62.4 Therefore, a diameter of 30 inches places the operation badly into the flooding region and a pressure drop calculation is meaningless. A larger diameter is necessary. Exercise 6.31 (continued) (f) From Eq. (6-95), with A = L/KV = (1,826)/1.75(768) = 1.36, NOG = Nt = A ln(1/A)/(1- A) = 7.3(1.36) ln (1/1.36)/(1 - 1.36) = 8.5 (g) From Table 6.7, HOG = V/KGaPS, with V = 792 lbmol/h, KGa = 2.5 lbmol/h-ft3-atm, P = 1 atm For column cross sectional area, S, assume 50% of flooding. At flooding, with Y = 0.18, by ratio with the Y of 0.440 for a 2.5 ft diameter, need uo at flooding = 17(0.18/0.44)0.5 = 10.9 ft/s At 50% of flooding, u = 5.45 ft/s. By ratio, column diameter = 2.5(17/5.45)0.5 = 4.4 ft Column cross sectional area = S = 3.14(4.4)2/4 = 15.2 ft2. Therefore, HOG = (792)/(2.5)(1)(15.2) = 20.8 ft, which is very high. From Eq. (6-89), lT = HOG NOG = 20.8(8.5) = 177 ft. Given value of KGa seems very low. Exercise 6.32 Subject: Absorption of NH3 from air into water with a packed column. Given: Conditions at a point in the column where P = 101.3 kPa, T = 20oC, 10 mol% NH3 in bulk gas, 1 wt NH3 in bulk liquid, NH3 partial pressure of 2.26 kPa at the interface, and ammonia absorption rate of 0.05 kmol/h-m2. Find: (a) X, Y, Yi , Xi , X*, Y*, KY, KX, kY, and kX in Fig. 6.49. (b) % of mass transfer resistance in each phase. 1 1 H = + (c) Verify that KY kY k X Analysis: (a) At the point, as shown in Fig. 6.49, the mole ratios in the bulk are X = (1/17)/(99/18) = 0.0107 and Y = 0.1/0.9 = 0.111. At the interface, Yi = 2.26/(101.3 - 2.26) = 0.0229. Then, using the equilibrium curve in Fig. 6.49, Xi = 0.028 The values of the mole fractions in equilibrium with the bulk values are obtained from the equilibrium curve in Fig. 6.49. Thus, in equilibrium with Y = 0.111 is X* = 0.114. In equilibrium with X = 0.0107 is Y* = 0.007 The given NH3 mass-transfer flux can be written with the following combinations of masstransfer coefficients and driving forces, using Table 6.7 as a guide: N = 0.05 = KY Y - Y * = K X X * - X = kY Y - Yi = k X X i - X (1) From Eq. (1), KY = 0.05/(Y - Y*) = 0.05/(0.111 - 0.007) = 0.480 kmol/h-ft2-mole ratio From Eq. (1), KX = 0.05/(X* - X) = 0.05/(0.114 - 0.0107) = 0.484 kmol/h-ft2-mole ratio From Eq. (1), kY = 0.05/(Y - Yi) = 0.05/(0.111 - 0.0229) = 0.568 kmol/h-ft2-mole ratio From Eq. (1), kX = 0.05/(Xi - X) = 0.05/(0.028 - 0.0107) = 2.89 kmol/h-ft2-mole ratio (b) To determine the relative mass-transfer resistances in each phase, Eq. (1) can be rearranged as in Eq. (6-78): kX (Y - Yi ) 2.89 =- =- = -5.08 kY ( X - Xi ) 0.568 Thus, the lowest coefficient and the highest driving force is that associated with the gas phase. From the ratio of 5.08, the largest mass-transfer resistance is in the gas phase. That resistance is about five times the resistance in the liquid phase, with 84% in the gas phase, 16% in the liquid. (c) From a mole ratio form of Eq. (6-80), 1 1 1 Yi - Y * = + (2) KY kY k X X i - X Therefore, Therefore, H = Yi - Y * 0.0229 - 0.007 = = 0.919 0.028 - 0.0107 Xi - X 1 H 1 0.919 1 1 + = + = 2.08 compared to = = 2.08 kY k X 0.568 2.89 KY 0.480 Exercise 6.33 Subject: Stripping of ammonia from an aqueous solution with air in a packed column. Given: Aqueous solution of 20 wt% NH3. Exit liquid to contain no more than 1 wt% NH3. Exit gas to be 1,000 ft3/h of a 10 mol% NH3 in air. Equilibrium data in Fig. 6.49. Temperature = 20oC = 68oF. Assumptions: No stripping of water. No absorption of air. Operation at 1 atm. Find: Volume of packing for KGa = 4 lbmol/h-ft3-atm. Analysis: First, determine the molar flow rates of components in the exit gas. From the ideal gas law applied to the exiting gas at 1 atm and 68oF, = P / RT = 1 / (0.7302)(68 + 460) = 0.00259 lbmol/ft3 Flow rate of exit gas = 1,000(0.00259) = 2.59 lbmol/h Gas contains 0.10(2.59) = 0.259 lbmol/h of NH3 and 0.90(2.59) = 2.331 lbmol/h of air. Next, determine the mole fractions in the inlet and exit liquid streams. Take as a basis, 100 lb of inlet liquid. Component H2O NH3 Total Inlet liquid: lb lbmol 80 4.44 20 1.18 100 5.62 mole fraction 0.79 0.21 1.00 Outlet liquid: lb lbmol 80.000 4.4400 0.808 0.0475 80.808 4.4875 mole fraction 0.9894 0.0106 1.0000 Determine the water flow rate, L', in and out by an ammonia balance: 0.21 0.0106 (1) L + 0 = L + 0.259 0.79 0.9894 Solving Eq. (1), L' = 1.015 lbmol/h of H2O. NH3 in entering liquid = (0.21/0.79)1.015 = 0.270 lbmol/h NH3 in leaving liquid = 0.270 - 0.259 = 0.011 lbmol/h Because of the concentrated solution, take into account the bulk-flow effect. For a partial pressure driving force, can write the rate of mass transfer, V for volume, as, nNH 3 = 0.259 = KG aV * pNH 3 - pNH 3 ( yair ) avg 0.259( yair ) avg LM (2) Rearranging Eq. (1), V= p * NH 3 - pNH 3 LM KG a (3) Exercise 6.33 (continued) Analysis: (continued) First consider the bulk-flow effect. At the top of the column, in the bulk exiting air, yair = 0.90 To obtain the corresponding value of y in the vapor in equilibrium with the entering liquid at the top of the column, use Fig. 6.49. In entering bulk liquid, moles NH3/moleH2O =1.18/4.44=0.266. From Fig. 6.49, by a rough extrapolation, moles NH3/mole H2O in equilibrium gas = 0.266. Therefore, y*air = 1 - 0.266/1.266 = 1 - 0.210 = 0.79. The average yair = (0.90 + 0.79)/2 = 0.845. At the bottom of the column, in the bulk entering air, yair = 1.00 To obtain the corresponding value of y* in the vapor in equilibrium with the leaving liquid at the column bottom, use Fig. 6.49. In leaving bulk liquid, moles NH3/mol H2O = 0.0475/4.4 = 0.0108. From Fig. 6.49, moles NH3/mole H2O in equilibrium gas = 0.007 = mole fraction of NH3 Therefore, y*air = 1 - 0.007 = 0.993. The average yair = (1.000 + 0.993)/2 = 0.997 The (yair)avg = (0.845 + 0.997)/2 = 0.921 Now, determine the log mean partial pressure driving force for ammonia, using Dalton's law with the above mole fractions. At the top of the column, for ammonia, (p*- p)top = P (y*- y)top = (1.0)(0.210 - 0.10) = 0.11 atm At the column bottom, for ammonia, (p*- p)bottom = P (y*- y)bottom = (1.0)(0.007 - 0.0) = 0.007 atm Therefore, for ammonia, (p*- p)LM = 0.037 atm From Eq. (3), packed volume = V = 0.259(0.921) = 1.61 ft3 (0.037)(4) Exercise 6.34 Subject: Absorption of NH3 from air into water in a packed column. Given: Entering gas is 2,000 lb/h of air (dry basis) containing NH3 with a partial pressure of 12 torr and saturated with water vapor at 68oF and 1 atm. Equilibrium data in Fig. 6.49. Absorb 99.6% of the NH3. Assumptions: No absorption of air. No stripping of water. Find: (a) Minimum water rate. (b) Column diameter and height for water rate of 2 times minimum and 50% of flooding gas velocity if 38-mm ceramic Berl saddles are used. (c) Same as Part (b) for 50-mm Pall rings. (d) Packing recommendation. Analysis: First compute the ammonia material balance. Flow rate of NH3-free air = 2,000/29 = 69 lbmol/h Flow rate of NH3 in the entering gas = 12(69)/(760-12) = 1.107 lbmol/h Flow rate of NH3 in the exiting liquid = 0.996(1.107) = 1.1026 lbmol/h Flow rate of NH3 in the leaving gas = 1.107 - 1.1026 = 0.0044 lbmol/h (a) Operation in dilute region, where we can apply Henry's law. From Fig. 6.49, using results from Exercise 6.33, when x=0.0108, y=0.007. Therefore, the equilibrium equation is y = 0.65x. Therefore, K = 0.65. Entering gas rate = V = 69 + 1.107 = 70.1 lbmol/h From Eq. (6-11), Lmin = VK(fraction absorbed) = 45.4 lbmol/h (b) L = 2Lmin = 2(45.4) = 90.8 lbmol/h Compute column diameter for 50% of flooding, using Fig. 6.36. Take properties of gas and liquid as those of air and water, respectively. From the ideal gas law, G = PM/RT = (1)(29)/(0.7302)(528) = 0.0752 lb/ft3 LM L V In Fig. 6.36(a), X = VM V L 1/ 2 (45.4)(18) 0.0752 = (70.1)(29) 62.4 1/ 2 = 0.014 2 uo FP V = 0.22 g H 2O Take f{L} and f{L} =1.0. From Fig. 6.36(a), at flooding, Y = (1) For 38-mm Berl saddles, the packing factor, FP, is not included in Table 6.8. However, it can be found in Perry's Chemical Engineers' Handbook, which gives FP = 65 ft2/ft3. Yg H2 O (0.22)(32.2) 62.4 2 Solving Eq. (1), uo = = = 90.4 FP V 65 0.0752 Therefore, uo = 9.5 ft/s. The fraction of flooding = f = 0.5. From Eq. (6-103), the column diameter is, 4VM V DT = fuo V 1/ 2 4(70.1/ 3, 600)(29) = (0.5)(9.5)(3.14)(0.0752) 1/ 2 = 1.42 ft and S=DT2/4 =3.14(1.42)2/4=1.58 ft2 Analysis: (b) (continued) For column height, use Eq. (6-89), because mass-transfer is controlled by the gas phase. For NOG , use Eq. (6-88) with xin = 0. Absorption factor = A = L/KV = (90.8)/(0.65)(70.1) = 2 From Eq. (6-93), using yin = 1.107/70.1 = 0.0158, yout = 0.0044/69 = 0.0000638 N OG = ln ( A - 1) / A ( yin ) / ( yout ) + (1 / A) ( A - 1) / A = ln (2 - 1) / 2 0.0158 / 0.0000638 + (1 / 2) (2 - 1) / 2 = 9.65 Exercise 6.34 (continued) For HOG, literature data can be used, as found on page 16-38 of Perry's Chemical Engineers' Handbook, 6th edition or on page 663 of "Equilibrium-Stage Separation Operations in Chemical Engineering" by Henley and Seader, as shown below. To use this plot, L" = 90.9(18)/1.58 = 1,040 lb/h-ft2 and V"=G" = 70.1(29)/1.58 = 1,290 lb/h-ft2 Then, HOG = 1.9 ft and column height = lT = NOGHOG = (9.65)(1.9) = 18.3 ft. Air-Ammonia-Water system with 1.5-inch Berl Saddles, AE units Exercise 6.34 (continued) Analysis: (c) (continued) (c) Assume the 50-mm Pall rings are metal. From Table 6.8, FP = 27 ft2/ft3. 2 Solving Eq. (1), uo = Yg H 2O (0.22)(32.2) 62.4 = = 235 FP V 27 0.0752 Therefore, uo = 15.3 ft/s. The fraction of flooding = f = 0.5. From Eq. (6-103), the column diameter is, 4VM V DT = fuo V 1/ 2 4(70.1/ 3, 600)(29) = (0.5)(15.3)(3.14)(0.0752) 1/ 2 = 1.12 ft and S = 3.14(1.12)2/4 = 0.985 ft2 The value of NOG = 9.65, the same as in Part (b) For HOG , can use Figs. 6.42 and 6.43 for 50 mm Pall rings. Gas capacity factor = F = uG(V)0.5 = f uo(V)0.5 = (0.5)(15.3)(0.0752)0.5 = 2.1 lb1/2-s-1-ft-1/2 or 5.12 kg1/2-s-1-m-1/2 By the continuity equation, liquid load = QL/S = uL = m/SL = (90.8/3600)(18)/(0.985)(62.4) =0.0074 ft/s or 0.00225 m/s or 2.25 x 10-3 m3/m2-s In Fig. 6.42, from the slope of the correlating line, kGa is proportional to F0.7 In Fig. 6.43, from the slope of the correlating line, kGa is proportional to uL0.45 Using the value of kGa = 3.1 s-1 at uL = 2.25 x 10-3 m3/m2-s and F = 1.16, then the two plots are represented by the equation, kGa = 43.4 F0.7 uL0.45 For our case, kGa = 43.4 (5.120.7)(0.002250.45) = 8.8 s-1 Note that this kGa is in concentration units for the driving force. Therefore, from Table 6.7 by analogy to the liquid phase case, HG = VMV/kGaSV = (70.1/3,600)(29)/(8.8)(0.985)(0.0752) = 0.87 ft The column height is lT = 9.65(0.87) = 8.4 ft. Exercise 6.34 (continued) Analysis: (d) Now compare the two packings on the basis of flooding and mass transfer: Packing 38-mm ceramic Berl saddles 50-mm metal Pall rings Column diameter, ft 1.42 1.12 Column height, ft 18.3 8.4 Clearly the 50-mm metal Pall rings are superior in performance. However, the selection would have to be made on the basis of cost. Because of the much smaller column and volume of packing, it is likely that the 50-mm Pall rings would be favored. From page 14-59 of Perry's Chemical Engineers' Handbook, 7th edition (1997), the cost of 38mm ceramic Berl saddles is approximately $21/ft3. The cost of 50-mm Pall rings depends on whether they can be made of carbon steel or must be made of stainless steel, with the former $19.9/ft3 and the latter $99.0/ft3. So the material of construction is a very important factor. Exercise 6.35 Subject: Absorption of Cl2 from air with water in a packed column. Given: 100 kmol/h of feed gas containing 20 mol% Cl2. Column operation at 20oC and 1 atm. Table of y-x data for Cl2. Exit gas to contain 1 mol% Cl2. Assumptions: No stripping of water. No absorption of air. Find: (a) Minimum water rate in kg/h. (b) NOG for a water rate of 2 times minimum. Analysis: Solve this problem using mole ratios. Feed gas is 80 kmol/h of air = G' , and 20 kmol/h of Cl2 with Yin = 20/80 = 0.25. Exit gas has Yout = 1/99 = 0.0101. Xin = 0.0. Highest value of Y in the table of equilibrium data is 0.06/0.94 = 0.0638. Therefore, data for higher values of Y are needed. Can obtain this from Perry's Handbook, page 3-102 in the 6th edition and page 2-126 in the 7th edition. The expanded table of equilibrium data with conversion to mole ratios includes the following values added from Perry's Handbook, using the conversions: y = pCl2, torr/760 torr x = [(grams Cl2/L)/71]/[1,000/18 + (grams Cl2/L/71)] pCl2, torr 100 150 200 grams Cl2/L 1.773 2.27 2.74 y 0.132 0.197 0.263 x 0.000450 0.000576 0.000695 These equilibrium data and those given in the exercise are now converted to mole ratios using, Y = y/(1 - y) and X = x/(1 - x) y 0.006 0.012 0.024 0.040 0.060 0.132 0.197 0.263 x 0.000100 0.000150 0.000200 0.000250 0.000300 0.000450 0.000576 0.000695 Y 0.00604 0.01215 0.02459 0.04167 0.06383 0.15207 0.24533 0.35685 X 0.000100 0.000150 0.000200 0.000250 0.000300 0.000450 0.000576 0.000695 (a) To determine the minimum water rate, it is best to do this graphically because of the high degree of curvature of the equilibrium curve. The pinch region of infinite stages at minimum absorbent rate occurs at the bottom of the column, such that the leaving liquid is in equilibrium with the entering gas, with a Y = 0.25. With mole-ratio coordinates, the operating line is straight and passes through the following point at the top of the column {X = 0.0, Y = 0.0101}. Exercise 6.35 (continued) Analysis: (a) (continued) The point at the bottom of the column must pass through the point {X* in equilibrium with Y = 0.25}. This is shown in the following plot, which is similar to Fig. 6.9. From the solute material balance, Eq. (6-6), the slope of the operating line is given by L'/V'. From the above plot, Xbottom = 0.000581, and, thus, the slope of that line is, L'/V' = (0.25 - 0.0101)/(0.000581 - 0.0) = 413. Since V' = 80 kmol/h, L' = 80(413) = 33,000 kmol/h. Alternatively, since the Cl2 absorption rate = V'(Ybottom - Ytop) = 80(0.25 - 0.0101) = 19.19 kmol/h, then L' = 19.19/ Xbottom = 19.19/0.000581 = 33,000 kmol/h. Take the result as 33,000 kmol/h or 33,000 (18) = 594,000 kg/h. (b) Twice the minimum water rate = 2(33,000) = 66,000 kmol/h. This gives a mole ratio of chlorine in the leaving water of Xbottom =19.19/66,000 = 0.000291 moles Cl2 / mole water. This operating line is shown in the plot below. To obtain NOG , using mole ratios, the integral in Eq. (6-138) or Table 6.7 can be used, Analysis: (b) (continued) Exercise 6.35 (continued) N OG = dY (Y - Y *) Y = 0.0101 Y = 0.25 (1) 1 (Y - Y *) as a function of Y, where for a given value of X, the value of Y is obtained from the operating line and Y* is obtained from the equilibrium curve, with values as follows. The values of X are taken from the above table, where the values of Y are the corresponding values of Y*. For given values of X, values of Y for the operating line are obtained from the following equation of the straight operating line passing through the points { X = 0.0, Y = 0.0101} and { X = 0.000291, Y = 0.25}, L 66, 000 Y = 0.0101 + X = 0.0101 + X = 0.0101 + 825 X V 80 Eq. (1) can be solved graphically or numerically. Both require a table of values of Note that several values of X are added by interpolation to increase the accuracy. Analysis: (b) (continued) X 0.000000 0.0000125 0.000025 0.000050 0.000075 0.000100 0.000150 0.000200 0.000250 0.000291 Y* 0.00000 0.00075 0.00150 0.00301 0.00452 0.00604 0.01215 0.02459 0.04167 0.06030 Exercise 6.35 (continued) Y 0.0101 0.0204 0.0307 0.0514 0.0720 0.0926 0.1339 0.1751 0.2164 0.2500 1/(Y-Y*) 99.01 50.86 34.22 20.69 14.82 11.55 8.22 6.64 5.72 5.27 Y 0.0103 0.0103 0.0207 0.0206 0.0206 0.0413 0.0412 0.0413 0.0336 Avg 1/(Y-Y*) 74.9 42.5 27.5 17.8 13.2 9.88 7.43 6.18 5.50 Area 0.772 0.438 0.569 0.367 0.272 0.408 0.306 0.256 0.185 The following is a plot of 1 versus Y. The area under the curve is the value of the (Y - Y *) integral for NOG . In the above table, the trapezoidal method is used to break the curve into 9 segments of Y each, varying in width from 0.0103 to 0.0413. For each segment (width), the average value of 1/(Y-Y*) is listed and the product = Area = Y times 1/(Y-Y*) is given in the last column. The sum of these areas is NOG = 3.573. Exercise 6.36 Subject: Absorption of ammonia from air into water with a packed column. Given: Operating conditions in Example 6.15. Entering gas is 100 lbmol/h containing 40 mol% NH3. Absorbent is 488 lbmol/h of water. Absorb 95% of the NH3. Equilibrium data in Fig. 6.44. Column operates at 1 atm and 25oC. Assumptions: No stripping of water. No absorption of air. Liquid has properties of water. Find: Column diameter and height if packing is 1.5-inch metal Pall rings, using conditions at bottom of tower. Analysis: From Example 6.15, the conditions at the bottom of the tower are: Entering gas: 60 lbmol/h of air and 40 lbmol/h of NH3 Exiting liquid: 488 lbmol/h of water and 38 lbmol/h of NH3 Operating temperature is 298 K (536oF) and pressure is 1 atm To compute column diameter use Fig. 6.36 to obtain flooding gas velocity, assume operation at 50% of flooding to obtain operating gas velocity, and use continuity equation to obtain column cross sectional area and column diameter. Average gas molecular weight = [(60)(29) + (40)(17)]/100 = 24.2 Gas density = V = PM/RT = (1)(24.2)/(0.7302)(536) = 0.0618 lb/ft3 From the abscissa of Fig. 6.36(a), LM L V X= VM V L From Fig. 6.36(a), Y = 0.105 = 1/ 2 = (526)(17.9) 0.0618 (100)(24.2) 62.4 1/ 2 = 0.123 (1) 2 uo FP V f { L } f { L } g H2 O For 1.5-inch (35 mm) metal Pall rings, from Table 6.8, FP = 40 ft2/ft3 At 25oC, from Figs. 6.36(b) and (c), f{L} = 1.0 and f{L} = 0.98 From a rearrangement of Eq. (1), 2 uo = Yg H2 O FP V 1 (0.105)(32.2) 62.4 1 = = 87 ft2/s2 f { L } f { L } 40 0.0618 (1)(0.98) Then, uo = 9.3 ft/s Take the operating superficial velocity = uV = fuo = 0.5 uo = 0.5(9.3) = 4.65 ft/s From Eq. (6-99), 4GM G DT = fuo G 1/ 2 4(100 / 3, 600)(24.2) = (0.5)(9.3)(3.14)(0.0618) 1/ 2 = 1.73 ft Analysis: (continued) Exercise 6.36 (continued) (2) From Eq. (6-127), column height = lT = HOGNOG From Example 6.15, NOG = 3.46 and from above, column cross section = S = D2/4 = 2.35 ft2. Mass-transfer data for 2-inch metal Pall rings for the absorption of NH3 from air into water are presented in Figs. 6.42 and 6.43. In the solutions to Exercise 6.34, these data are fitted to the (3) equation, kGa = 43.4 F0.7 uL0.45 , s-1 0.5 0.5 1/2 -1 where the gas capacity factor = F = uV(V) = f uo (V) in kg -s -m-1/2 and from the continuity equation, uL = m/SL in m/s or m3/m2-s For our case, F = (0.5)(9.3)(0.0618)0.5 = 1.16 lb1/2-s-1-ft-1/2 or 2.83 kg1/2-s-1-m-1/2 and uL = [(526)(17.9)/3,600)]/(2.35)(62.4) = 0.018 ft/s = 0.00547 m3/m2-s Substitution of these values into Eq. (3) gives, kGa = 43.4(2.83)0.7(0.00547)0.45 = 8.6 s-1 Note that this kGa is in concentration units for the driving force. Therefore, from Table 6.7 by analogy to the liquid phase case, HOG = HG = VMV/kGaSV = (100/3,600)(24.2)/(8.6)(2.35)(0.0618) = 0.54 ft Now, we must correct this value to 1.5-inch metal Pall rings. From Table 6.8, using interpolation when necessary, we have the following characteristics: Packing 1.5-inch Pall rings 2.0-inch Pall rings a, m2/m3 139.4 112.6 0.965 0.951 Ch 0.644 0.784 CV 0.373 0.410 From Eqs. (6-136) to (6-140), aPh/a is proportional to 1/2/a . Therefore, aPh / a for 1.5-inch Rings 0.965 = aPh / a for 2.0-inch Rings 0.951 1/ 2 112.6 = 0.814 139.4 From Eq. (6-133), if we ignore holdup in the term ( - hL), then, 1 a H G is proportional to 1.25 CV a aPh 1 0.965 H G for 1.5-inch Rings 0.341 139.41.25 1 Therefore, = = 1.15 1 0.951 H G for 2.0-inch Rings 0.814 0.410 112.61.25 This ratio should be about the same for HOG . Therefore, for the 1.5-inch rings, HOG = 0.54(1.15) = 0.62 ft. From Eq. (2), column height = 3.46(0.62) = 2.1 ft. Exercise 6.37 Subject: Absorption of acetone from air with water in a packed column. Given: Column operates at 60oF and 1 atm. Entering air is 50 ft3/min at 60oF and 1 atm, containing 3 mol% acetone. 97% of the acetone is to be absorbed. Maximum allowable gas superficial velocity is 2.4 ft/s. Equilibrium equation is Y = 1.75 X (1) Assumptions: No stripping of water. No absorption of air. Find: (a) (b) (c) (d) (e) (f) Minimum water-to-air ratio. Maximum acetone concentration possible in the water. Nt for absorbent flow rate equal to 1.4 times the minimum. Number of overall gas transfer units. Height of packing for Kya = 12.0 lbmol/h-ft3-mole ratio difference. Height of packing as a function of molar flow ratio for constant V and HTU. Analysis: First compute material balance. Total gas in = 50(60)/379 = 7.916 lbmol/h Acetone in entering gas = 0.03(7.916) = 0.237 lbmol/h Air in entering gas = 7.916 - 0.237 = 7.679 lbmol/h Acetone in exiting gas = 0.03(0.237) = 0.007 lbmol/h Acetone in entering liquid = 0.0 lbmol/h Acetone in exiting liquid = 0.237 - 0.007 = 0.230 lbmol/h Xin = 0.0 Yin = 0.03/0.97 = 0.03093 Yout = 0.007/7.679 = 0.000912 (a) For minimum water rate, the exiting liquid will be in equilibrium with the entering air. Therefore, using the equilibrium equation, Eq. (1), Xout = Yin /1.75 = 0.03093/1.75 = 0.01767 Therefore, water rate out and in = 0.230/0.01767 = 13.02 lbmol/h Minimum water-to-air molar ratio = 13.02/7.679 = 1.70 (b) Maximum possible acetone concentration in the liquid is at min. water rate, X = 0.01767 This corresponds to an acetone mole fraction of 0.01767/(1 + 0.01767) = 0.01736. (c) Use Lin = L' = 1.4 Lmin = 1.4(13.02) = 18.23 lbmol/h Using mole ratios with no acetone in entering liquid, and K' = constant = 1.75, can apply Eq. (5-29), which was derived for extraction but applies to absorption using: Yout = Yin 1 N n=0 where E = En L 18.23 = = 1.356 K G (175)(7.679) . Therefore, Yout 0.000912 = = 0.0295 = Yin 0.03093 1 N n =0 Solving, N = 7.5 equilibrium stages n 1356 . Analysis: (c) (continued) Could also solve this using Kremser's equation or by a graphical construction like Fig. 6.12. (d) Because operating line and equilibrium line are straight, use Eq. (6-95) with A = E = 1.356 N OG = N t ln(1/ A) ln(1/1.356) = 7.5 = 8 .7 (1 - A) / A (1 - 1.356) /1.356 Exercise 6.37 (continued) (e) From Eq. (6-141), lT = HOGNOG From Table 6.7, H OG = V 7.679 0.64 = = , where S = cross sectional area of tower in ft2 KY aS (12.0) S S The allowable gas superficial velocity = uV = 2.4 ft/s. Take this at the bottom of the tower where gas rate is the highest. From the continuity equation, Q = uV S = gas volumetric rate Therefore, S = Q/uV = (50/60)/2.4 = 0.347 ft2. Therefore, HOG = 0.64/0.347 = 1.84 ft Packing height = (1.84)(8.7) = 16 ft. (f) To obtain packed height as a function of L'/V' with V' and HOG constant, the only change will be the value of NOG . Use Eq. (6-93), which for xin = 0 and yin/yout = 0.03/0.000911 = 32.9, is: ln 32.9 N OG = A -1 1 + A A A -1 A (2) where for a dilute system, we can take A = L'/K'V' = (L'/V')/K' = (L'/V')/1.75 The results obtained from Eq. (2) are as follows: L'/V' 1.75 2.0 2.5 3.0 5.0 10.0 A 1.00 1.143 1.429 1.714 2.857 5.714 1/A 1.00 0.875 0.700 0.583 0.350 0.175 NO'G 12.9 7.86 6.38 4.74 4.01 Height, ft 23.6 14.5 11.7 8.72 7.38 Exercise 6.38 Subject: Absorption of NH3 from air into water with a packed column. Given: Column operating at 68oF and 1 atm. Gas enters at 2,000 ft3/min at 68oF and 1 atm, containing 6 mol% NH3 in air. Entering water rate is twice the minimum. Absorb 99% of the NH3. Gas velocity at 50% of flooding velocity. Assumptions: No stripping of water. No absorption of air. Find: Tower diameter and packed height. Analysis: Entering gas rate is 2,000 60 + 460 60 = 312 lbmol /h 379 68 + 460 Compute material balance. NH3 in entering gas = 0.06(312) = 18.72 lbmol/h NH3 in exiting gas = 0.01(18.72) = 0.19 lbmol/h Air in entering and exiting gas = V' = 312 - 18.72 = 293.28 lbmol/h NH3 in exiting water = 18.72 - 0.19 = 18.53 lbmol/h Therefore, Xin = 0.0 Yin = 18.72/293.28 = 0.06383 Yout = 0.19/293.28 = 0.000648 At the minimum water rate, the exiting liquid is in equilibrium with the entering gas. From Fig. 6.49, X*out = 0.072 for Yin = 0.06383. Therefore, L'min = 18.53/0.072 = 257 lbmol/h Operating liquid rate = 2 times minimum = 2(257) = 514 lbmol/h Now, compute column diameter using Fig. 6.36. Corrections for liquid density and viscosity are neglected. From the ideal gas law on the entering gas, with an average molecular weight = 28.3, V =PM/RT = (1)(28.3)/(0.7302)(68+460) = 0.0734 lb/ft3 From Table 6.8, packing factor for 1-inch metal Pall rings = 56 ft2/ft3. LM L V The abscissa in Fig. 6.36 is, X = VM V L 2 From Fig. 6.36, Y = 0.18. Therefore, uo = 1/ 2 (514 + 18.5)(18) 0.0734 = (312)(28.3) 62.4 1/ 2 = 0.0373 Yg H 2O (0.18)(32.2) 62.4 = = 88 FP V 56 0.0734 Therefore, flooding velocity = uo = 9.4 ft/s For 50% of flooding, f = 0.5 4GM G From Eq. (6-103), DT = fuo G 1/ 2 4(312 / 3, 600)(28.3) = (0.5)(9.4)(3.14)(0.0734) 1/ 2 = 3.0 ft. Exercise 6.38 (continued) Analysis: (continued) Assume gas phase controls mass transfer rate. Then, from Eq. (6-89), lT = HOGNOG From Eq. (6-93), for xin = 0, yin = 0.06, and yout = 0.00065, ln 92.3 N OG = A -1 1 + A A A -1 A (1) For the absorption factor, A = L/KV. Take L = 514 lbmol/h. Take V = 312 lbmol/h. From Fig. 6.49, K = 0.82. Therefore, A = (514)/(0.82)((312) = 2.0 From Eq. (1), 2 -1 1 ln 92.3 + 2 2 = 7.7 N OG = 2 -1 2 Mass-transfer data for 2-inch metal Pall rings for the absorption of NH3 from air into water are presented in Figs. 6.42 and 6.43. In the solutions to Exercise 6.34, these data are fitted to the equation, (3) kGa = 43.4 F0.7 uL0.45 , s-1 where the gas capacity factor = F = uV(V)0.5 = f uo (V)0.5 in kg1/2-s-1-m-1/2 and from the continuity equation, uL = m/SL in m/s or m3/m2-s. Here, S = 3.14(3)2/4 = 7.07 ft2 For our case, F = (0.5)(9.4)(0.0734)0.5 = 1.27 lb1/2-s-1-ft-1/2 or 3.10 kg1/2-s-1-m-1/2 and uL = [(532.5)(18)/3,600)]/(7.07)(62.4) =0.006 ft/s = 0.00182 m3/m2-s Substitution of these values into Eq. (3) gives, kGa = 43.4(3.10)0.7(0.00182)0.45 = 5.6 s-1 Note that this kGa is in concentration units for the driving force. Therefore, from Table 6.7 by analogy to the liquid phase case, HOG = HG = VMV/kGaSV = (312/3,600)(28.3)/(5.6)(7.07)(0.0734) = 0.84 ft Now, we must correct this value to 1-inch metal Pall rings. From Table 6.8, using interpolation when necessary, we have the following characteristics: Packing 1.0-inch Pall rings 2.0-inch Pall rings a, m2/m3 223.5 112.6 0.954 0.951 Ch 0.719 0.784 CV 0.336 0.410 Analysis: (continued) Exercise 6.38 (continued) From Eqs. (6-136) to (6-140), aPh/a is proportional to 1/2/a . Therefore, aPh / a for 1.0-inch Rings 0.954 = aPh / a for 2.0-inch Rings 0.951 1/ 2 112.6 = 0.505 223.5 From Eq. (6-133), if we ignore holdup in the term ( - hL), then, H G is proportional to 1 a 1.25 CV a aPh Therefore, 1 0.954 H G for 1.0-inch Rings 0.336 223.51.25 = 1 0.951 H G for 2.0-inch Rings 0.410 112.61.25 1 = 1.03 0.505 This ratio should be about the same for HOG . Therefore, for the 1.5-inch rings, HOG = 0.84(1.03) = 0.87 ft. From Eq. (2), column height = 7.7 (0.87) = 6.7 ft. Exercise 6.39 Subject: Absorption of SO2 from air by water in a packed column. Given: Mass velocity of entering gas = 6.90 lbmol/h-ft2 on SO2-free basis, containing 80 mol% air and 20 mol% SO2. Water enters at a mass velocity of 364 lbmol/h-ft2. Gas exits with 0.5 mol% SO2. Tower operates at 30oC and 2 atm. Equilibrium y-x equation given for SO2. Tower is packed with Montz B1-200 metal structured packing. Assumptions: No stripping of water. No absorption of air. Find: (a) Molar material balance operating line. (b) NOG Analysis: (a) A material balance for the solute, SO2, from the top of the column to an intermediate point gives: Yn +1V + X in L = YoutV + X n L Solving Eq. (1), the operating line is given by, Y= L L X + Yout - X in V V (2) (1) From the given data, L'/V' = 364/6.90 = 52.75. Xin = 0.0 and Yout = 0.005/0.995 = 0.005025 Rearrange Eq. (2), letting Y = y/(1-y) and X = x/(1-x), which is approximately x because of the dilute condition, x V y V y X= x= - Yout = 0.01896 - 0.000095 (3) L 1 - y L 1- x 1- y which is very close to the given operating line. (b) From Eq. (6-138), N OG = dY Y -Y * Y = 0.00503 Y = 0.25 (4) 1 (Y - Y *) as a function of Y, where for a given value of Y, the value of X is obtained from Eq. (3). Then, x is computed from X. Then, y = y* is computed from the given equilibrium equation, and Y* is 1 then obtained from the equilibrium curve. Below is a plot of versus Y. The area under (Y - Y *) the curve is the value of the integral for NOG . In the table below, the trapezoidal method is used Eq. (4) can be solved graphically or numerically. Both require a table of values of Analysis: (b) (continued) to break the curve into segments of Y. each, varying in width from 0.0103 to 0.0413. For each segment (width), the average value of 1/(Y-Y*) is used to compute the product = Area = Y times 1/(Y-Y*), which is given in the last column. The sum of these areas is NOG = 5.16. Exercise 6.39 (continued) Exercise 7.1 Subject: Comparison of absorption, distillation, and stripping Find: Differences between absorption and distillation, and stripping and distillation Analysis: Absorption Feed is gas Second phase (absorbent) is added Operation can be adiabatic Single-section cascade Can be almost isothermal Can not separate a closeboiling mixture Tray efficiency can be low Stripping Feed is liquid Distillation Feed is liquid, vapor or a mixture of the two Second phase (stripping agent) Second phase is created by is added heat transfer Operation can be adiabatic Single-section cascade Can be almost isothermal Can not separate a closeboiling mixture Tray efficiency can be moderate Must have heat transfer at the top and bottom stages Two-section cascade Can have a large temperature range Can separate a close-boiling mixture Tray efficiency can be high Exercise 7.2 Subject: Emergence of packing to replace trays. Find: Reasons why some existing trayed towers are being retofitted with packing, and some large-diameter columns are being designed for packing. Analysis: New random packings and structured packings have been introduced with higher capacity, lower pressure drop, and higher efficiency than trays. Also for packed columns, liquid distributors have been greatly improved. Channeling of liquid in packed columns is now much less of a problem. Exercise 7.3 Subject: Condenser coolant for distillation of a methane-ethane mixture. Find: Appropriate coolant. Analysis: Assume that the distillate is nearly pure methane, the most volatile of the two components. To use cooling water, the distillate temperature would be approximately 120oF. But the critical temperature of methane is -115oF. Therefore could not condense it. The critical pressure of methane is 673 psia. Using Fig. 7.16, could consider operation at 415 psia because that is safely below the critical pressure (P/Pc = 0.62 and could operate as high as a P/Pc = 0.8). From Perry's Handbook (p. 3-203, 6th edition), at P = 415 psia (28.6 bar), the saturation temperature of methane is 176 K or -143oF. To condense at just below this temperature, could use a partial condenser with refrigerant R-14 (carbon tetrafluoride), which from Perry's Handbook (p. 12-26, 6th edition) has a saturation pressure of 70 psia at -150oF. Exercise 7.4 Subject: Operating pressure for distillation of an ethylene-ethane mixture Find: Suitable operating pressure. Analysis: Assume that the distillate is nearly pure ethylene, the most volatile of the two components. To use cooling water, the distillate temperature would be approximately 120oF. But the critical temperature of ethylene is 9.7oC or 49.4oF . Therefore could not condense it. The critical pressure of ethylene is 50.5 atm or 742 psia. Using Fig. 7.16, could consider operation at 415 psia because that is safely below the critical pressure (P/Pc = 0.56 and could operate as high as a P/Pc = 0.8). From Perry's Handbook (p. 3-203, 6th edition), at P = 415 psia (28.6 bar), the saturation temperature of ethylene is 258 K or 4.7oF. To condense at just below this temperature could use a partial condenser with refrigerant R-1270 (propylene), which from Perry's Handbook (p. 12-26, 6th edition) has a saturation pressure of 35 psia at 0oF. Exercise 7.5 Subject: Need for testing or piloting a distillation separation. Find: Circumstances that require laboratory or pilot-plant testing of a proposed distillation. Analysis: Laboratory and/or pilot-plant testing is recommended for: 1. A new mixture not previously separated by distillation. 2. A sharp and critical separation. 3. A mixture with uncertain vapor-liquid equilibria data. 4. A lack of experience with tray efficiency for the mixture. Exercise 7.6 Subject: Economic tradeoff in distillation. Find: Reasons for tradeoff between trays and reflux. Analysis: It is well known that for a given separation, as the number of trays is increased, the reflux ratio can be decreased. Thus, as the tower height is increased, the vapor and liquid traffic up and down the column can be decreased. Therefore, the column diameter can be decreased. Also the condenser and reboiler duties and sizes, and the utility requirements can be decreased. Therefore, there is a tradeoff. Exercise 7.7 Subject: McCabe-Thiele method for binary distillation. Find: Reasons for the survival of the McCabe-Thiele graphical method. Analysis: For a binary mixture, the McCabe-Thiele method shows clearly the ease or difficulty of the separation. Pinched regions are readily seen. The effect of feed location is readily seen. The method is reasonably accurate. Azeotropes are readily accommodated. Exercise 7.8 Subject: Separation of ethyl alcohol and water at 1 atm. with two countercurrent cascades. Given: One cascade (a) with given liquid feed to top stage and given vapor feed to bottom stage. Another cascade (b) with total condenser and reflux, and given vapor feed to bottom stage. Vapor-liquid equilibrium data for 1 atm. Assumptions: Constant molar overflow to give straight operating lines on a y-x diagram. Find: (a) (b) (c) (d) Compositions of V4 and L1 for 4 stages in cascade (a). Number of equilibrium stages for 85 mol% alcohol in exit vapor of cascade (a). Compositions of D and L1 for 4 stages in cascade (b). Number of equilibrium stages for 50 mol alcohol in D of cascade (b). Analysis: From the given vapor-liquid equilibrium data, in the composition range of interest, ethyl alcohol is more volatile than water. Therefore, the y and x coordinates in a y-x plot pertain to ethyl alcohol. (a) Since L = 100 mol and V = 100 mol, the slope of the operating line from Eqs. (7-6) or (7-11) = L/V = 100/100 = 1. The terminal points on the operating line as (y, x) are: (?, 0.7) at the top and (0.3, ?) at the bottom. To determine the compositions of V4 and L1 for 4 stages, this operating line is located so that exactly 4 stages are stepped off in a y-x diagram, as shown below. From the diagram, the ethanol compositions are 76 mol% in V4 and 24 mol% in L1. Analysis: (continued) (b) It is impossible to obtain an overhead vapor with 85 mol% ethanol. With an infinite number of stages, the highest concentration of ethanol in the overhead vapor corresponds to that in equilibrium with the top liquid feed containing 70 mol% ethanol. From the given vapor-liquid equilibrium data, the highest concentration is an ethanol mole fraction of 0.82. (c) Since the bottom vapor feed, V0 = 100 mol and D = 50 mol, by overall material balance, L1 = V0 - D = 100 - 50 = 50 mol. Because of the assumption of constant molar overflow, L = LR = L1 = 50 mol. By material balance around the condenser or because of constant molar overflow, V = V4 = LR + D = 50 + 50 = 100 mol. The slope of the operating line from Eqs. (7-6) or (7-11) = L/V = 50/100 = 0.5. To determine the compositions of D and L1 for 4 stages, an operating line of this slope is located so that exactly 4 stages are stepped off in a y-x diagram, as shown below. From the diagram, the ethanol compositions are 45 mol% in D and 16 mol% in L1. (d) Since the distillate is 50 mol% ethanol, 25 moles of ethanol and 25 moles of water leave in the distillate. Because the feed is 30 moles of ethanol and 70 moles of water, L1 , the leaving liquid, contains 5 moles of ethanol and 45 moles of water. Thus, the terminal points on the operating line, because of the total condenser, as (y, x), are: (0.5, 0.5) at the top and (0.3, 0.1 ) at the bottom. However, point (0.3, 0.1) above the equilibrium line is impossible. Exercise 7.8 (continued) Exercise 7.9 Subject: Separation of air in a reboiled stripper Given: Reboiled stripper with total reboiler operating at 1 atm. Liquid air (79.1 mol% N2 and 20.9 mol% O2) fed to top stage. 60% of O2 in the feed is drawn off in vapor product from the reboiler. Bottoms vapor product contains 0.2 mol% N2. Vapor-liquid equilibria data are given. Assumptions: Feed is a saturated liquid. Find: (a) Mol% N2 in vapor from top stage. (b) Moles of vapor generated in reboiler per 100 moles of feed. (c) Number of equilibrium stages required. Analysis: (a) Take a basis of F = 100 mol/h. Therefore, 79.1 mol/h of N2 and 20.9 mol/h of O2 in the feed. Bottoms product vapor contains 0.6(20.9) = 12.54 mol/h of O2 and (0.2/99.8)(12.54) = 0.025 mol/h N2. By material balance, the overhead vapor contains 20.9 - 12.54 = 8.36 mol/h O2 and 79.1 - 0.025 = 79.075 mol/h of N2. The mol% N2 in the overhead vapor = 79.075/(8.36 + 79.075) x 100% = 90.4 %. (b) Assume constant molar overflow. Then because the feed is assumed to be a saturated liquid, the moles of vapor generated in the reboiler per 100 moles of feed = mol/h of overhead vapor = 8.36 + 79.075 = 87.435 moles per 100 moles of feed. (c) Use a y-x diagram for N2 because it is the more volatile. The slope of the operating line is L/V = 100/87.435 = 1.14. At the top of the column, the operating line terminates at a (y-x) of (0.904, 0.791). At the bottom of the column, with a total reboiler, the operating line terminates at a (y-x) of (0.002, 0.002). To determine the number of equilibrium stages, it is convenient to use two diagrams, the usual one and a second one for just the very low mole fraction region so as to gain accuracy in the region of the lower end of the operating line. From the two diagrams, it is seen that just less than 8 equilibrium stages are required. Analysis (d) (continued): Exercise 7.9 (continued) Exercise 7.10 Subject: Stage composition data for distillation of an A-B mixture. Given: Saturated liquid feed of 40 mol% A. Test results for vapor and liquid compositioins for 3 successive stages between the feed stage and a total condenser. ____________Mol % A_____________ Test 1 Test 2 Stage Vapor Liquid Vapor Liquid M+2 79.5 68.0 75.0 68.0 M+1 74.0 60.0 68.0 60.5 M 67.9 51.0 60.5 53.0 Assumption: Constant molar overflow. Find: Reflux ratio and overhead composition for each test. Analysis: Test 1: By material balance for component A around stage M + 1, Vy M + Lx M + 2 = Vy M +1 + Lx M +1 (1) Solving Eq. (1) for L/V and substituting for y and x values from above table, L y - yM 0.740 - 0.679 = M +1 = = 0.763 V x M + 2 - x M +1 0.680 - 0.600 From Eq. (7-7), reflux ratio = R = L/D = (L/V)/(1 - L/V) = 0.763/(1 - 0.763) = 3.22 D/V = (L/V)/(L/D) = 0.763/3.22 = 0.237. Noting that stages in the rectifying section are counted here from the bottom up instead of the top down, Eq. (7-5) becomes, L D y M +1 = x M + 2 + x D (2) V V Solving Eq. (2) for xD , L y M +1 - x M +2 0.740 - (0.763)(0.680) V xD = = = 0.933 D /V 0.237 Therefore the composition of the distillate is 93.3 mol% A and 6.7 mol% B. Test 2: Because y M +1 = x M + 2 and y M = x M +1 , operation is at total reflux, i.e. reflux ratio = infinity. In this case, Eq. (2) can not be solved for xD because D/V = 0. We can not determine the composition of the distillate. We do know that it must at least 75 mol% A. Exercise 7.11 Subject: Five procedures for continuous distillation of a mixture of benzene (A) and toluene (B). Given: Operation at 1 atm to produce a distillate of 80 mol% benzene (i.e. yD = 0.8) from a saturated liquid feed of 70 mol% benzene (xF = 0.7) . Procedures are: 1. No column. Just a partial condenser on top of a partial reboiler. Feed is to the reboiler. Reflux ratio, L/D = 0.5. Vapor distillate is totally condensed. 2. Same as 1 except that one equilibrium stage sits between the condenser and reboiler. 3. Same as 1 except two equilibrium stages between the condenser and reboiler. 4. Same as 3 except that reflux bypasses the top equilibrium stage. 5. Same as 2 except that feed is sent to the stage between the condenser and the reboiler. Assumptions: Constant molar overflow. Constant relative volatility = , = 2.5. Find: For each procedure, determine: (a) Moles of distillate per 100 moles of feed. (b) Moles of total vapor generated per mole of distillate. (c) Mole percent of benzene in the bottoms (residue). (d) y-x diagram, indicating compositions of distillate, reflux, and residue. Also, (e) For maximization of benzene recovery (in the distillate), which procedure is preferred. Analysis: For each procedure, the partial condenser and the partial reboiler are equilibrium stages. Benzene is the more volatile component, so the y-x diagram is based on benzene. Because the relative volatility = constant = 2.5, the equilibrium relationship is given by Eq. (7-3), x 2.5x = (1) 1 + x ( - 1) 1 + 15x . Take as a basis, 100 mol/s of feed. Therefore, the feed contains 70 mol/s of A and 30 mol/s of B. From the reflux ratio, L = 0.5D, V = L + D = 1.5D. Therefore, D/V = 2/3 and L/V = 1/3. Use a subscript of C for streams leaving the condenser, R for streams leaving the reboiler, 1 for the top stage when used, and 2 for the second stage when used. y= Procedure 1: Solve with material balances and Eq. (1). The liquid leaving the partial condenser is in equilibrium with the vapor distillate of yC = yD = 0.8. Solving Eq. (1), yC 0.8 = = 0.615 yC + (1 - yC ) 0.8 + 2.5(1 - 0.8) Benzene material balance around condenser, xC = (2) y RV = yC D + xC L or y R = yC D L 2 1 + xC = 0.80 + 0.615 = 0.738 V V 3 3 Exercise 7.11 (continued) Analysis: Procedure 1 (continued) The vapor from the reboiler is in equilibrium with the liquid bottoms (residue). From the lefthand part of Eq. (2), yR 0.738 xR = = = 0.530 yR + (1 - yR ) 0.738 + 2.5(1 - 0.738) Overall total material balance, F = 100 = D + B (3) Overall benzene material balance, xFF = yCD + xRB or 70 = 0.8D + 0.530B (4) Solving Eqs. (3) and (4), D = 62.9 mol/s or 62.9 mol/100 mol feed, and B = 37.1 mol/s. Therefore, vapor generated = V = 1.5D = 1.5(62.9) = 94.4 mol/s. The operating line for the y-x diagram passes through the (y, x) point (0.8, 0.8) with a slope, L/V = 1/3, as shown in the diagram below. Analysis: (continued) Procedure 2: The slope and top point of the operating line are the same as for Procedure 1. We just have to step off one more stage. Therefore from the results above, we have: yC = 0.80 xC = 0.615 y1 = 0.738 x1 = 0.530 Benzene material balance around Stage 1, y RV + xC L = y1V + x1 L (5) Solving for yR, yR = y1 + ( x1 - xC ) Exercise 7.11 (continued) L 1 = 0.738 + (0.530 - 0.615) = 0.710 V 3 The vapor from the reboiler is in equilibrium with the liquid bottoms (residue). From the lefthand part of Eq. (2), yR 0.710 xR = = = 0.495 yR + (1 - yR ) 0.710 + 2.5(1 - 0.710) Overall total material balance, F = 100 = D + B (6) Overall benzene material balance, xFF = yCD + xRB or 70 = 0.8D + 0.495B (7) Solving Eqs. (6) and (7), D = 67.2 mol/s or 67.2 mol / mol feed, and B = 32.8 mol/s. Therefore, vapor generated = V = 1.5D = 1.5(67.2) = 100.8 mol/s The operating line for the y-x diagram passes through the (y, x) point (0.8, 0.8) with a slope, L/V = 1/3, as shown in the diagram below. Analysis: (continued) Procedure 3: The slope and top point of the operating line are the same as for Procedures 1 and 2. We just have to extend Procedure 2 by stepping off a second equilibrium stage. From above, the results for the condenser and stage 1 are: yC = 0.80 xC = 0.615 y1 = 0.738 x1 = 0.530 y2 = 0.710 x2 = 0.495 Benzene material balance around Stage 2, y RV + x1 L = y2V + x2 L (8) L 1 Solving for yR, y R = y2 + ( x2 - x1 ) = 0.710 + (0.495 - 0.530) = 0.698 V 3 The vapor from the reboiler is in equilibrium with the liquid bottoms (residue). From the lefthand part of Eq. (2), yR 0.698 xR = = = 0.480 yR + (1 - yR ) 0.698 + 2.5(1 - 0.698) Overall total material balance, F = 100 = D + B (10) Overall benzene material balance, xFF = yCD + xRB or 70 = 0.8D + 0.480B (11) Solving Eqs. (10) and (11), D = 68.8 mol/s or 68.8 mol / 100 mol feed, and B = 31.2 mol/s. Therefore, vapor generated = V = 1.5D = 1.5(68.8) = 103.2 mol/s The operating line for the y-x diagram passes through the (y, x) point (0.8, 0.8) with a slope, Exercise 7.11 (continued) L/V = 1/3, as shown in the diagram below. Analysis: (continued) Procedure 4: If the reflux bypasses the top stage, the vapor and liquid pass through that stage without change. Therefore, this procedure is the same as Procedure 2, i.e. just one stage in the column. Procedure 5: The slope and top point of the operating line are the same as for Procedure 1. We just have to add the feed to the stage in the column. Therefore from the results above, we have: yC = 0.80 xC = 0.615 y1 = 0.738 x1 = 0.530 Benzene material balance around Stage 1, which now includes the feed, xFF + y RV + xC L = y1V + x1 L (12) Solving for yR, V L L F V L L 100 y R = y1 + x1 - xC - xF = 0.738 + 0.530 - 0.615 - 0.70 (13) V V V V V V V V Because the feed is a saturated liquid, , V =V and Exercise 7.11 (continued) L = L + 100 From above, V = 1.5D and L/V = 1/3. Also, L / V = L / V + 100 / V = 1 / 3 + 100 / V Therefore, Eq. (13) becomes, 1 100 1 100 17 11.33 (14) y R = 0.738 + 0.530 + - 0.615 - 0.70 = 0.710 - = 0.710 - 3 V 3 V V D The vapor from the reboiler is in equilibrium with the liquid bottoms (residue). From the lefthand part of Eq. (2), yR xR = (15) y R + 2.5(1 - y R ) Overall total material balance, F = 100 = D + B (16) (17) Overall benzene material balance, xFF = yCD + xRB or 70 = 0.8D + xRB Solving Eqs. (14), (15), (16), and (17), yR = 0.565, xR = 0.342, D = 78.3 mol/s or 78.3 mol/100 mol feed, B = 21.7 mol/s Therefore, vapor generated = V = 1.5D = 1.5(78.3) = 117.5 mol/s The rectifying section operating line for the y-x diagram passes through the (y, x) point (0.8, 0.8) with a slope, L/V = 1/3, the q-line (feed line) is a vertical line, and the stripping section operating line passes throught the (y, x) point (0.565, 0.342) and the intersection of the other two lines, as shown in the diagram on the following page. The results from the 5 procedures are summarized as follows: Procedure D/100 moles of feed V/mole of D x of benzene in B 1 62.9 1.5 0.530 2 67.2 1.5 0.495 3 68.8 1.5 0.480 4 67.2 1.5 0.495 5 78.3 1.5 0.342 Analysis: Procedure 5: (continued) Exercise 7.11 (continued) (e) Procedure 5 is recommended because it produces by far the most distillate, which corresponds to the highest recovery of benzene. Exercise 7.12 Subject: Distillation of a mixture of benzene (A) and toluene (B) at 101 kPa. Given: Column consisting of a partial reboiler, one theoretical plate, and a total condenser. Produce a distillate of 75 mol% benzene from a saturated liquid feed of 50 mol% benzene. Assumptions: Constant molar overflow. Constant relative volatility = , = 2.5. Find: Number of moles of distillate per 100 moles of feed for: (a) Feed to the reboiler and no reflux. (b) Feed to the reboiler and a reflux ratio, L/D = 3. (c) Feed to the plate and a reflux ratio of 3. (d) Same as (c) except a partial condenser. (e) Feed to the reboiler with minimum reflux. (f) Feed to the reboiler with total reflux. Analysis: Either a graphical or analytical method can be used. Because the relative volatility is assumed constant, use an analytical method. For each part, the theoretical plate and the partial reboiler are equilibrium stages. Benzene is the more volatile component, so the y-x diagram is based on benzene. Because the relative volatility = constant = 2.5, the equilibrium relationship is given by Eq. (7-3), x 2.5x y= = (1) 1 + x ( - 1) 1 + 15x . Take as a basis, 100 moles of feed. Therefore, the feed is 50 moles of A and 50 moles of B. (a) With no reflux, separation occurs only in the reboiler. The vapor leaving the reboiler is totally condensed to become the distillate with yD = xD = 0.75. Solve Eq. (1) for equilibrium x, 0.750 yD = = 0.545 (2) y D + (1 - y D ) 0.750 + 2.5(1 - 0.750) Because the distillate and bottoms have benzene mole fractions greater than the mole fraction of the feed (0.5), it is impossible to obtain a distillate with a benzene mole fraction of 0.75. xB = (b) From the reflux ratio, L = 3D, V = L + D = 4D. Therefore, D/V = 1/4 and L/V =3/4. Use a subscript of D for distillate, R for reflux, B for streams leaving the reboiler, and 1 for the theoretical plate, when used. With 1 theoretical plate, from part (a), y1 = 0.75 xD = 0.75 x1 = 0.545 Benzene material balance around plate 1, y BV + x D L = y1V + x1 L (3) Solving for yB L 3 y B = y1 + ( x1 - x D ) = 0.750 + (0.545 - 0.750) = 0.596 V 4 Analysis: (b) (continued) Exercise 7.12 (continued) The mole fraction of benzene in the bottoms product is in equilibrium with yB =0.596. Therefore, the form of Eq. (2) applies, yB 0.596 xB = = = 0.371 y B + (1 - y B ) 0.596 + 2.5(1 - 0.596) Overall total material balance, F = 100 = D + B (4) Overall benzene material balance, xFF = xDD + xBB or 50 = 0.75D + 0.371B (5) Solving Eqs. (4) and (5), D = 34.2 moles or 34.2 mol/100 mol feed, and B = 65.8 moles. (c) With the feed to the theoretical plate, the following results apply from part (b), y1 = 0.75 xD = 0.75 x1 = 0.545 Benzene material balance around Stage 1, which now includes the feed, xFF + y BV + x D L = y1V + x1 L (6) Solving for yB, V L L F V L L 100 y B = y1 + x1 - xD - xF = 0.750 + 0.545 - 0.750 - 0.50 (7) V V V V V V V V Because the feed is a saturated liquid, , V =V and L = L + 100 From above, V = 4D and L/V = 3/4. Also, L / V = L / V + 100 / V = 3 / 4 + 100 / V Therefore, Eq. (7) becomes, 3 100 3 100 4.5 1125 . (8) y B = 0.750 + 0.545 + - 0.750 - 0.50 = 0.596 - = 0.596 - 4 V 4 V V D The vapor from the reboiler is in equilibrium with the liquid bottoms (residue). From the lefthand part of Eq. (2), yB xB = (9) y B + 2.5(1 - y B ) Overall total material balance, F = 100 = D + B (10) Overall benzene material balance, xFF = xDD + xBB or 50 = 0.75D + xBB (11) Solving Eqs. (8), (9), (10), and (11), yB = 0.647, xB = 0.423, D = 23.5 moles or 23.5 mol/100 mol feed, B = 76.5 moles (d) With a partial condenser, the mole fraction of the liquid reflux is that in equilibrium with the vapor distillate. Therefore, from the above results, yD = 0.75 xR = 0.545 y1 = 0.596 x1 = 0.371 Benzene material balance around the theoretical plate, which includes the feed, xFF + y BV + x R L = y1V + x1 L (12) Solving for yB, V L L F V L L 100 y B = y1 + x1 - xR - xF = 0.596 + 0.371 - 0.545 - 0.50 (13) V V V V V V V V Analysis: (d) (continued) Because the feed is a saturated liquid, From above, V = 4D and L/V = 3/4. Therefore, Eq. (13) becomes, 3 100 y B = 0.596 + 0.371 + - 0.545 4 V Exercise 7.12 (continued) , V =V and L = L + 100 Also, L / V = L / V + 100 / V = 3 / 4 + 100 / V 3 100 12.9 3.23 - 0.50 = 0.466 - = 0.466 - 4 V V D (14) The vapor from the reboiler is in equilibrium with the liquid bottoms (residue). From the lefthand part of Eq. (2), yB xB = (15) y B + 2.5(1 - y B ) Overall total material balance, F = 100 = D + B Overall benzene material balance, xFF = yDD + xBB or 50 = 0.75D + xBB Solving Eqs. (14), (15), (16), and (17), yB = 0.405, xB = 0.214, D = 53.4 moles or 53.4/100 mol feed, (16) (17) B = 46.6 moles (e) At minimum reflux, with the feed sent to the still pot (partial reboiler), an infinite number of theoretical plates is needed between the condenser and reboiler. This part is not completely specified. In order to compute the distillate, we must assume a bottoms benzene mole fraction less than that in the feed. Suppose we choose that mole fraction to be 0.45. Then the operating line will intersect the equilibrium line at x = 0.45, creating the pinch zone of infinite stages. The value of y at the intersection is given by Eq. (1): 2.5(0.45) = 0.672 1 + 15(0.45) . Therefore the operating line passes through the two points, as {y, x}, of {0.75, 0.75} and {0.672, 0.450}. Therefore, the slope = L/V = (0.75 - 0.672)/(0.75 - 0.45) = 0.260. Now compute the overall material balances: Overall total material balance, F = 100 = D + B (18) Overall benzene material balance, xFF = yDD + xBB or 50 = 0.75D + 0.45B (19) Solving Eqs. (18) and (19), D = 16.67 moles or 16.67 mol/100 mol feed, and B = 83.33 moles y= Calculations for other values of the benzene mole fraction in the bottoms can be made in the same manner. (f) At total reflux, there is no distillate, but there is a boilup. The moles of distillate per 100 moles of feed = 0. Exercise 7.13 Subject: Distillation of a mixture of benzene and toluene at 101 kPa for specified reflux ratio and product compositions. Given: Feed of 30 kg/h of saturated liquid feed containing 40 mass% benzene and 60 mass% toluene. Distillate to contain 97 mass% benzene and bottoms to contain 98 mass% toluene. Reflux ratio = 3.5 and feed is to optimal stage. Table of vapor-liquid equilibrium data in mole fractions. At 101 kPa. Assumptions: Total condenser and partial reboiler. Saturated liquid reflux. Constant molar overflow. Find: (a) Flow rates of distillate and bottoms. (b) Number of equilibrium stages needed. Analysis: First solve the material balance in mass units. Then convert to moles and mole fractions so that the McCabe-Thiele method can be used for part (b). Overall total mass balance: 30 = D + B (1) (2) Overall benzene mass balance: 0.40(30) = 12 = 0.97D + 0.02B D = 12.1 kg/h B = 17.9 kg/h Solving Eqs. (1) and (2): Converting to moles with molecular weights of 78.11 for benzene and 92.13 for toluene, Benzene Toluene Product kmol/h Mass fraction Mole fraction Mass fraction Mole fraction Distillate 0.154 0.97 0.974 0.03 0.0235 Bottoms 0.196 0.02 0.026 0.98 0.9765 1.00 1.000 1.00 1.0000 Total: 0.350 (b) Because benzene is the more volatile component of the feed, the x and y coordinates will be those of benzene in the diagram on the next page. . In moles, the feed consists of: Component kmol/h Mole fraction Benzene 0.154 0.44 Toluene 0.196 0.56 Total: 0.350 1.00 For a saturated liquid feed, the q-line is vertical and passes through x = 0.44. The slope of the rectifying operating line, L/V, is obtained from Eq. (7-7), using the specified reflux ratio = 3.5, L/V = R/(1 + R) = 3.5/4.5 = 0.778 For saturated liquid reflux, the rectifying operating line passes through the point {0.974, 0.974}. See the McCabe-Thiele construction on the next page, where it is seen that slightly more than 10 stages + a partial reboiler acting as an equilibrium stage are required. The top 5 stages are in the rectifying section. Analysis: (b) (continued) Exercise 7.13 (continued) McCabe-Thiele Diagram Exercise 7.14 Subject: Distillation of a mixture of benzene and chlorobenzene with specified number of equilibrium stages, boilup ratio, and reflux ratio. Given: Feed is a saturated liquid of 54.5 mol% benzene. Column contains two equilibrium plates with feed to the bottom plate. Column is equipped with total condenser and partial reboiler. Boilup is V/F = 0.855. Reflux ratio, L/V = 0.5. Vapor-liquid equilibrium data given. Find: Compositions of distillate and bottoms, assuming constant molar overflow. Analysis: Take as a basis, F = 100 mol/s. Therefore, vapor generated in reboiler = 0.855(100) = 85.5 mol/s. Since the feed is a saturated liquid, this vapor rate continues up the column to the condenser. L/V = 0.5, which is the slope of the operating line. Therefore, L = 0.5(85.5) = 42.75 mol/s. Therefore, the distillate rate = 85.5 - 42.75 = 42.75 mol/s. Passing to the reboiler is a liquid rate of 42.75 + 100 = 142.75 mol/s. The bottoms rate = 142.75 - 85.5 = 57.25 mol/s. The slope of the stripping section operating line is L / V = 142.75/85.5 = 1.67. The q-line is a vertical line because the feed is a saturated liquid. To solve for the compositions of the distillate and bottoms on a McCabe-Thiele diagram, we must locate the operating lines to obtain three equilibrium stages that satisfy an overall benzene material balance given by, xFF = 54.4 = xDD + xBB = 42.75xD + 57.25xB Solving Eq. (1), xB = 0.952 - 0.7467 xD (1) (2) Therefore, an approach to solving this exercise is to assume a value of xD and then compute the value of xB from Eq. (2). Then construct the McCabe-Thiele diagram with the above operating lines and q-line to see if three stages are required with the feed to the second plate. See plot below, where benzene mole fractions are plotted because it is the more volatile component. It is seen that for benzene, xD = 0.90 and xB = 0.28. Exercise 7.14 (continued) McCabe-Thiele Diagram Exercise 7.15 Subject: Effect of loss of plates in a distillation column separating a benzene-toluene mixture. Given: Saturated vapor feed of 13,600 kg/h of 40 wt% benzene and 60 wt% toluene. Column with 14 plates above the feed location. Plate efficiency is 50%. Reflux ratio is 3.5. Previously, with 10 plates in the stripping section, column could achieve a distillate of 97 wt% benzene and a bottoms of 98 wt% toluene. Vapor-liquid equilibrium data in Exercise 7.13. Assumptions: Constant molar overflow. Total condenser and partial reboiler. Find: (a) If column with 10 inoperable plates can yield a distillate of 97 wt% benzene, assuming that we no longer can achieve the 98 wt% bottoms product. (b) The distillate flow rate. (c) The composition of the bottoms. Analysis: (a) First convert the feed to kmol/h and mole fractions, using molecular weights of 78.11 for benzene and 92.13. The result is: Component kmol/h Mole fraction Benzene 69.65 0.44 Toluene 88.57 0.56 Total: 158.22 1.00 For a distillate of 97 wt% benzene, the mole fraction for benzene, the more volatile of the two components, is, 97 7811 . xD = = 0.974 97 3 + 78.11 92.13 With a reflux ratio of 3.5, from Eq. (7-7), the slope of the rectifying section operating line is, L/V = R/(1 + R) = 3.5/4.5 = 0.778 The q-line is a horizontal line at y = 0.44. For 14 plates with 50% efficiency, the column has the equivalent of 7 equilibrium stages + 1 for the partial reboiler. The McCabe-Thiele construction is shown on the next page, where it is seen that it is possible to obtain the desired distillate composition. (b) and (c) From the McCabe-Thiele diagram, the mole fraction of benzene in the bottoms is xB = 0.24. As a weight fraction, this corresponds to, 0.24(78.11) = 0.211 weight fraction or 21.1 wt% benzene 0.24(7811) + 0.76(92.13) . Compute the distillate rate by overall molar material balances. Overall total mass balance: 158.22 = D + B (1) Overall benzene mass balance: 69.65 = 0.974D + 0.240B (2) Solving Eqs. (1) and (2): D = 43.16 kmol/h B = 115.06 kmol/h By weight, D = 43.16[0.974(78.11) + 0.026(92.13)] = 3,387 kg/h Analysis: (a) (continued) Exercise 7.15 (continued) McCabe-Thiele Diagram Exercise 7.16 Subject: Effect on the separation of A from B by distillation when 3 of 7 theoretical plates rust and drop to the bottom of the column. Given: . Column has 7 theoretical plates + partial reboiler. Saturated liquid feed of 100 kmol/h of 50 mol% A is sent to plate 5 from the top. Distillate contains 90 mol% A. The L/V = 0.75 in the rectifying section. Vapor-liquid equilibrium data. Assumptions: Constant molar overflow. Total condenser. Find: Case 1: Column before the 3 plates rust and drop. (a) Composition of the bottoms product. (b) The L/V in the stripping section. (c) The kmol/h of bottoms product. Case 2: If plates 5, 6, and 7 counted down from the top are lost: (a) Composition of bottoms product. Case 3: Same as Case 2, except replace reflux with the same molar flow rate of product containing 80 mol% A: (a) Composition of distillate. (b) Composition of bottoms. Analysis: Case 1: Apply the McCabe-Thiele method in terms of component A, which is more volatile than B. The rectifying section operating line passes through [0.90, 0.90] with a slope of 0.75. The q-line is vertical through x = 0.50. Step off 4 stages in the rectifying section. Then, by trial and error, find an xB with a corresponding stripping section operating line that gives 4 equilibrium stages in the stripping section. The result is shown on the following page, where: (a) Bottoms contains 7 mol% A and 93 mol% B. (b) The slope of the stripping section operating line from the coordinates of the line is: L / V = {[0.90 - 0.75(0.90 - 0.50] - 0.07}/(0.50 - 0.07) = 1.23 (c) By material balances, F = D + B and FxF = 50 = 0.9D + 0.07B. Solving these two equations, distillate flow rate = 51.8 kmol/h and bottoms flow rate = 48.2 kmol/h Case 2: We now have 4 equilibrium stages and a partial reboiler, with the feed being sent to the reboiler. Assume that utility rates are such that L/V and L / V are the same as in Case 1. Then, on the McCabe-Thiele diagram, the values of xD and xB must shift so that 5 stages are stepped off, with the fifth, which is the reboiler, intersecting the stripping section operating line at the 45o line. This is shown on the McCabe-Thiele diagram on the following page. (a) From this diagram, the mole fractions of benzene in the distillate and bottoms are 0.80 and 0.21, respectively. Case 3: Since the distillate composition in Case 2 is 80 mol% benzene, the results would be the same as Case 2 if an 80 mol% benzene stream from another column were used as reflux. Exercise 7.16 (continued) Analysis: Case 1 (continued) Exercise 7.16 (continued) Analysis: Case 2 (continued) Exercise 7.17 Subject: Distillation of a mixture of benzene and toluene with different feed conditions. Given: Column with 7 equilibrium plates, total condenser, and partial reboiler. Feed is 50 mol% benzene and 50 mol% toluene. Operation at 101 kPa to produce a distillate of 96 mol% benzene. Equilibrium data from Exercise 7.13. Assumptions: Constant molar overflow. Use a basis of 100 mol/s for the feed. (a) For a saturated liquid feed sent to tray 5 from the top, (1) minimum reflux ratio, R = L/D, (2) bottoms composition for twice the minimum reflux ratio, and (3) moles of products per 100 moles of feed. (b) Same as (a) except feed is saturated vapor to tray 5 from the top. (c) For a saturated vapor feed to the reboiler and a reflux ratio, L/V, of 0.9, determine, (1) bottoms composition and (2) moles of products per 100 moles of feed. Find: Analysis: (a) (1) For a saturated liquid feed, minimum reflux corresponds to a pinch point located at the intersection of a vertical q-line passing through xF = 0.5 and the equilibrium curve. From the equilibrium data, this intersection is at y = 0.72 and x = 0.5. Then, the slope of the rectifying section operating line, (L/V)min is (0.96 - 0.72)/(0.96 - 0.50) = 0.522. From a rearrangement of Eq. (7-7), Rmin = (L/V)min /[1 - (L/V)min ] = 0.522/(1 - 0.522) = 1.092. (2) Reflux ratio = 2(1.092) = 2.18. From Eq. (7-7), the slope of the rectifying section operating line = L/V = R/(1 + R) = 2.18/3.18 = 0.686. To determine the bottoms composition, use a McCabe-Thiele diagram in terms of benzene, the more volatile component. The q-line and the rectifying section operating line are fixed and 4 trays are stepped off from the top, starting at the distillate mole fraction for benzene, xD , of 0.96. Then, the stripping section operating line is positioned by trial and error so that 3 more stages plus the reboiler stage are stepped off to arrive at the point where the assumed location of the stripping section operating line intersects the 45o line. The result is shown on the next page where it is seen that xB = 0.18. (3) The products are now computed by overall material balances: F = 100 = D + B and 50 = xDD + xBB = 0.96D + 0.18B. Solving these two equations, D = 41.0 mol/100 mol feed and B = 59.0 mol/100 mol feed. Analysis: (a) (continued) Exercise 7.17 (continued) (b) (1) For a saturated vapor feed, minimum reflux corresponds to a pinch point located at the intersection of a horizontal q-line passing through xF = y = 0.5 and the equilibrium curve. From the equilibrium data, this intersection is at y = 0.50 and x = 0.293. Then, the slope of the rectifying section operating line, (L/V)min is (0.96 - 0.50)/(0.96 - 0.293) = 0.690. From a rearrangement of Eq. (7-7), Rmin = (L/V)min /[1 - (L/V)min ] = 0.69/(1 - 0.69) = 2.23. (2) Reflux ratio = 2(2.23) = 4.46. From Eq. (7-7), the slope of the rectifying section operating line = L/V = R/(1 + R) = 4.46/5.46 = 0.817. To determine the bottoms composition, use a McCabe-Thiele diagram in terms of benzene, the more volatile component. The q-line and the rectifying section operating line are fixed and 4 trays are stepped off from the top, starting at the distillate mole fraction for benzene, xD , of 0.96. Then, the stripping section operating line is positioned by trial and error so that 3 more stages plus the reboiler stage are stepped off to arrive at the point where the assumed location of the stripping section operating line intersects the 45o line. The result is shown below, where it is seen that xB = 0.08. Analysis: (b) (continued Exercise 7.17 (continued) (3) The products are now computed by overall material balances: F = 100 = D + B and 50 = xDD + xBB = 0.96D + 0.08B. Solving these two equations, D = 47.7 mol/100 mol feed and B = 52.3 mol/100 mol feed. Analysis: (continued) (c) (1) A saturated vapor feed is fed to the reboiler. The slope of the rectifying section operating line, (L/V), is 0.9. To determine the bottoms composition, use a McCabe-Thiele diagram in terms of benzene, the more volatile component. The q-line and the rectifying section operating line are fixed and 7 trays are stepped off from the top, starting at the distillate mole fraction for benzene, xD , of 0.96. Then, the stripping section operating line is positioned by trial and error so that the reboiler stage is stepped off to arrive at the point where the assumed location of the stripping section operating line intersects the 45o line. The result is shown below, where it is seen that xB = 0.07. (2) The products are now computed by overall material balances: F = 100 = D + B and 50 = xDD + xBB = 0.96D + 0.07B. Solving these two equations, D = 48.3 mol/100 mol feed and B = 51.7 mol/100 mol feed. Exercise 7.17 (continued) Exercise 7.18 Subject: Conversion of a distillation column to a reboiled stripper to obtain very pure toluene from a mixture of benzene and toluene at 101 kPa. Given: A column with 8 theoretical plates, a total condenser, and a partial reboiler. Feed contains 36 mol% benzene and 64 mol% toluene. Reboiler produces 100 kmol/h of vapor. To obtain nearly pure toluene bottoms, feed is introduced to the top plate as a saturated liquid, with no reflux. Vapor-liquid equilibrium data are in Exercise 7.13. Assumptions: Constant molar overflow. Find: (a) Minimum feed rate and corresponding bottoms composition. (b) Bottoms rate and composition for a feed rate 25% above the minimum. Analysis: (a) The minimum feed rate corresponds to a rate equal to the boilup rate so as to give an L / V = 1.0. Thus, the minimum feed rate = 100 kmol/h. Under these conditions, no bottoms product is withdrawn and the reboiler performs as a total reboiler. The vapor leaving the top of the column has the same composition as the feed. Therefore, we consider only a total of 8 equilibrium stages. In the McCabe-Thiele diagram below, the stripping section operating line is coincident with the 45o line. The 8 stages are stepped off from the top at y = 0.36 and x = 0.36. The mole fraction of benzene in the reboiler is found to be 0.008 from the low concentration region plot. Therefore, the corresponding toluene mole fraction = 0.992, which is quite pure. Note that 2 diagrams below are used, with the second, covering the stages at the bottom of the column in the low concentration region. To obtain accuracy in this low-end region, the vaporliquid equilibrium data for x = 0.1 and x = 0.2 were fitted to a quadratic equation passing through the origin, with the result, y = 2.35x - 2.5x2 . Analysis: (a) (continued) Exercise 7.18 (continued) Analysis: (continued) (b) F = Feed rate = L = 1.25(100) = 125 kmol/h. V = Vapor rate = 100 kmol/h. Therefore, with no condenser and constant molar overflow, B = F -V = 125 - 100 = 25 kmol/h. The slope of the stripping section operating line = L / V = 125/100 = 1.25. The q-line is vertical for a saturated liquid feed, passing through xF = 0.36. The stripping section operating line is positioned by trial and error so that 8 + 1 equilibrium stages will be stepped off. Because part (a) shows that a very low mole fraction of benzene is obtained in the bottoms, it is suspected that the operating line will intersect the 45o line almost at the origin. If this were true, then the operating line, with a slope of 1.25, would intersect the vertical q-line at 1.25(0.36) = 0.45. Use this as a first approximation and adjust it downward until only 9 stages can be stepped off. The final result is shown in the two plots on the next page, where the second plot is for the low concentration region. As seen, the mole percent of benzene in the bottoms = 0.004, giving 99.6 mol% toluene in the bottoms. The overhead vapor contains 0.449 mole fraction benzene. By material balances, F = 125 = D + B and 0.36F = 0.449D + 0.004B, D = 100.0 and B = 25.0 kmol/h. Exercise 7.18 (continued) Analysis: (b) (continued) Exercise 7.18 (continued) Low-concentration Region Exercise 7.19 Subject: Normal and abnormal operation of a distillation column separating a methanol water mixture at 101 kPa. Given: Column with 7 theoretical plates, a total condenser, and a partial reboiler. A feed of 100 kmol/h of 50 mol% methanol in water is sent to plate 3 from the bottom. During normal operation, distillate is 90 mol% methanol and bottoms is 5 mol% methanol, with a reflux rate of 1 mole per mole distillate. During abnormal operation, the following data are obtained: Stream kmol/h mol% methanol Feed 100 51 Bottoms 62 12 Distillate 53 80 Reflux 94 Vapor-liquid equilibrium data are given at 101 kPa, where methanol is the more volatile species. Assumptions: Constant molar overflow. Find: Most probable cause for abnormal operation. Recommended further tests. If 90 mol% methanol distillate could be obtained by increasing the reflux ratio for a constant vapor rate. Analysis: First determine whether the normal operation can be verified by the McCabe-Thiele method. With L/D = R = 1, from Eq. (7-7), the slope of the rectifying operating line = L/V = R/(1 + R) = 1/2 = 0.5. Also, xF = 0.5, xD = 0.90, and xB = 0.05. What is not known is the phase condition of the feed. If a saturated liquid feed is assumed, giving a vertical q-line as shown in the plot below, stepping stages up from the bottom, with the feed stage to plate 3 from the bottom, less than 2 theoretical plates are needed in the rectifying section, while 4 are present. The construction is shown on the next page. Therefore, it appears that the feed is not a saturated liquid, but is partially vaporized. Analysis: Normal Operation (continued) Exercise 7.19 (continued) By trial and error, using q-lines of various slopes, the following McCabe-Thiele diagram is consistent with the given data. It shows a q-line with a slope of -0.34. slope = q /(q - 1) Therefore, q = slope/(slope-1)=-0.34/(-0.34-1.0)=0.25 From Eq. (7-19), molar fraction vaporized = 1 - q = 1 - 0.25 = 0.75 Analysis: Normal operation (continued) Exercise 7.19 (continued) The material balance for the normal operation is as follows, using the overall balances, F = 100 = D + B and 0.5F = 0.5(100) = 50 = xDD + xBB = 0.90D + 0.05B. Stream Feed Bottoms Distillate Reflux kmol/h 100 47.06 52.94 52.94 mol% methanol 50 5 90 90 Analysis: Abnormal operation For the abnormal operation, first check the overall total material balance using the given data. F = 100 kmol/h. D + B = 53 + 62 = 115 kmol/h. Therefore, it appears that we have 115 - 100 = 15 kmol/h more flow out of the distillation system. Now check the methanol overall material balance using the given data. Methanol flow rate in = 0.51(100) = 51 kmol/h. Methanol flow rate out = 0.80(53) + 0.12(62) = 49.84 kmol/h. Therefore, the methanol balance is close, with only about a 2% discrepancy. Now check the water overall material balance using the given data. Water flow in = 0.49(100) = 49 kmol/h. Water flow out = 0.20(53) + 0.88(62) = 65.16 kmol/h. Therefore, we have 65.16 - 49 = 16.16 kmol/h more water out than in. This is a significant discrepancy. It appears certain that water is leaking into the distillation system. Two possibilities are: (1) leakage of condenser cooling water into the condensate, or (2) leakage of reboiler steam into the boilup vapor. A reboiler steam leak may not be serious because the steam might not get to the top of the column to dilute the methanol product. A condenser cooling water leak could be very serious because part of it would end up in the distillate, thereby diluting the methanol product. Because of the impure methanol distillate for the abnormal operation, it appears that a condenser cooling water leak is the fault. Check this next. We note that the distillate flow rate for the abnormal operation is almost exactly the same as that for the normal operation. A flow rate equal to that of he leakage passes out the bottom of the column. In normal operation, the water passing out in the distillate = 0.1(53) = 5.3 kmol/h, while for the abnormal operation, the water passing out in the distillate = 0.2(53) = 10.6 kmol/h. Thus, an additional 5.3 kmol/h of water leaves in the distillate. For the abnormal operation, the overhead vapor rate = 53 + 94 = 147 kmol/h and, therefore, 53/147 x 100% = 36% of the overhead vapor (total condensate) is distillate. Thus, if 15 kmol/h of water leaked into the overhead vapor, then, we would expect 0.36(15) = 5.4 kmol/h would be expected to leave with the distillate. This compares very well with the 5.3 kmol/h additional water calculated above by material balance. If the degree of fractionation within the column is about the same as for the normal operation, it could be concluded that a condenser cooling water leak is to blame. To check the cooling water leak, could meter the cooling water in and out of the condenser and see if there is a difference. If the vapor rate is kept constant and the reflux rate is increased, then the distillate rate must be decreased. Assume a vapor rate of 147 kmol/h, with 30 kmol/h to distillate and 117 kmol/h to reflux. Then, 30/117 x 100% = 25.6% of the overhead vapor is distillate. Therefore, the water leak to the distillate would be 0.256(15) = 3.84 kmol/h. If the fractionation were otherwise the same as for normal operation so that the overhead vapor was 90 mol% methanol, the dilution with leakage would result in 0.1(30 - 3.84) + 3.84 = 6.46 kmol/h of water in 30 kmol/h. Thus, methanol purity = (30 - 6.46)/30 x 100% = 78.5 mol%. However, the higher reflux ratio would increase the fractionation, so as to increase the purity above this value. A further increase in fractionation could be achieved, if the feed were condensed to a saturated liquid and additional heat was transferred in the reboiler. But, even if a pure methanol overhead vapor were achieved, the methanol purity after dilution with the water leakage would be : (30 - 3.84)/30 x 100% = 87.2 mol% methanol. Must eliminate the leak. Exercise 7.19 (continued) Exercise 7.20 Effect on reflux and boilup compositions of reducing the feed rate to a distillation Subject: column when the reflux and boilup rate are held constant Given: Column with 3 theoretical plates, a total condenser, and a partial reboiler. Feed is a saturated liquid of 50 mol% A and 50 mol% B, fed to the bottom tray. At a feed rate of 100 kmol/h, desired products of distillate with 90 mol% A and bottoms of 20 mol% A can be achieved, when a reflux corresponding to L/V = 0.75 is used. Relative volatility of A to B is constant at 3.0. Assumptions: Constant molar overflow. Saturated liquid reflux. Find: Compositions of reflux and boilup when feed rate is inadvertently reduced to 25 kmol/h, while keeping the reflux and boilup flow rates constant. Analysis: First, verify the separation for a feed rate of 100 kmol/h. This is shown in the McCabe-Thiele plot below in terms of mole fractions of A, the more volatile component. The equilibrium curve is computed from Eq. (7-3), 3x x y= = 1 + x ( - 1) 1 + 2 x The given mole fractions are: xF = 0.50 xD = 0.90 xB = 0.20 The rectification section operating line has a slope of 0.75 and passes through point {0.9, 0.9}. The q-line is vertical at x = 0.5. The plot shows almost perfect agreement with the desired separation for a feed rate of 100 kmol/h. Analysis: (continued) Exercise 7.20 (continued) For the base case of F = 100 kmol/h, the material balance equations are F = D + B and xFF = 50 = xDD + xBB = 0.90D + 0.20B. Solving, these equations, along with V = L + D, V/L = 0.75, V = V , and L = L + F , gives the following results: Stream Feed Distillate Bottoms Reflux, L Overhead vapor, V Liquid to reboiler, L Vapor from reboiler, V Flow rate, kmol/h 100.00 42.86 57.14 128.58 171.44 228.58 171.44 Mol% A 50 90 20 Mol% B 50 10 80 When the feed rate is reduced to 25 kmol/h, the reflux rate, L, is maintained at 128.58 kmol/h and the boilup, V , is maintained at 171.44 kmol/h. Therefore by material balances, L = L + F = 128.58 + 25 = 153.58 and B = L -V = 153.58 - 171.44 = -17.86 kmol/h. This is impossible. Therefore, the column can not be operated with the same boilup rate. That rate would have to be reduced to achieve a desired bottoms rate, e. g. the 57.14 % of the feed, as in the base case or 0.5714(25) = 14.29 kmol/h. If this were done, we would now have, V = L - B = 153.58 - 14.29 = 139.29 kmol/h = V. Thus, in the rectifying section, L/V = 128.58/139.29 = 0.923 and in the stripping section, L /V = 153.58/139.29 = 1.103. The resulting distillate and bottoms compositions are determined by positioning the operating lines so that 3 stages + a reboiler can be stepped off. The result is shown below, where the mole fractions of A are 0.93 in the distillate and reflux, 0.18 in the bottoms, and 0.38 in the reboiler vapor. Analysis: (continued) Exercise 7.20 (continued) Exercise 7.21 Subject: Distillation of a saturated vapor of maleic anhydride (A) and benzoic acid (B) under vacuum at 13.3 kPa. Given: Feed contains 90 mol% A and 10 mol% B. Distillate to contain 99.5 mol% anhydride and bottoms to contain 0.5 mol% acid. Vapor pressure data. Assumptions: Raoult's law to compute K-values from vapor pressure data. Constant molar overflow. Find: Number of theoretical plates needed if a reflux ratio, L/D = 1.6 times minimum. Analysis: First compute an equilibrium y, x curve using Raoult's law with the vapor pressure data. Eq. (2-44 ) applies, as well as the sum of the mole fractions in the phases in equilibrium. Thus, Ps T Ps T y y KA = A = A , KB = B = B (1, 2) xA P xB P yA + y B = 1 , xA + x B = 1 (3, 4) Equations (1) to (4) can be reduced to the following equations for the mole fractions of maleic anhydride (A) in terms of the K-values: 1 - KB (5, 6) , y A = KA x A KA - KB If the given vapor pressure data are fitted to Antoine equations, we obtain: xA = PAs = exp 16.6541 - PBs = exp 23.0155 - 4088.12 T + 203.924 9336.97 T + 321434 . (7) (8) Where vapor pressure is in torr and temperature is in oC. Solving, Eqs. (1) to (8), In solving the equations, P = 13.3 kPa or 99.8 torr. The results are tabulated on the next page. Below the table is a McCabe-Thiele plot of yA versus xA for determining the minimum reflux for xD = 0.995 and a horizontal q-line at y = 0.90, which intersects the equilibrium curve at x = 0.572. Therefore, the slope of the rectifying section operating line at minimum reflux is (L/V)min = (0.995 - 0.90)/(0.995 - 0.572) = 0.225. From a rearrangement of Eq. (7-7), Rmin = (L/V)min/[1 - (L/V)min] = 0.225/(1-0.225) = 0.290. Therefore, the reflux ratio for operation = 1.6Rmin = 1.6(0.290) = 0.464. Analysis: (continued) Exercise 7.21 (continued) T, oC Ps of A, torr Ps of B, torr 135.3 99.8 13.1 136.5 104.1 13.8 137.8 108.9 14.6 140.3 118.8 16.3 142.8 129.5 18.2 145.3 140.9 20.3 147.8 153.1 22.6 166.2 25.1 150.3 152.8 180.2 27.9 211.1 34.2 157.8 160.3 228.1 37.8 162.8 246.3 41.8 165.3 265.6 46.2 167.8 286.1 50.9 307.9 56.1 170.3 172.8 331.1 61.8 175.3 355.6 68.0 177.8 381.6 74.7 180.3 409.2 81.9 438.3 89.9 182.8 185.3 469.1 98.5 185.7 474.2 99.9 KA 1.000 1.043 1.092 1.191 1.297 1.411 1.534 1.665 1.805 2.115 2.286 2.468 2.661 2.867 3.085 3.317 3.563 3.824 4.100 4.392 4.701 4.751 KB 0.131 0.138 0.147 0.164 0.183 0.203 0.226 0.251 0.279 0.343 0.379 0.419 0.463 0.510 0.562 0.619 0.681 0.748 0.821 0.901 0.987 1.001 xA 1.000 0.953 0.903 0.814 0.733 0.659 0.592 0.530 0.472 0.371 0.326 0.284 0.244 0.208 0.173 0.141 0.111 0.082 0.055 0.028 0.004 0.000 yA 1.000 0.993 0.986 0.970 0.951 0.931 0.908 0.882 0.853 0.784 0.744 0.700 0.650 0.596 0.535 0.468 0.395 0.313 0.224 0.125 0.017 0.000 Analysis: (continued) Exercise 7.21 (continued) Exercise 7.21 (continued) Analysis: (continued) Now determine the tray requirements for actual operation. Using Eq. (7-7), with the operating reflux ratio of 0.464, L/V = R/(1 + R) = 0.464/(1 + 0.464) = 0.317. Because such high purity distillate and bottoms products are to obtained, use 3 McCabe-Thiele diagrams. The first diagram is for the high purity region of component A from y and x = 0.9 to 1.0. The operating line for the rectifying section begins at {0.995, 0.995} and, with a slope of 0.317, intersects the vertical axis for x = 0.90 at y = 0.965. The entire region is covered in the second diagram, the feed stage is located optimally. The low concentration region is covered in the third diagram. From these three diagrams, it is seen that 8 theoretical plates plus a partial reboiler are needed. The feed is sent to plate 4 from the top. Analysis: (continued) Exercise 7.21 (continued) Analysis: (continued) Exercise 7.21 (continued) Exercise 7.22 Subject: Distillation of a mixture of A and B based on boilup, rather than reflux, requirements. Given: A bubble-point feed mixture of 5 mol% A and 95 mol% B. Distillate to contain 35 mol% A and a bottoms to contain 0.2 mol% A. Relative volatility, A,B = 6 = a constant. Column equipped with partial condenser and partial reboiler. Assumptions: Constant molar overflow. Find: Using algebraic methods, (a) Minimum number of equilibrium stages. (b) Minimum boilup ratio, VB = V / B . (c) Number of equilibrium stages for a boilup ratio = 1.2 times minimum. Analysis: From a rearrangement of the equilibrium equation, Eq. (7-3), y y x= = (1) y + (1 - y ) 6 - 5 y (a) For minimum stages, have total reflux, so that y = x for passing streams. Begin calculations from the top. yD = y1 = 0.35. From Eq. (1), x1 = 0.35/[6 - 5(0.35)] = 0.0824. Therefore, y2 = x1 = 0.0824. From Eq. (1), x2 = 0.0824/[6-5(0.0824)] = 0.0147. Therefore, y3 = x2 = 0.0147. From Eq. (1), x3 = 0.0147/[6-5(0.0147)] = 0.0025. This is close to but not quite equal to the desired value of 0.002. Thus, we need just slightly more than 3 minimum equilibrium stages. (b) For minimum boilup ratio, the stripping section operating line connects the two points for {y, x} of {0.002, 0.002} and {y in equilibrium with x = 0.05}. From a rearrangement of Eq. (1), the y in equilibrium with x = 0.05 is: y = x/[1 + x( - 1)] = 6(0.05)/[1 + 0.05(6 - 1)] = 0.24. The slope of the operating line = ( L / V ) = (0.24 0.002)/(0.05 - 0.002) = 4.96. From a rearrangement of Eq. (7-12), (VB)min = 1/ [( L / V ) - 1] = 1/(4.96 - 1) = 0.253. (c) The boilup ratio = VB = 1.2(0.253) = 0.3036. From Eq. (7-12), the slope of the stripping section operating line = L / V = (VB + 1)/VB = (0.3036 + 1)/0.3036 = 4.294. This line intersects the vertical q-line (xF = 0.05) at 0.002 + 4.294(0.05 - 0.002) = 0.2081. Therefore, the slope of the rectifying line = L/V = (0.35 - 0.2081)/(0.35 - 0.05) = 0.4730. From a rearrangement of Eq. (7-8), R = (L/V)/[1 - (L/V)] = 0.473/(1 - 0.473) = 0.8975. The equation for the rectifying section operating line, using Eq. ((7-9), with a modification for a partial condenser as determined from Fig. 7.18, is, y n +1 = R 1 xn + y D = 0.473xn + 0.184 5 R +1 R +1 (2) Analysis: (c) (continued) Exercise 7.22 (continued) The equation for the stripping section operating line, using Eq. (7-12) is, 1 V +1 y m +1 = B xm - x B = 4.294 xm - 0.00659 VB VB (3) We can now calculate stage by stage down from the top, starting from yD = 0.35, alternating between the equilibrium curve, Eq. (1) and the appropriate operating line, Eq. (2) or (3). We begin using Eq. (2), but switch to Eq. (3), when x < xF = 0.05. The calculations are terminated when x < xB = 0.002. The calculations can be done with a spreadsheet, with the following results, given as mole fractions of A leaving an equilibrium stage. The optimal feed stage is the top plate. Equilibrium stage Partial condenser 1 2 3 4 5 6 7 8 Partial reboiler yA 0.350 0.223 0.190 0.155 0.121 0.0894 0.0626 0.0407 0.0235 0.0106 xA 0.0824 0.0458 0.0376 0.0296 0.0224 0.0161 0.0110 0.00701 0.00400 0.00178 The calculations show that besides the partial condenser and partial reboiler, 8 equilibrium stages are needed in the column. Exercise 7.23 Subject: Distillation of a mixture of methanol and water with a subcooled liquid feed. Given: Liquid feed of 14,460 kg/h methanol and 10,440 kg/h water at q = 1.12. Distillate of 99 mol% methanol and a bottoms of 99 mol% water are desired. Column has a total condenser and a partial reboiler. Operation at 1 atm with a reflux ratio of L/D = R = 1.0. Vapor-liquid equilibrium data in Exercise of 7.19. Assumptions: Constant molar overflow. Find: Feed stage location and number of equilibrium stages. Analysis: First, determine feed composition in mol%. Using molecular weights of 32.04 for methanol and 18.02 for water, Component Methanol Water Total: kg/h 14,460 10,440 24,900 kmol/h 451.3 579.4 1,030.7 Mol% 43.79 56.21 100.00 Using the vapor-liquid equilibrium data, a y-x plot for the McCabe-Thiele method is made and smoothed with a spreadsheet, noting that methanol is the more volatile. Therefore, xF = 0.4379, xD = 0.99, xB = 0.01. From Eq. (7-26), slope of q-line = q/(q - 1) = 1.12/(1.12 - 1) = 9.333. Therefore, on the McCabeThiele diagram, a line is drawn with a slope of 9.333 that intersects the point {0.438,0.438}. From Eq. (7-7), the slope of the rectifying section operating line = L/V = R/(R + 1) = 1/(1 + 1) = 0.5. This line is drawn on the McCabe-Thiele diagram with a slope of 0.5 that intersects the point {0.99,0.99}. The stripping section operating line is drawn to intersect the point {0.01,0.01} and the point where the q-line intersects the rectifying section operating line. In the McCabe-Thiele graphs below, the first for the high mole-fraction region, it is seen that the specified reflux ratio is, fortunately, above the minimum value for the specified distillate mole fraction. From the plots, it is seen that 20 theoretical stages plus a partial reboiler are needed. The optimal feed stage is number 17 down from the top. Analysis: (continued) Exercise 7.23 (continued) Analysis: (continued) Exercise 7.23 (continued) Exercise 7.24 Subject: condenser. Partial separation of a benzene-toluene mixture with a partial reboiler and a partial Given: Saturated liquid feed of 69.4 mol% benzene (B) in toluene (T) fed to a partial reboiler. Vapor from the reboiler passes to a partial condenser. Vapor from the partial condenser passes to a total condenser. Reflux from the partial condenser is sent to the partial reboiler. Distillate is to contain 90 mol% benzene (yD = 0.9) at a rate of 25 moles per 100 moles of feed. The relative volatility of benzene with respect to toluene = = 2.5. Find: Moles of vapor generated in the reboiler per 100 moles of feed by analytical and graphical methods. Analysis: First compute overall material balance. The total condenser need not be considered. Basis: 100 moles of feed. Overall total material balance: F = 100 = D + B = 25 + B (1) Solving Eq. (1), B = 75 moles Overall benzene material balance: FxF = 69.4 = DyD + BxB = 25(0.9) + 75xB (2) Solving Eq. (2), xB = 0.625 Analytical Method: Write material balances around the partial reboiler and partial condenser, using subscripts B for streams leaving the reboiler and D for streams leaving the partial condenser. Partial reboiler: Total material balance: F + LD = B + VB or 100 + LD = 75 + VB (3) Benzene material balance: FxF + LDxD = BxB + VByB (4) or 69.4 + LDxD = 75(0.625) + VByB = 46.9 + VByB Partial condenser: Total material balance: VB = D + LD = 75 + LD (5) Benzene material balance: VByB = DyD + LDxD = 75(0.9) + LDxD = 67.5 + LDxD (6) Assume equilibrium in the partial condenser and partial reboiler. Using the given with its definition in Eq. (7-2). For the partial condenser, = 2.5 = yD (1 - xD)/xD(1 - yD) = 0.9(1 - xD)/xD(1 -0.9) = 9(1 - xD)/xD (7) Solving Eq. (7), xD = 0.783 For the partial reboiler, = 2.5 = yB (1 - xB)/xB(1 - yB) = yB (1 - 0.625)/0.625(1 - yB) = 0.6 yB/(1 - yB) (8) Solving Eq. (8), yB = 0.806 Eqs. (3) through (6), are 4 equations in 2 unknowns, VB and LD. We only need 2 of the 4 equations. Using Eqs. (3) and (4), 100 + LD = 75 + VB (3) 69.4 + LD(0.783) = 46.9 + VB(0.806) (9) Solving Eqs. (3) and (9), LD = 98 moles /100 moles of feed and VB = 123 moles/100 moles of feed Analysis: (continued) Graphical Method: On the McCabe-Thiele diagram below, the equilibrium curve is obtained from Eq. (7-3), x 2.5x y= = 1 + x ( - 1) 1 + 15x . The rectification section operating line is located, as shown, so that two equilibrium steps, one for the partial condenser and one for the partial reboiler, are stepped off between xC = 0.9 (from the total condenser) and xB = 0.625. The measured slope of the operating line = LD/VB = 0.8. Combining this with Eq. (3), VB = 125 moles/100 moles of feed, which is close to the analytical value. Exercise 7.24 (continued) Exercise 7.25 Subject: Rectification of a mixture of benzene and chlorobenzene at total reflux. Given: Feed of 100 kmol of 20 mol% benzene and 80 mol% chlorobenzene. Column has 4 theoretical plates, a total condenser, a reflux drum, and a still to vaporize the feed. At an operating pressure of 1 atm, relative volatility of benzene with respect to chlorobenzene = = 4.13. Operate at total reflux with holdups only in the reflux drum and the still. Want liquid in the still with 0.1 mol% benzene. Assumptions: Perfect mixing to give uniform compositions in the reflux drum and the still. Find: Moles of liquid in the still at steady state. Analysis: This exercise can be solved analytically or graphically. Since benzene is the more volatile component, Eq. (7-3) gives the equilibrium relation at the still or any of the 4 plates, n, as, yn = xn 4.13xn = 1 + xn ( - 1) 1 + 313xn . (1) At total reflux, numbering stages up from the bottom, xn+1 = yn (2) Analytical Method: Start at the bottom, stage 1, with x1 = 0.001. Solve for y1 (vapor leaving the still) with Eq. (1). Then, from Eq. (2), x2 = y1. Continue in this manner, solving alternately Eq. (1) and then Eq. (2), until y5 (vapor leaving the top plate) is reached. Then, x in the reflux drum = y5. The results from a spreadsheet are, Equilbrium stage 1 (still) 2 (bottom plate 3 4 5 (top plate) reflux drum x 0.00100 0.00412 0.01679 0.06587 0.22554 0.54603 y 0.00412 0.01679 0.06587 0.22554 0.54603 By overall total material balance, F = 100 = D + B (3) By overall benzene material balance, FxF = (100)(0.20) = 20 = DxD + BxB = D(0.54603) + B(0.001) (4) Solving Eqs. (3) and (4), D = distillate in reflux drum = 36.51 kmoles B = bottoms in still = 63.49 kmoles Exercise 7.25 (continued) Graphical Method: On the McCabe-Thiele plot on the next page, the equilibrium curve is computed from Eq. (1). The rectification section operating line is the 45o line. Five stages are stepped off from the bottoms of xB = 0.001. The resulting y5 is essentially the same as that for the analytical method. Thus, again, B = bottoms in still = 63.49 kmoles. Analysis: Graphical method (continued) Exercise 7.25 (continued) Exercise 7.26 Subject: Distillation of acetone and isopropanol, taking into account tray efficiency. Given: Saturated liquid feed of 50 mol% acetone and 50 mol% isopropanol. Column is equipped with a total condenser, and a partial reboiler. Reflux ratio, L/D, is 0.5. Murphree vapor efficiency = 50%. Vapor-liquid equilibrium data at 1 atm are given, with acetone being the more volatile component. Assumptions: Constant molar overflow. Find: Number of actual trays required to achieve a distillate of 80 mol% acetone and a bottoms of 25 mol% acetone, inserting the feed at the optimal location. Analysis: In the McCabe-Thiele diagram below, the equilibrium curve is plotted from the given data. The q-line is vertical, passing through x = 0.5. The rectification operating line has a slope, L/V, from Eq. (7-9), of R/(R + 1) = 0.5/1.5 = 0.333. This operating line passes through the point, y = 0.8, x = 0.8. From Eq. 7-9), the equation for the rectifying section operating line is, y= R 1 x+ x D = 0.333x + 0.667(0.8) = 0.333x + 0.533 R +1 R +1 From this equation, the intersection of the rectifying section operating line and the vertical q-line is at y = 0.333(0.5) + 0.533 = 0.70. The stripping section operating line passes through the {y, x} points {0.70, 0.50} and {0.25, 0.25}, giving it the equation, y = 1.80x-0.20. Except for the reboiler stage, the stages are stepped off from an efficiency line, which for a Murphree vapor efficiency of 0.5 is positioned vertically half way between the equilibrium curve and the operating line, as governed by Eq. (7-41), EMV = 0.5 = (yn - yn+1)/(yn* - yn+1), where yn+1 is the location on the operating line, yn is the location on the efficiency line, and yn* is the location on the equilibrium line. However, the reboiler is assumed to have a 100% efficiency. As seen in the plot, just over 8 trays are required plus the partial reboiler. The feed plate is 4 from the top. Analysis (continued) Exercise 7.26 (continued) Exercise 7.27 Subject: Distillation of carbon disulfide and carbon tetrachloride. Given: Partially vaporized feed (q = 0.5) of 40 mol% CS2. Operation with a reflux ratio, L/D, of 4 and a Murphree vapor efficiency of 80%. Partial reboiler and total condenser. Vapor-liquid equilibrium data. Assumptions: Constant molar overflow. Find: For a distillate of 95 mol% CS2 and a bottoms of 5 mol% CS2, determine: (a) Minimum reflux ratio, minimum boilup ratio, and minimum number of stages. (b) Number of trays. Analysis: (a) In the McCabe-Thiele plot on the next page, rectifying section and stripping section operating lines are shown for determining minimum reflux and boilup ratios. Note that the q-line has a slope given in Eq. (7-26) as q/(q - 1) = 0.5/(0.5 - 1) = -1 and intersects the point {0.4, 0.4}. The rectifying section operating line intersects the point {0.95, 0.95} and the point where the equilibrium curve and the q-line intersect. The slope of that line is measured to be L/V = 0.642. From Eq. (7-9), L/V = R/(R + 1). Rearranging, Rmin = (L/V)/[1 - (L/V)] = 0.642/(1 - 0.642) = 1.79 The stripping section operating line intersects the point {0.05, 0.05} and the point where the equilibrium curve and the q-line intersect. The slope of that line is measured to be L / V = 2.043. From Eq. (7-14), L / V = (VB + 1)/VB. Rearranging, V 1 1 = = = 0.959 (VB )min = B min ( L / V ) - 1 2.043 - 1 The McCabe-Thiele plot for minimum stages at total reflux is also shown on the next page. The operating lines are coincident with the 45o line. It is seen that 6 equilibrium stages are needed. (b) For a reflux ratio, R = L/D, of 4, the slope of the rectifying section operating line from Eq. (7-9) is L/V = R/(R + 1) = 4/5 = 0.8. This line and the stripping section operating line are shown on the third McCabe-Thiele diagram below. Except for the reboiler stage, the stages are stepped off from an efficiency line, which for a Murphree vapor efficiency of 0.8 is positioned 80% of the vertical distance from the operating line to the equilibrium curve, as governed by Eq. (7-41), EMV = 0.8 = (yn - yn+1)/(yn* - yn+1), where yn+1 is the location on the operating line, yn is the location on the efficiency line, and yn* is the location on the equilibrium line. However, the reboiler is assumed to have a 100% efficiency. As seen in the plot, just over 9 trays are required plus the partial reboiler. Call it 10 trays plus the reboiler. The feed plate is 7 from the top. Analysis: (a) (continued) Exercise 7.27 (continued) Analysis: (a) (continued) Exercise 7.27 (continued) Analysis: (b) (continued) Exercise 7.27 (continued) Exercise 7.28 Subject: Preliminary design calculations for the distillation of a benzene-toluene mixture. Given: Bubble-point feed of 50 mol% benzene and 50 mol% toluene. Equipment to include a partial reboiler, total condenser, and a bubble-cap tray column with an overall plate efficiency of 65%. Column to operate at 1 atm to produce a distillate of 95 mol% benzene and a bottoms of 95 mol% toluene. Vapor-liquid equilibrium data from Exercise 7.13. Enthalpy data for reboiler. Assumptions: Constant molar overflow and saturated liquid reflux. Find: (a) (b) (c) (d) (e) (f) Minimum reflux ratio (infinite stages). Minimum number of actual plates (total reflux). Number of actual plates for R = 1.5 Rmin. Kilograms per hour of products for a feed of 907.3 kg/h. Kg/h of saturated steam at 273.7 kPa for reboiler heat duty using given enthalpy data. Rigorous enthalpy balance around the reboiler. Analysis: McCabe-Thiele plots are made in terms of benzene mole fractions, since benzene is the more volatile component. The equilibrium curve is plotted from the data in Exercise 7.13. (a) For a saturated liquid feed, minimum reflux corresponds to a pinch point located at the intersection of a vertical q-line passing through xF = 0.5 and the equilibrium curve as shown in the McCabe-Thiele diagram below. From the equilibrium data, this intersection is at y = 0.72 and x = 0.5. Then, the slope of the rectifying section operating line, (L/V)min is (0.95 - 0.72)/(0.95 - 0.50) = 0.511. From a rearrangement of Eq. (7-7), Rmin = (L/V)min /[1 - (L/V)min ] = 0.511/(1 0.511) = 1.045. (b) The McCabe-Thiele plot for minimum stages at total reflux is shown below. The operating lines are coincident with the 45o line. Equilibrium stages are stepped off starting from xB = 0.05 to xD =0.95. It is seen that just less than 7 equilibrium stages are needed. Call it Nt = 7. From Eq. (6-21), for an overall plate efficiency of 65%, i.e. Eo = 0.65, the actual minimum number of plates = Na = Nt / Eo = 7/0.65 = 10.8. (c) Operating reflux ratio = R = 1.5 Rmin = 1.5(1.045) = 1.57. From Eq. (7-7), the slope of the operating line for the rectifying section = L/V = R/(1 + R) = 1.57(1 + 1.57) = 0.611. On the McCabe-Thiele diagram on the next page, the rectifying section operating line has this slope and passes through the point, y=0.95, x=0.95. the stripping section operating line passes through the point, y=0.05, x=0.05 and intersects the vertical q-line at the point where the rectifying section operating line intersects the q-line. As seen, the equilibrium stages are stepped off starting at the top, with a switch from the rectifying section to the stripping section to minimize the number of stages and, thus, locating the optimal feed stage. The result is just over 10 equilibrium stages plus a partial reboiler. Call it 11 equilibrium stages plus a partial reboiler. Applying Eq. (6-21), Na = 11/0.65 = 16.9 or 17 actual plates plus the partial reboiler. Analysis: (continued) Exercise 7.28 (continued) Analysis: (continued) Exercise 7.28 (continued) Analysis: (continued) Exercise 7.28 (continued) Analysis: (continued) (d) MW of benzene = 78.11. MW of toluene = 92.14. Let F = kmol/h of feed. Then by mass material balance with an equimolar feed, 0.5F (78.11) + 0.5F(92.14) = 907.3 Solving, F = 10.66 kmol/h. For the equimolar feed, the component flow rates in the feed are: 5.33 kmol/h each for benzene and toluene Next calculate the distillate and bottoms flow rates from, overall total mole balance: F = nF = 10.66 = D + B (1) overall benzene mole balance: FxF = 5.33 = 0.95D + 0.05B (2) Solving Eqs. (1) and (2), D = 5.33 kmol/h and B = 5.33 kmol/h Therefore in terms of mass flow rates, total total distillate rate is, mD = 0.95(5.33)(78.11) + 0.05(5.33)(92.14) = 420.1 kg/h Therefore the bottoms rate = mB = 907.3 - 420.1 = 487.2 kg/h (e) First compute the kmol/h of vapor leaving the reboiler, using the assumption of constant molar overflow. From part (c), the reflux ratio = 1.57. Therefore, the reflux rate = 1.57(5.33) = 8.37 kmol/h. Below the feed stage, the liquid rate = 8.37 + 10.66 = 19.03 kmol/h. The vapor rate leaving the reboiler = 19.03 - 5.33 = 13.70 kmol/h. From the plot above, the composition of the reboiler vapor = 12 mol% benzene. Neglecting the sensible heat and using the enthalpy data given, after converting from Btu/lbmol to kJ/kmol, the reboiler heat duty is, QR = 2.324[0.12(13.7)(18,130 - 4,900) + 0.88(13.7)(21,830 - 8,080)] = 436,000 kJ/kmol. From Perry's Handbook, latent heat of vaporization of steam at 273.7 kPa (404 K) = 2,172 kJ/kg Therefore, we need 436,000/2,172 = 200.7 kmol/h or 3,616 kg/h. (f) A rigorous enthalpy balance around the reboiler takes into account the sensible heat effect since the temperature of the liquid entering the reboiler is not the same as the temperatures of the equilibrium liquid and vapor leaving the reboiler. Let N = conditions leaving the reboiler and N-1 be the conditions leaving the stage above the reboiler. Then, QR = VN HVN + LN H LN - LN -1 H LN -1 Note that in the simplified enthalpy balance of part (e), the following equation was applied, QR = VN HVN - H LN Since, VN = LN - LN -1 , this is equivalent to assuming H LN = H LN -1 Exercise 7.28 (continued) Exercise 7.29 Subject: Preliminary design for the distillation of a mixture of ethanol and water at 1 atm. Given: Bubble-point feed containing 20 mol% ethanol in water. Unit consisting of a perforated-tray column, partial reboiler, and total condenser. Distillate to contain 85 mol% alcohol and a 97% recovery of alcohol. Vapor-liquid equilibrium data. Assumptions: Constant molar overflow. Find: (a) (b) (c) (d) Molar concentrations in the bottoms product. Minimum values of L/V, L/D, and VB/B. Minimum number of equilibrium stages and actual plates for Eo = 0.55. Number of actual plates for L/V = 0.80. Analysis: From the vapor-liquid equilibrium data for 1 atm., it is seen that ethanol is more volatile than water for ethanol mole fractions in the liquid from 0 to 0.8943, which is the azeotrope concentration. The distillate composition is within this region. (a) Take a basis of F = 100 kmol/h. (1) Overall total material balance: F = 100 = D + B (2) Ethanol recovery: 0.97FxF = 0.97(100)(0.20) = 19.4 = DxD = 0.85D Solving Eq. (2), D = 22.82 kmol/h. From Eq. (1), B = 100 - 22.82 = 77.18 kmol/h Ethanol in bottoms = 20 - 19.4 = 0.6 kmol/h Therefore, ethanol mole fraction in bottoms = 0.6/77.18 = 0.00777 Water mole fraction in bottoms = 1.0 - 0.00777 = 0.99223 (b) In the McCabe-Thiele diagram on the next page, the given equilibrium data are plotted. For a bubble-point liquid feed, the q-line is vertical at xF = 0.2. The minimum reflux in terms of L/V is obtained from the slope of the rectifying section operating line, which passes through the point, y = xD = 0.85 and is tangent to the equilibrium curve, rather than being drawn through the intersection of the q-line and the equilibrium curve because that would cause the operating line to mistakenly cross over the equilibrium curve. The slope of the operating line = (L/V)min = 0.65. From Eq. (7-27), Rmin = (L/D)min = 0.65/(1-0.65) = 1.86. The liquid rate in the rectifying section = L = 1.86D = 1.86(22.82) = 42.44 kmol/h. Below the feed plate, L = L + F = 42.44 + 100 = 142.44 kmol/h. Vapor rate from the reboiler = VB = L - B = 142.44 - 77.18 = 65.26 kmol/h. Therefore, boilup ratio = VB/B = 65.26/77.18 = 0.846. (c) In the second McCabe-Thiele diagram on the next page, the minimum number of stages is determined by stepping off stages between the equilibrium curve and the 45o line (total reflux) from the points 0.85 and 0.00777 on the 45o line. The result is approximately 10 minimum equilibrium stages. For a stage efficiency of 0.55, using Eq. (6-21), Na = Nt/Eo = 10/0.55 = 18.2 minimum plates. Analysis: (b and c) (continued) Exercise 7.29 (continued) Analysis: (b and c) (continued) Exercise 7.29 (continued) Analysis: (continued) (d) For an operating reflux ratio = L/V = 0.8, the reflux ratio, R = L/D = 0.8/(1-0.8) = 4. On the McCabe-Thiele diagram below, the rectifying section operating line has a slope of 0.8 and passes through the point, y=0.85, x=0.85. the stripping section operating line passes through the point, y=0.00777, x=0.0777 and intersects the vertical q-line at the point where the rectifying section operating line intersects the q-line. As seen, the equilibrium stages are stepped off starting at the top, with a switch from the rectifying section to the stripping section to minimize the number of stages and, thus, locating the optimal feed stage. The result is just less than 15 equilibrium stages. Call it 14 equilibrium stages plus a partial reboiler. Applying Eq. (6-21), Na = 14/0.55 = 25.5 or 26 actual plates plus the partial reboiler as an equilibrium stage. Exercise 7.29 (continued) Exercise 7.30 Subject: Recovery by distillation with open steam of solvent A from water in two feeds. Given: Two saturated liquid feeds, each containing 50 kmol/h of A. Feed 1 contains 40 mol% A and Feed 2 contains 60 mol% A. Unit consists of a column and a total condenser. Open steam is used in lieu of a partial reboiler. Relative volatility = 3.0 for A with respect to water. Distillate is to contain 95 mol% A with a 95% recovery. Assumptions: Constant molar overflow. Open steam enters bottom stage as saturated vapor. Both feeds enter at optimal locations. Find: For an overall plate efficiency of 70% and an R = L/D = 1.33 times minimum, determine the number of actual plates. Compute the bottoms composition. Determine analytically the location of all three operatiing lines. Analysis: The total flow rate of Feed 1 = 50/0.4 = 125 kmol/h. The total feed rate of Feed 2 = 50/0.6 = 83.3 kmol/h. The total feed rate of A = 50 + 50 = 100 kmol/h. For a recovery of 95% of A in the distillate, the flow rate of A in the distillate = 0.95(100) = 95 kmol/h. With a mole fraction of 0.95 for A in the distillate, the total flow rate of the distillate = 95/0.95 = 100 kmol/h. From Eq. (7-3) for = 3, the equilibrium mole fractions of A are related by, 3x x (1) y= = 1 + ( - 1) x 1 + 2 x Equation (1) is plotted in the McCabe-Thiele diagram on the next page. Because Feed 2 is richer in A than Feed 1, Feed 2 enters the column above Feed 1. At minimum reflux, the pinch condition will occur at either Feed 1 or Feed 2. Assume that the pinch occurs at Feed 2. For a saturated liquid feed, using Eq. (1), the upper section operating line will intersect the equilibrium curve for xF = 0.6 at y = 3(0.6)/[1 + 2(0.6)] = 0.818. Therefore, the slope of this operating line is, (L/V)min = (0.95 - 0.818)/(0.950 - 0.6) = 0.377 Correspondingly, using Eq. (7-17), R = L/D = (L/V)min/[1 - (L/V)min] = 0.377/(1 - 0.377) = 0.605 and Lmin = 0.605(100) = 60.5 kmol/h. Now check the middle section to see if the operating line there is below the equilibrium curve. The liquid rate in the middle section = L' = L + F2 = 60.5 + 83.3 = 143.8 kmol/h. The vapor rate in the middle section = V' = V = L + D = 60.5 + 100 = 160.5 kmol/h. Therefore the slope of the operating line in the middle section = L'/V' = 143.8/160.5 = 0.896. As seen in the McCabe-Thiele diagram on the next page, this operating line does not cross over the equilibrium curve. Therefore, the pinch does occur at Feed 2 (the upper feed). For an operating reflux ratio of 1.33 times minimum, L = 1.33(60.5) = 80.5 kmol/h. The vapor rate in the upper section = L + D = 80.5 + 100 = 180.5 kmol/h. Analysis: (continued) Exercise 7.30 (continued) Therefore the upper section operating line has a slope, L/V = 80.5/180.5 = 0.446 and passes through the point y = x = 0.95. It intersects the vertical q-line at x = 0.6 and for the slope of 0.446 = (0.95 - y)/(0.95 - 0.6), y = 0.794. For the middle section, L' = L + F2 = 80.5 + 83.3 = 163.8 kmol/h and V' = V = 180.5 kmol/h Therefore, the middle section operating line has a slope of L'/V' = 163.8/180.5 = 0.908 and intersects the q-line for x = 0.6 at y = 0.794. It intersects the vertical q-line at x = 0.4 and for the slope of 0.908 = (0.794 - y)/(0.6 - 0.4), y = 0.613. For the lower section, L" = L' + F1 = 163.8 + 125 = 288.8 kmol/h and V"=V' = 180.5 kmol/h Therefore, the lower section operating line has a slope of L"/V" = 288.8/180.5 = 1.60 and intersects the q-line for x = 0.4 and y = 0.613. As seen in Fig. 7.27(c), the mole fraction of A in the bottoms, xB , is determined from the intersection of the operating line for the lower section with the y-axis. Thus, 1.60 = (0.613 - 0)/(0.4 - xB). Solving, xB = 0.0169 for component A. Since the bottoms contains 5 kmol/h of A, the bottoms rate = B = 5/0.0169 = 295.9 kmol/h. Thus, the bottoms contains 290.9 kmol/h of water. But, the flow rate of water entering in the two feeds = 125 + 83.3 - 100 = 108.3 kmol/h. Therefore, the open steam flow rate = 290.9 + 5 - 108.3 = 187.6 kmol/h In the McCabe-Thiele diagram on the next page, the three operating lines are drawn and the equilibrium stages are stepped off so as to place the two feeds at their optimal locations. As seen, the number of equilibrium stages required = Nt = 14. From Eq. (6-21), for a plate efficiency of 70%, the actual number of trays = Na = Nt /Eo = 14/0.7 = 20 plates. Analysis: (continued) Exercise 7.30 (continued) Exercise 7.31 Subject: intercooler. Distillation of a mixture of n-hexane and n-octane in a column with an Given: Saturated liquid feed of 40 mol% hexane in octane. Intercooler at second stage from the top removes heat so as to condense 50 mol% of the vapor rising from the third. Distillate is to contain 95 mol% of hexane and bottoms is to contain 5 mol% of hexane. Reflux ratio, L/D, at the top, is equal to 0.5. Vapor-liquid equilibrium data for 1 atm is plotted in Fig. 4.4. Assumptions: Constant molar overflow. Total condenser and partial reboiler. Operating pressure of 1 atm. Find: (a) Equations to locate operating lines. (b) Number of equilibrium stages if optimal feed stage location is used. Analysis: First compute overall material balance. Take a basis of F = 100 kmol/h. Overall total mole balance: F = 100 = D + B (1) Overall hexane mole balance: FxF = 40 = DxD + BxB = 0.95D + 0.05B (2) Solving Eqs. (1) and (2), D = 38.9 kmol/h and B = 61.1 kmol/h (a) For a reflux ratio of 0.5, in the section above the intercooler, L = 0.5D = 19.45 kmol/h. The overhead vapor rate is V = L + D = 19.45 + 38.9 = 58.35 kmol/h. The slope of the operating line = L/V = 19.45/58.35 = 0.333. Using Eq. (7-6), the equation for the operating line y = 0.333x + DxD/V = 0.333x + (38.9)(0.95)/(58.35) = 0.333x + 0.633 (3) is, Now consider the section of stages between the intercooler at stage 2 from the top and the feed stage. Because 50 mol% of the vapor from this section is condensed at stage 2 by the intercooler, the vapor rate in this section = V' = 2V = 2(58.35) = 116.7 kmol/h. The liquid rate in this section is L' = V' - D = 116.7 - 38.9 = 77.8 kmol/h. The slope of the operating line = L'/V' = 77.8/116.7 = 0.667. In this section, by hexane material balance, yV' = xL' + xDD or, y = (L'/V')x + DxD/V' = 0.667x + (38.9)(0.95)/116.7 = 0.667x + 0.317 (4) In the section below the feed stage, for a saturated liquid feed, L"= L' + F = 77.8 + 100 = 177.8 kmol/h. The vapor rate = V" = V' = 116.7 kmol/h. The slope of the operating line = L"/V" = 177.8/116.7 = 1.524. From Eq. (7-11), y = (L"/V")x - BxB/V"= 1.524x - (61.1)(0.05)/116.7 = 1.524x - 0.026 (5) (b) A McCabe-Thiele diagram in terms of hexane, the more volatile component, is shown on the next page, where the equilibrium curve is obtained from Fig. 4.4 and the operating lines for the three sections are drawn from Eqs. (3), (4), and (5). The q-line is vertical, passing through x = 0.4. Note that the upper and middle section operating lines both pass through the point {0.95, 0.95}. The theoretical stages are stepped off starting from the top, switching to the middle section operating line after stage 2, and switching to the stripping section so as to locate the feed stage optimally. The result is just slightly less than 5 equilibrium stages or, say, 4 stages plus a partial reboiler. Analysis: (continued) Exercise 7.31 (continued) Exercise 7.32 Subject: Distillation of a mixture of ethyl alcohol and water at 1 atm using open steam instead of a reboiler. Given: 100 kmol/h of a saturated liquid feed containing 12 mol% ethyl alcohol in water. Distillate to contain 85 mol% alcohol with a recovery of 90%. Reflux ratio, L/D = 3 with saturated liquid reflux. Feed stage located optimally. Vapor-liquid equilibrium data in Exercise 7.29. Assumptions: Constant molar overflow. Total condenser. Find: (a) (b) (c) (d) Open steam requirement, kmol/h Number of equilibrium stages Optimal feed stage location. Minimum reflux ratio. Analysis: First compute material balance. Because the ethanol mole fraction in the distillate is less than that of the azeotrope (89.43 mol% in Exercise 7.29), the ethanol is always the more volatile component. The feed contains 12 kmol/h of ethanol and 88 kmol/h of water. For a 90% recovery, the distillate contains 0.9(12) = 10.8 kmol/h of ethanol. Since the distillate is 85 mol% ethanol, the total distillate rate = D = 10.8/0.85 = 12.7 kmol/h. The bottoms contains 12 - 10.8 = 1.2 kmol/h of ethanol. The distillate contains 12.7 - 10.8 = 1.9 kmol/h of water. The bottoms contains 88 - 1.9 + open steam = 89.9 + open steam in kmol/h. (a) For a reflux ratio of 3, L = 3D = 3(12.7) = 38.1 kmol/h. Overhead vapor rate = V = L + D = 38.1 + 12.7 = 50.8 kmol/h. Below the feed stage, L' = L + F = 38.1 + 100 = 138.1 kmol/h. Boilup rate = V' = V = 50.8 kmol/h = flow rate of open steam. (b) The bottoms rate = 138.1 kmol/h. The bottoms consists of 1.2 kmol/h of ethanol and 138.1 - 1.2 = 136.9 kmol/h of water. The mole fraction of ethanol in the bottoms = 1.2/138.1 = 0.0087. The McCabe-Thiele diagram is given on the next page, where the equilibrium curve is obtained from Exercise 7.29 and the q-line is vertical at x = 0.12. The rectifying section operating line passes through the point {0.85, 0.85}and has a slope, L/V = 38.1/50.8 = 0.75. The stripping section operating line has a slope, L'/V' = 138.1/50.8 = 2.72 and, as shown in Fig. 7.27(c), passes through the point x = xB = 0.0087 at y = 0. Because the stages are so crowded at the high mole fraction end, a second McCabe-Thiele diagram is shown for the region above y = x = 0.7. As shown, with the use of the two diagrams, just less than 20 equilibrium stages are needed. (c) From the first McCabe-Thiele plot, the optimal feed stage is Stage 18 from the top. (d) From the third McCabe-Thiele diagram on the next page, the minimum reflux in terms of L/V is obtained from the slope of the rectifying section operating line, which passes through the point, y = xD = 0.85 and is tangent to the equilibrium curve, rather than being drawn through the intersection of the q-line and the equilibrium curve because that would cause the operating line to mistakenly cross over the equilibrium curve. The slope of the operating line = (L/V)min = 0.667. From Eq. (7-27), Rmin = (L/D)min = 0.667/(1-0.667) = 2.0. Analysis: (b, c, and d) (continued) Exercise 7.32 (continued) Exercise 7.32 (continued) Analysis: (b, c, and d) (continued) Exercise 7.32 (continued) Analysis: (b, c, and d) (continued) Exercise 7.33 Subject: Distillation of a mixture of isopropyl alcohol and water at 1 atm using either a partial reboiler or open steam. Given: Bubble-point feed containing 10 mol% isopropyl alcohol in water. Distillate to contain 67.5 mol% isopropyl alcohol with a 98% recovery. Vapor-liquid equilibrium data , with an azeotrope at 68.54 mol% alcohol. Assumptions: Constant molar overflow. Total condenser. Find: For a reflux ratio, R = L/D = 1.5 times minimum, determine number of stages, if: (a) Partial reboiler is used. (b) Open saturated steam is used. and (c) Minimum number of equilibrium stages. Analysis: In the composition region of operation, the alcohol is the most volatile component. First, compute the distribution of the alcohol. Take a basis of 100 kmol/h of feed. Then, the feed contains 10 kmol/h alcohol and 90 kmol/h of water. For a recovery of 98 mol%, alcohol, distillate contains 9.8 kmol/h of alcohol. For an alcohol purity of 67.5 mol%, the distillate rate = D = 9.8/0.675 = 14.52 kmol/h. Water in the distillate = 14.52 - 9.8 = 4.72 kmol/h. Alcohol in the bottoms = 10 - 9.8 = 0.2 kmol/h. (a) With a partial reboiler, no other water enters the system. Therefore, water in the bottoms = 90 - 4.72 = 85.28 kmol/h. Total bottoms rate = B = 85.28 + 0.2 = 85.48 kmol/h. Mole fraction of alcohol in bottoms = 0.2/85.48 = 0.0023. The minimum reflux is determined from the McCabe-Thiele diagram on the next page, where the equilibrium curve is drawn from the given data and the q-line is vertical, passing through x = 0.10. The rectifying section operating line for minimum reflux usually is a straight line that connects the distillate mole fraction on the 45o line to the intersection of the equilibrium curve and the q-line as shown by the dashed line on the diagram. However, in this case the line mistakenly crosses over the equilibrium curve. Therefore, instead, the operating line is drawn tangent to the equilibrium curve from the point x = xD = 0.675 as shown on the diagram by a solid line. The slope of the operating line = L/V = 0.467. From Eq. (7-27), Rmin = (L/D)min = 0.467/(1-0.467) = 0.876. The operating reflux ratio = R = 1.5Rmin = 1.5(0.876) = 1.314. From Eq. (7-7), L/V = R/(1+R) = 1.314/(1+1.314) = 0.568. On a set of three McCabe-Thiele diagrams on the next page, the rectifying section operating line has this slope and passes through the point, y=0.675, x=0.675. the stripping section operating line passes through the point, y=0.0023, x=0.0023 and intersects the vertical q-line at the point where the rectifying section operating line intersects the q-line. As seen, the equilibrium stages are stepped off starting at the top, with a switch from the rectifying section to the stripping section to minimize the number of stages and, thus, locating the optimal feed stage. To achieve accuracy, one diagram covers the high-concentration region, one the middle region, and one the low-concentration region. The result is between 14 and 15 equilibrium stages. Call it 14 stages plus a partial reboiler. Exercise 7.33 (continued) Analysis: (a) (continued) Exercise 7.33 (continued) Analysis: (a) (continued) Exercise 7.33 (continued) Analysis: (a) (continued) Exercise 7.33 (continued) Analysis: (a) (continued) Analysis: (continued) Exercise 7.33 (continued) (b) When open (live) steam is used with the same reflux ratio, the rectification section operating line and the q-line are identical to part (a) for a partial reboiler. Thus, the part (a) McCabe-Thiele diagram for the high concentration region applies for open steam. However, the stripping section operating line and the bottoms mole fraction change as follows. The liquid rate in the rectification section = L = 1.314D = 1.314(14.52) = 19.08 kmol/h. The vapor rate in the rectifying section = V = L + D = 19.08 + 14.52 = 33.6 kmol/h. The liquid rate below the feed stage = L' = L + F = 19.08 + 100 = 119.08 kmol/h. The vapor rate in the stripping section = V' = V = 33.6 kmol/h = open steam flow rate. The bottoms rate = B = L' = 119.08 kmol/h. The mole fraction of isopropanol in the bottoms = 0.2/119.08 = 0.00168. The change to the part (a) McCabe-Thiele diagrams on the preceding page for the middle concentration region is extremely small because the location of the stripping section operating line changes only slightly. However, the change is important in the low-concentration region. The new McCabeThiele diagram for the low concentration region is shown below. The operating line for the stripping section has a slope of L'/V' = 119.08/33.6 = 3.54 and passes through the point {y = 0, x = 0.00168}. The number of stages remains about the same as for part (a). Thus, without a reboiler, use 15 equilibrium stages in the column. Exercise 7.33 (continued) Analysis: ( c) (continued) (c) The minimum number of stages is determined as shown in the McCabe-Thiele diagrams on the next page by stepping off stages between the equilibrium curve and the 45o line from xB = 0.0023 and xD = 0.675. The number of minimum equilibrium stages = just more than 8 equilibrium stages. Exercise 7.33 (continued) Analysis: ( c) (continued) Exercise 7.33 (continued) Analysis: ( c) (continued) Exercise 7.34 Subject: Stripping of isopropyl alcohol from water at 1 atm using either a partial reboiler or open (live) steam. Given: Bubble-point liquid feed containing 10 mol% alcohol. Vapor overhead to contain 40 mol% alcohol. Boilup, V/F = 0.246. Vapor-liquid equilibrium data from Exercise 7.33. Assumptions: Constant molar overflow. Find: Determine the number of equilibrium stages for: (1) Partial reboiler, (2) Open steam. Analysis: Take a basis of F = 100 kmol/h. Vapor rate leaving top of column = V = 0.246F = D = 0.246(100) = 24.6 kmol/h. Alcohol in overhead vapor = 0.4(24.6) = 9.84 kmol/h. Water in the overhead vapor = 24.6 - 9.84 = 14.76 kmol/h. With a partial reboiler, bottoms rate = B = F - D = 100 - 24.6 = 75.4 kmol/h. Alcohol in bottoms = 10 - 9.84 = 0.16 kmol/h. Mole fraction of alcohol in bottoms = xB = 0.16/75.4 = 0.0021. With isopropanol as the most volatile component, the McCabe-Thiele diagram is given below, where the equilibrium curve is obtained from the data in Exercise 7.33 and the q-line is vertical through x = 0.10. The stripping section operating line passes through the point {x=0.0021, y=0.0021}with a slope = L/V = F/V =100/24.6 = 4.065. It also passes through the point {x=0.1, y=0.4}. From the plot, the number of equilibrium stages = just less than 3. Call it 2 equilibrium stages in the column + partial reboiler. (2) The open steam rate = V = 24.6 kmol/h. The liquid rate = L = 100 kmol/h. Therefore, the slope of the stripping section operating line is the same as for part (1), i.e. L./V = 100/24.6 = 4.065. Now the mole fraction of alcohol in the bottoms = xB = 0.16/100 = 0.0016. Thus, as shown in the McCabe-Thiele diagram below, the operating line passes through the points {x=0.0016, y=0}and {x=0.10, y=0.40}, with the slope of 4.065. Now, The number (1) of equilibrium stages is equal to 3, all of them in the column. Exercise 7.34 (continued) Analysis: Partial Reboiler Case: Exercise 7.34 (continued) Analysis: Open Steam Case: Exercise 7.35 Subject: Distillation of two feeds of mixtures of water and acetic acid at 1 atm. Given: Feed 1 is a bubble-point liquid of 100 kmol/h containing 75 mol% water. Feed 2 is 50 mol% vaporized of 100 kmol/h containing 50 mol% water. Unit consists of a plate column, total condenser, and partial reboiler. Distillate is to contain 98 mol% water. Bottoms is to contain 5 mol% water. Reflux ratio, L/D = R = 1.2 times minimum. Vapor-liquid equilibrium data. Assumptions: Constant molar overflow. Find: Optimal feed stage locations and number of equilibrium stages. Analysis: Water is the more volatile component. Compute flow rates of distillate and bottoms. Overall total material balance: F1 + F2 = 100 + 100 = 200 = D + B (1) Overall water balance: (0.75)(100) + 0.5(100) = 125 = xDD + xBB = 0.98D + 0.05B (2) Solving Eqs. (1) and (2), D = 123.66 kmol/h and B = 76.34 kmol/h. Assume the minimum reflux is controlled by the upper feed. This is verified in the McCabe-Thiele diagram below, where the equilibrium curve is plotted from the data, the q-line for Feed 1 is vertical through the point, x = 0.75, the q-line for Feed 2 has a slope of -1 starting from x = 0.50, and the operating line for the upper section between Feed 1 and the condenser is drawn through the two points, {x=0.98, y=0.98} and the intersection of the equilibrium curve and the q-line for Feed 1. From the plot, for the upper section, L/V = (0.98-0.828)/(0.98-0.75) = 0.661. From Eq. (7-27), R = L/D = (L/V)/[1 - (L/V)] = 0.661/(1 - 0.661) = 1.95. Therefore, L = 1.95D = 1.95(123.66) = 241.1 kmol/h and V = L + D = 241.1 + 123.66 = 364.8 kmol/h. In the middle section, between the two feeds, L' = L + F1 = 241.1 + 100 = 341.1 kmol/h and V' = V = 364.8 kmol/h. Therefore, the slope of the middle section operating line = L'/V' = 341.1/364.8 = 0.935. As seen in the diagram below, this line does not cause a pinched region at Feed 2. Therefore, the assumption is correct and Rmin = 1.95. For an operating reflux ratio of 1.2 times minimum, L = 1.2(241.1) = 289.3 kmol/h. The vapor rate in the upper section = V = L + D = 289.3 + 123.66 = 413 kmol/h. Therefore the upper section operating line has a slope, L/V = 289.3/413 = 0.700 and passes through the point y = x = 0.98. It intersects the vertical q-line at x = 0.75 and, for the slope of 0.700 = (0.98 - y)/(0.98 - 0.75), at y = 0.819. For the middle section, L' = L + F1 = 289.3 +100 = 389.3 kmol/h and V' = V = 413 kmol/h. Therefore, the middle section operating line has a slope of L'/V' = 389.3/413 = 0.943 and intersects the q-line for x = 0.75 at y = 0.819. It intersects the q-line for Feed 2 at y = 0.543. For the lower section, L" = L' + 0.5F2 = 389.3 + 50 = 439.3 kmol/h and V"=V' - 0.5F2= 413 - 50 = 363 kmol/h. Therefore, the lower section operating line has a slope of L"/V" = 439.3/363 = 1.210 and intersects the q-line for Feed 2 at y = 0.543 and the 45o line at xB = 0.05. In the McCabe-Thiele diagrams below and on the next page for the high, middle, and low mole fraction regions, the three operating lines are drawn and the equilibrium stages are stepped off so as to place the two feeds at their optimal locations. As seen, the number of equilibrium stages required = just less than 33 equilibrium stages. Call it 32 equilibrium stages in the column and a partial reboiler. Optimal feed stages are located at Stages 17 and 27 from the top. Analysis: (continued) Exercise 7.35 (continued) Analysis: (continued) Exercise 7.35 (continued) Analysis: (continued) Exercise 7.35 (continued) Analysis: (continued) Exercise 7.35 (continued) Exercise 7.36 Subject: Distillation at 1 atm of a mixture of methanol (M) and ethanol (E) to obtain a distillate, bottoms, and a liquid sidestream. Given: 100 kmol/h of a 25 mol% vaporized mixture of 75 mol% methanol in ethanol. Distillate is 96 mol% methanol and bottoms is 5 mol% methanol. Unit consists of a total condenser, plate column, and partial reboiler. Sidestream is 15 kmol/h of 20 mol% methanol. Reflux ratio, R = 1.2 times minimum. Assumptions: Constant molar overflow. Raoult's law K-values. Find: Number of theoretical stages and optimal locations of feed and sidestream. Analysis: First compute the material balance. Overall total material balance: F = 100 = D + B + S = D + B + 15 (1) (2) Overall methanol material balance: 75 = 0.96D + 0.05B + 0.20(15) Solving Eqs. (1) and (2), D = 74.45 kmol/h and B = 10.55 kmol/h Now determine the equilibrium curve. At 1 atm (14.696 psia), methanol boils at 64.7oC and ethanol boils at 78.4oC. From Fig. 2.4, the vapor pressure of ethanol at 64.7oC = 8.2 psia. The vapor pressure of methanol at 78.4oC = 25 psia. For the Raoult's law K-value, Eq. (2-44) applies. Combining this equation with the definition of the relative volatility of Eq. (2-21), gives, for methanol with respect to the less volatile ethanol, , = (Ps)M/ (Ps)E . At 64.7oC, , = 14.696/8.2 = 1.79. At 78.4oC, , = 25/14.696 = 1.70. Since these values are close (within about 5%), use an equilibrium curve based on a constant = (1.70 + 1.79)/2 = 1.745. From Eq. (7-3), the curve is computed from, y = x/[1 + x(-1)] = 1.745x/(1+0.745x) (3) The McCabe-Thiele diagram on the next page shows the operating lines for determining the minimum reflux ratio. It assumes that the pinch region occurs at the feed stage and not at the sidestream stage. From Eq. (7-18), q for 25 mol% vaporized = 0.75. From Eq. (7-26), the slope of the q-line = q/(q-1) = 0.75/(0.75-1) = -3. The upper section operating line intersects the q-line and equilibrium curve at y=0.823 and x=0.727. Thus, slope of the upper section operating line = L/V = (0.96-0.823)/(0.96-0.727) = 0.588. From Eq. (7-27), Rmin = (L/V)/[1 - (L/V)] = 0.588/(10.588) = 1.427. Therefore, L = 1.427(74.45) = 106.3 kmol/h and V = L + D = 106.3 + 74.45 = 180.75 kmol/h. In the middle section between the feed stage and sidestream stage, for 25 mol% vaporization, L' = L + 0.75(100) = 106.3 + 75 = 181.3 kmol/h and V' = V - 0.25(100) = 180.75 25 = 155.75 kmol/h. Thus, L'/V' = 181.3/155.75 = 1.164. The middle section operating line has this slope and passes through y=0.823 and x=0.727. In the lower section operating line below the sidestream, for a liquid sidestream flow rate of 15 kmol/h, L" = L' - S = 181.3 - 15 = 166.3 kmol/h and V" = V' = 155.75 kmol/h. Thus, L"/V" = 166.3/155.75 = 1.068. This line passes through the point y=0.05 and x = 0.05 with the slope of 1.068. It also intersects the middle section operating line at the sidestream composition, xs = 0.20. These three operating lines are drawn in the McCabe-Thiele diagram, showing that the middle section operating line lies below Analysis: (continued) Exercise 7.36 (continued) the equilibrium curve. Therefore, the assumption that the minimum reflux is controlled by the feed stage is verified and the minimum reflux ratio is 1.427. For actual operation, reflux ratio = 1.2Rmin = 1.2(1.427) = 1.712. Reflux rate in upper section = L = 1.2Lmin = 1.2(106.3) = 127.6 kmol/h. Vapor rate = V = L + D = 127.6 + 74.45 = 202.05 kmol/h. Slope of upper section operating line = L/V = 127.6/202.05 = 0.632. For constant molar overflow, the flow rates in the other sections are: Middle section: L' = 202.6 kmol/h V' = 177.05 kmol/h L'/V' = 1.144 Lower section: L" = 187.6 kmol/h V" = 177.05 kmol/h L"/V" = 1.060 Analysis: (continued) Exercise 7.36 (continued) In the McCabe-Thiele diagrams below, upper, middle, and lower operating lines are based on these values starting from the upper line, which passes through the point x = 0.96 on the 45o line. Equilibrium stages are stepped off starting from the distillate composition and switching operating lines at the appropriate times to determine the optimal feed stage and sidestream stage locations. A separate McCabe-Thiele diagram is used for the upper section to achieve better accuracy. As seen 19 equilibrium stages plus a partial reboiler are required. The optimal feed stage is 9 from the top and the optimal sidestream stage is 17 from the top. Analysis: (continued) Exercise 7.36 (continued) Exercise 7.37 Subject: Distillation at 1 atm of a mixture of toluene and phenol for a given boilup ratio, with an alternative using an interreboiler. Given: 1,000 kmol/h of a saturated liquid feed of 25 mol% toluene. Distillate is 98 mol% toluene and bottoms is 2 mol% toluene. The base unit consists of a total condenser, a plate column, and a partial reboiler, with a boilup ratio, VB = V / B = 1.15 times minimum. The alternative unit adds an interreboiler midway in the stripping section to provide 50% of the boilup. A table of T-y-x phase equilibrium data Assumptions: Constant molar overflow. Find: For each unit, determine the number of theoretical stages. For the alternative unit, determine the temperature of the interreboiler stage. Analysis: Toluene is the more volatile unit. Base unit: From the McCabe-Thiele diagram below, the minimum boilup ratio is determined from the slope of the stripping section operating line that intersects the equilibrium curve at the feed composition of xF = 0.25. ( L /V ) From Eq. (7-28), min = 0.815 - 0.02 = 3.46 0.25 - 0.02 (VB )min = ( L /V ) 1 min -1 = 1 = 0.407 3.46 - 1 Analysis: Base Case (continued) Exercise 7.37 (continued) For column operation, VB = 1.15(VB)min = 1.15(0.407) = 0.468 From Eq. (7-12), the slope of the stripping section operating line is, V + 1 0.468 + 1 L /V = B = = 314 . VB 0.468 As shown in the McCabe-Thiele diagram, below, a line of this slope is drawn through the point x = y = 0.02 until it intersects the q-line. Equilibrium stages are stepped off, starting from the distillate point at y = x = 0.98. The optimal feed stage location is located as shown at stage 5 from the top. The total number of stages required is just less than 8, with one of those stages being the partial reboiler. Exercise 7.37 (continued) Analysis: Base Case (continued) Exercise 7.37 (continued) Analysis: (continued) Alternative unit with Interreboiler: For column operation with an interreboiler that provides 50% of the reboiler duty, the operating boilup ratio between the reboiler and the interreboiler is 50% of 0.468 or 0.234. From Eq. (7-12), the slope of the operating line in this region is: L /V = VB + 1 0.234 + 1 = = 5.27 VB 0.234 Between the interreboiler and the feed stage, the slope of the operating line remains at 3.14, based on a material balance around the column section from the bottoms to the region between the interreboiler and the feed stage. These two operating lines are shown in the McCabe-Thiele diagram below, where both pass through the point y = x = 0.02. Stepping off stages from the bottom, it is seen that 3 stages are needed below the feed stage. The first stage is the partial reboiler. The interreboiler is located at the second equilibrium stage. A total of 8 equilibrium stages is required, just slightly more than that when all of the heat input is to the partial reboiler at the bottom of the column. The vapor composition of the interreboiler stage is 0.345, which from the given T-y-x phase equilibrium data corresponds to a temperature of approximately 173oC. The interreboiler could also be located at the third stage from the bottom. The would increase the number of stages by about half of a stage and lower the temperature of the interreboiler stage to 162oC. Exercise 7.37 (continued) Analysis: (continued) Alternative unit with Interreboiler: Exercise 7.38 Subject: Effect of the addition of an intercondenser and interreboiler to a distillation column separating n-butane and n-pentane. Given: Distillate and bottoms compositions of actual operation (before addition of intercondenser and interreboiler) compared to design specification. Assumptions: Constant molar overflow. Constant relative volatility. Column is large enough in diameter to handle increased reflux and boilup. Raoult's law (ideal solutions and ideal gas law). Find: Whether the addition can improve the operation because of the increased reflux and boilup produced by the intercondenser and interreboiler. Analysis: First, estimate the average relative volatility for nC4/nC5. Assume a distillate temperature of 120oF so that cooling water can be used in the condenser. This corresponds to a saturation pressure of about 70 psia. Using Fig. 2.8 with Eqs. (2-21) and (2-44), the relative volatility of butane with respect to pentane is = 1.1/0.38 = 2.9. Assuming a 5 psi pressure drop, gives a bottoms pressure of 75 psia and a corresponding temperature of 200oF. Using Fig. 2.8 with Eqs. (2-21) and (2-44), = 2.1/0.9 = 2.3. Take the average relative volatility as 2.6 and draw a y-x equilibrium curve using Eq. (7-3), y= 2.6 x x = 1 + x ( - 1) 1 + 16 x . (1) The equilibrium curve, based on Eq. (1) is shown below in a McCabe-Thiele diagram. Included on the diagram are arbitrary operating lines and a q-line for an equimolar feed that is 50 mol% vaporized. Using these lines, 15 equilibrium stages are stepped off between the compositions of the actual operation, xD = 1 - 0.1349 = 0.8651 and xB = 0.0428. The slope, L/V, of the rectifying section operating line is 0.52. Analysis: Actual operation before addition (continued) Exercise 7.38 (continued) Exercise 7.38 (continued) Analysis: Addition (continued) When an interreboiler is added between the reboiler and the feed stage and an intercondenser is added between the feed stage and the overhead condenser, the column is made up of 4 sections instead of 2. Each section has its own operating line as shown in the McCabeThiele diagram below. In order to maintain the same reflux ratio and boilup ratio, the intercondenser is designed to condense a molar flow rate equal to that produced by the interreboiler. Thus, in Section 2 between the intercondenser and the feed stage, the L/V ratio is higher than the value in Section 1 between the overhead condenser and the intercondenser. Thus, the two operating lines above the feed stage have different slopes, but by material balance both lines pass through the distillate composition on the 45o line. Correspondingly, in Section 3 between the interreboiler and the feed stage, the L/V ratio is lower than the value in Section 4 between the reboiler and the interreboiler. Thus, the two operating lines below the feed stage have different slopes, but by material balance both lines pass through the bottoms composition on the 45o line. Also note that by material balance, the operating lines for Sections 1 and 4 intersect on the 45o line, and the operating lines for Sections 2 and 3 intersect on the 45o line. Starting from xD = 0.9974, stages are stepped off from the top between the operating line for Section 1 and the equilibrium curve for a few stages before switch to the operating line for Section 2. Then 5 stages are stepped off until the feed stage is reached. Then, 4 stages are stepped off between the operating line for Section 3 and the equilibrium curve before switching to the operating line for Section 4. The switches are made to maintain the same number of total stages, 15, and the same location for the feed stage. Other combinations of arbitrary operating lines and a q-line can be used to illustrate the potential of an intercondenser and interreboiler. The important thing to note is that the addition of an intercondenser and interreboiler increases the distance between the equilibrium curve and the operating lines for Sections 2 and 3 so that the steps in Sections 2 and 3 accomplish larger composition changes. Exercise 7.38 (continued) Analysis: Addition (continued) Exercise 7.39 Subject: Distillation of a mixture of para-dichlorobenzene (P) and ortho-dichlorobenzene (O), two close-boiling isomers, using the McCabe-Thiele diagram, with the Kremser equation to accurately calculate the separations at the two ends of the column. Given: Feed of 62 mol% P and 38 mol% O that is slightly vaporized with q = 0.9. Distillate is liquid of 98 mol% P. Bottoms is 96 mol% O. Pressures at top and bottom are 20 psia and 15 psia, respectively. Reflux ratio, R = 1.15 times minimum. Assumptions: Constant molar overflow. Total condenser and partial reboiler. Average relative volatility based on Raoult's law. Find: Number of theoretical stages from McCabe-Thiele diagram, using Kremser supplement for the two ends of the column. Analysis: From a simulation program or handbook, the temperatures at the top and bottom of the column, based on the given pressures and product compositions, are determined to be approximately 350 and 380oF. From vapor pressure data, e.g. from CHEMCAD, using Eq. (7-1), P,O at 350o F = PPs 15.65 P s 2316 . = = 1167 and P,O at 380o F = Ps = . = 1159 . s PO 13.41 PO 19.98 Take the average relative volatility as (1.167 + 1.159)/2 = 1.163. From Eq. (7-26), the slope of the q-line = q/(q - 1) = 0.9/(0.9 - 1) = -9. Feed is 10 mol% vaporized. Apply the Fenske-Underwood-Gilliland method to obtain an initial estimate of reflux and stage requirements. Can use the SHOR model in CHEMCAD. The results are: Nmin = 48, Rmin = 9.37 R = 1.15Rmin = 10.77 N = 103 (102 + reboiler) N of feed = stage 50 from the top To construct the equilibrium curve for the McCabe-Thiele method, use Eq. (7-3), yP = P-O xP 1163xP . = 1+ xP P-O - 1 1 + 0163xP . (1) From Eq. (7-9), the slope of the rectifying section operating line is, L/V = R/(R+1) = 10.77/(10.77 + 1) = 0.9150 Analysis: (continued) This line passes through x = 0.98 on the 45o line. The equation for this line is given by Eq. (7-9), L 1 1 (2) x+ x D = 0.915x + (0.98) = 0.915x + 0.08326 V R +1 11.77 + 1 Below the feed stage, with 10 mol% vaporization of the feed, L / V = 10535. From a . rearrangement of Eq. (7-12), the boilup ratio, VB, is 18.6916. The equation of the stripping section operating line is given by Eq. 7-14), L 1 1 y= x- x D = 10535x - . (0.04) = 10535x - 0.00214 . (3) V VB 18.6916 y= The equation for the q-line is given by Eq. (7-26), y= q 1 0.9 1 x- zF = x- (0.62) = -9 x + 6.2 q -1 q -1 0.9 - 1 0.9 - 1 (4) Exercise 7.39 (continued) Based on Eqs. (1) to (4), the McCabe-Thiele diagram in terms of P, the more volatile component, is drawn below for three regions: (2) x = 0.2 to 0.4, (3) x = 0.4 to 0.6, and (4) x = 0.6 to 0.8, in order to gain accuracy. In these three regions, 28 stages are stepped off in the rectifying section up to x = 0.8, and 34.3 stages are stepped off in the stripping section down to x = 0.2. Let region (1) extend from x = 0.8 to 0.98 (i.e. xD). Apply the Kremser equation, Eq. (7-39) to this region. Apply this equation to the heavy component, O. Obtain the K-value for O from the top of 1.167, taking the K-value for P = 1.00. Therefore, KO = 1/1.167 = 0.857. Therefore, the absorption factor, A = L/KV = 0.915/0.857 = 1.067. Other quantities needed in Eq. (7-39) are: xo = 1 - (xD)P = 1 - 0.98 = 0.02 y1 = xo = 0.02 From Eq. (2), for xN = 0.8, yN+1 for P = 0.8153. Therefore, for O, yN+1 = 1 - 0.8153 = 0.1847. log NR = 1 1 + 1- A A logA yN +1 - xo K y1 - xo K log = 1 1 + 1- 1.067 1067 . 0.1847 - 0.02(0.857) 0.02 - 0.02(0.857) log1.067 = 23.6 Let region (5) extend from x = 0.04 (i.e. xB) to 0.20. Apply the Kremser equation, Eq. (7-40) to this region. Apply this equation to the light component, P. Obtain the K-value for P from the bottom of 1.159, taking the K-value for O = 1.00. Therefore, KP = 1.159. Therefore, the absorption factor in the stripping section is A = L / KV = 10535 / 1159 = 0.9085 . Other . . values needed in Eq. (7-40) are x1 = xB = 0.04 and xN+1 = 0.20. Therefore, x -x /K 0.20 - 0.04 / 1159 . log A + 1 - A N +1 1 log 0.9085 + 1 - 0.9085 x1 - x1 / K 0.04 - 0.04 / 1159 . NS = = = 135 . 1 1 log log A 0.0985 Analysis: (continued) Exercise 7.39 (continued) Summing the above results, Number of equilibrium stages in the rectifying section = 28 + 23.6 = 51.6 stages Number of equilibrium stages in the stripping section = 34.3 + 13.5 = 47.8 stages Call it 99 stages in the column plus a partial reboiler with the feed to stage 52 from the top. This compares to 102 stages in the column plus a partial reboiler with the feed to stages 50 from top as determined by the Fenske-Underwood-Gilliland method. Analysis: (continued) Exercise 7.39 (continued) Exercise 7.39 (continued) Analysis: (continued) Exercise 7.40 Subject: Use of a McCabe-Thiele diagram to determine stage requirements for the distillation of air into nitrogen and oxygen using a Linde double column. Given: As shown in Fig. 7.45, the distillation consists of an upper column (UC) operating at 1 atm, on top of a lower column (LC) operating at 4 to 5 atm. Compressed air, containing 79 mol% N2 is condensed to supply heat in the reboiler of LC, and then is fed as liquid air to an intermediate tray of LC. Bottoms liquid from LC, containing about 55 mol% N2, is the feed to an intermediate tray in UC. The reboiler of UC is the condenser for LC. Condensate from the top of LC is nearly pure N2, which is sent as reflux to the top of UC. The reboiler at the bottom of UC provides almost pure O2 boilup for UC. Nearly pure liquid O2 is withdrawn from the UC reboiler sump at the bottom of UC. The UC has no condenser. Nearly pure gaseous N2 leaves the top of UC. This is consistent with the fact that N2 with a normal boiling point of -195.8oC (77.4 K) is more volatile than O2 with a normal boiling point of -183oC (90.2 K). Assumptions: Constant molar overflow. Constant relative volatility at each pressure. Find: Construction lines on a McCabe-Thiele diagram that enable the determination of stage requirements. Analysis: In LC, the N2 composition ranges from 55 mol% at the bottom to about 99 mol% at the top, with a feed of 79 mol%. Based on calculations using K-values from the SRK equation of state at 4.5 atm, the average relative volatility in LC is 2.5. In UC, the N2 composition ranges from about 1 mol% at the bottom to 99 mol% at the top, with a feed of 55 mol%. Based on calculations using K-values from the SRK equation of state at 1 atm, the average relative volatility in UC is 4.0. Using Eq. (7-3), equilibrium curves for these constant cases are shown in the McCabe-Thiele diagram on the next page. However, so as not to clutter the diagram, the curve for UC at 1 atm is based on N2, using, yN 2 = N 2 - O2 xN 2 1+ xN 2 N 2 - O 2 - 1 = 4 xN 2 1+ 3xN 2 (1) while the curve for LC at 4.5 atm is based on O2, using, yO 2 = O 2 - N 2 xO 2 1+ xO 2 O 2 - N 2 - 1 = (1 / 2.5) xO 2 1 + xO 2 1 / 2.5 - 1 = 0.4 xO 2 1- 0.6 xO 2 (2) Note that the equilibrium curve for 1 atm is above the 45o line, while that for 4.5 atm is below the 45o line. Typical operating lines and q-lines are shown for determining the stage requirements. Analysis: (continued) Exercise 7.40 (continued) Exercise 7.41 Subject: Comparison of measured with predicted plate efficiency for distillation of a methanol-water mixture. Given: Performance data for a distillation column. Vapor-liquid equilibrium data. Assumptions: Partial reboiler is an equilibrium stage. Constant molar overflow. Find: (a) (b) (c) (d) Overall plate efficiency from performance data. Predicted plate efficiency from Drickamer-Bradford correlation. Predicted plate efficiency from O'Connell correlation. Predicted efficiency from Chan-Fair correlation. Analysis: (a) Convert the performance data for feed and product flow rates and compositions from mass units into mole units, molecular weights of 32.04 for methanol and 18.02 for water. The results are as follows: Component Methanol Water Total: Flow rate, lbmol/h: Feed Distillate Bottoms 709.1 702.7 6.4 1260.8 65.3 1195.5 1969.9 768.0 1201.9 Mole fraction: Feed Distillate 0.360 0.915 0.640 0.085 1.000 1.000 Bottoms 0.0053 0.9947 1.0000 Use the McCabe-Thiele method, based on methanol mole fractions, to find the number of equilibrium stages needed. The slope of the rectifying section operating line is given by Eq. (77). L/V = R/(R + 1)=0.947/(1.947)=0.486. This line intersects the 45o line at x = 0.915. The slope of the stripping section operating line is given by Eq. (7-12). L / V = VB + 1 / VB = (1.138 + 1)/1.138 = 1.88. This line intersects the 45oi line at x = 0.0056. The McCabe-Thiele diagram is given below in two parts to achieve accuracy. From the two parts of the diagram, it is observed that the two operating lines do not intersect on the q-line. This is probably because the assumption of constant molar overflow is not valid and the operating lines have some curvature. Assuming the feed stage in the actual column is near the optimal location, the questionable McCabe-Thiele diagram gives 5 equilibrium stages in the stripping section, one of which is the partial reboiler, and 4.2 equilibrium stages in the rectifying section. Thus, the total equilibrium stages in the column = Nt = 8.2 + the partial reboiler. The column contains 12 plates + the partial reboiler. From Eq. (6-21), the overall plate efficiency is, Eo = Nt /Na = 8.2/12 = 0.68 or 68%. The column contains 5 plates in the rectifying section. This is equivalent to a plate efficiency in that section of 4.2/5 = 0.84 or 84%. The stripping section contains 6 plates plus the feed plate. Analysis: (a) (continued) Exercise 7.41 (continued) Therefore, the efficiency in this section is 4/7 = 0.57 or 57%. Note these results are subject to degree of curvature of the operating lines and the placement of the feed in the McCabe-Thiele method. In the McCabe-Thiele method, if one more equilibrium stage is added to the stripping section, so as to give 5 stages plus the partial reboiler, then the rectifying section only needs 3.8 equilibrium stages. Then the overall plate efficiency is 73%, with 71% in the stripping section and 76% in the rectifying section. This still does not account for the effect of curvature of the operating lines. Analysis: (a) (continued) Exercise 7.41 (continued) (b) From Eq. (7-42), Eo = 13.3 - 66.7 log Take the viscosity as that of the feed = 0.34 cP Eo = 13.3 - 66.7 log (0.34) = 44.6% This is poor agreement with the performance data. (c) From Eq. (7-43), Eo = 50.3()-0.226 At the feed composition, x = 0.36 and y = 0.71. Therefore, from Eqs (2-19) and (2-21 combined, the relative volatility is, = ( y/x)/[(1 - y)/(1 - x)] = (0.71/0.36)/(0.29/0.64) = 4.4 Exercise 7.41 (continued) Analysis: Eo = 50.3[(4.4)(0.34)]-0.226 = 45.9% Now correct for length of liquid path from Fig. 7.5. Column diameter = 6 ft. Assume length of liquid path = 70% of column diameter = 0.7(6) = 4.2 ft. From Fig. 7.5, correction to be added = 10%. Therefore corrected Eo = 45.9 + 10 = 55.9% This also appears to be low. (d) From Eq. (6-56), NOG = - ln (1 - EOV). Therefore, EOV = 1 - exp(-NOG) Use Eqs. (6-62, (6-64), (6-66), and (6-67) as in Example 6.7. Carry out the calculations at the bottom tray based on methanol diffusion. Conditions are: Molar flow rate, kmol/h Molecular weight Density, kg/m3 Gas 630 18.5 0.657 Liquid 1,192 18.0 940 Estimate liquid diffusivity of methanol in water at 212oF = 373 K. From Perry's Handbook, DL = 1.6 x 10-5 cm2/s at 25oC. The viscosity of water at 212oF = 0.25 cP. Using Eq. (3-39) to correct for temperature and viscosity, DL = 1.6 x 10-5 (373/298)(1/0.25) = 8 x 10-5 cm2/s Estimate gas diffusivity of methanol in water vapor from Eq. (3-36). T = 373 K MAB = 2/[(1/32) + (1/18)] = 23. Using Table 3.1, = 15.9 + 2.31(4) + 611 = 31.3, . = 131 . V V methanol water 0.00143(373)1.75 DV = = 0.30 cm2/s 1/ 2 1/ 3 1/ 3 2 (151 / 14.7)23 [31.3 + 131 ] . . Required tray dimensions: DT = 6 ft, A = 3.14(6)2/4 = 28.3 ft2 = 2.63 m2 Aa = 0.91 A = 0.91(2.63) = 2.39 m2 = 23,900 cm2, Lw = 42.5 in. = 1.08 m Tray conditions: = 1 - 0.617 = 0.383, qL = 47,300/(60)(8.33) = 94.6 gpm = 5,970 cm3/s From Eq. (6-54), Cl = 0.362 + 0.317 exp[-3.5(2)] = 0.362 From Eq. (6-51), f = 0.40, 94.6 hl = 0.383 2.0 + 0.362 (42.5)(0.383) 2/3 = 1.21 in. = 3.08 cm2 From Eq. (6-64), t L = (3.08)(23,900)/5,970 = 12.3 s Analysis: (d) (continued) The continuity equation is, mV = U a Aa V , so U a , ft/s = Exercise 7.41 (continued) mV 630(18.5)(2.205) / 3600 = = 6.8 ft/s = 2.07 m/s Aa V 28.3(0.91)(0.657 /16.02) From Eq. (6-65), tG = (0.617)(3.08) = 0.024 s (0.383)(6.8)(2.54)(12) From below Eq. (6-67), 0.5 F=Ua V = 2.07(0.657)0.5 = 1.68 (kg/m)0.5/s From Eq. (6-67), k L a = 78.8(8 10 -5 ) 0.5 (168 + 0.425) = 1.48 s-1 . From Eq. (6-66), k G a = 1,030(0.30)1/ 2 0.40 - 0.842(0.40) 2 3.081/ 2 = 85.3 s-1 From Eq. (6-63), N L = k L at L = 148(12.3) = 18.2 . From Eq. (6-62), N G = k G at G = 85.3(0.024) = 2.05 From the vapor-liquid equilibrium data at the bottom of the column, Kmethanol = 0.156/0.0246 = 6.34 Absorption factor = KV/L = (6.34)(630)/1,192 = 3.35 From Eq. (6-61), 1 1 ( KV / L) 1 3.35 = + = + = 0.488 + 0.184 = 0.672 N OG N G NL 2.05 18.2 NOG = 1/0.672 = 1.488 Thus, the gas-phase resistance is more important than the liquid-phase resistance. From Eq. (6-56), EOV = 1 - exp(- N OG ) = 1 - exp(-1.488) = 0.774 or 77.4% This is in very good agreement with the performance data. Exercise 7.42 Subject: Estimation of efficiencies, EMV and Eo from EOV for methanol-water mixture, as measured with a small Oldershaw column. Given: Column conditions from Exercise 7.41. Find: EMV and Eo Analysis: Case 1: Assume complete mixing on the trays. EMV = EOV = 0.65 or 65% Case 2: Assume plug flow of liquid with no longitudinal diffusion. Take conditions at the top of the column. From Eq. (7-7), L/V = R/(R + 1) = 0.947/(1 + 0.947) = 0.486 From Eq. (6-33), = m/(L/V) From the vapor-liquid equilibrium data given in Exercise 7.41, m = dy/dx = (1 - 0.915)/(1 - 0.793) = 0.41 = 0.41/0.486 = 0.844 From Eq. (6-32), EMV = 1 1 [ exp(EOV ) - 1] = {exp [0.844(0.65)] - 1} = 0.754 or 75.4% 0.844 The actual value of EMV probably lies inbetween 65% and 75.4%, or say 70%. From Eq. (6-37), assuming that the equilibrium and operating lines are straight, Eo = log [1 + EMV ( - 1) ] log [1 + 0.70(0.844 - 1) ] = = 0.68 or 68% log log(0.844) Exercise 7.43 Subject: Estimation of column diameter at top tray of a distillation column separating benzene from monochlorobenzene. Given: Conditions at the top tray in Fig. 7.46. Find: Column diameter for a valve tray at 85% of flooding. Analysis: Use Fig. 6.24, where the value of FLV is needed. V = 336.5 lbmol/h, L = 274.7 lbmol/h Because both the vapor and liquid are almost pure benzene, take MV and ML = 78.1 From Perry's Handbook, the liquid density of benzene at 204oF and 23 psia = L = 52 lb/ft3 From the ideal gas law, V = PMV/RT = (23)(78.1)/[(10.73)(664)] = 0.252 lb/ft3 FLV LM L V = VMV L 1/ 2 = (274.7)(78.1) 0.252 (336.5)(781) 52 . 0.5 = 0.057 Assume 24-inch tray spacing. From Fig. 6.24, CF = 0.37 From below Eq. (6-44), Ad/A = 0.1 From the Handbook of Chemistry and Physics, for benzene under its own vapor, = 30 dyne/cm. From below Eq. (6-42), FST = (30/20)0.2 = 1.084 Assume that the foaming factor, FF = 1.0 and take FHA = 1.0 . From Eq. (6-42), C = 1.084(1.0)(1.0)(0.37) = 0.40 ft/s From Eq. (6-40), - V Uf =C L V 1/ 2 52 - 0.252 = 0.40 0.252 0.5 = 5.73 ft / s From Eq. (6-44), 4VM V DT = fU f (1 - Ad / A)V 1/ 2 4(336.5 / 3600)(78.1) = 0.85(5.73)(3.14)(1 - 0.1)(0.252) 0.5 = 2.9 ft Exercise 7.44 Subject: Diameters and heights for separation of propylene from propane in two columns. Given: Feed and product conditions and numbers of trays for sieve-tray columns in Fig. 7.47. Assumptions: Constant molar overflow. 85% of flooding. Find: Column diameters, tray efficiencies, numbers of actual trays, and column heights. Analysis: Column diameters: Assume that the diameter of the second column is controlled at the top tray, and than the diameter of the first column is controlled by the bottom tray. Top tray of second column: Use the properties of pure propylene. Distillate rate = 3.5 + 360 - 12.51 = 350.99 lbmol/h With an R = 15.9, L = 15.9(350.99) = 5,580 lbmol/h. V = L + D = 351 + 5,580 = 5,931 lbmol/h. From Perry's Handbook, at 116oF = 576oR = 320 K, vapor density = V = 2.64 lb/ft3 and liquid density = L = 29.0 lb/ft3. FLV LM L V = VMV L 1/ 2 (5,580)(42.1) 2.64 = (5,931)(42.1) 29.0 0.5 = 0.284 Assume 24-inch tray spacing. From Fig. 6.24, CF = 0.25 From below Eq. (6-44), Ad/A = 0.1+ (0.284 - 0.1)/9 = 0.120 Assume surface tension, , = 10 dyne/cm. From below Eq. (6-42), FST = (10/20)0.2 = 0.87 Assume that the foaming factor, FF = 1.0 and take FHA = 1.0. From Eq. (6-42), C = 0.87(1.0)(1.0)(0.25) = 0.22 ft/s - V From Eq. (6-40), U f = C L V From Eq. (6-44), 4VM V DT = fU f (1 - Ad / A)V 1/ 2 1/ 2 = 0.22 29.0 - 2.64 2.64 0.5 = 0.70 ft / s 4(5931/ 3600)(42.1) = 0.85(0.70)(3.14)(1 - 0.12)(2.64) 0.5 = 8.0 ft Bottom tray of first column: Use the properties of pure propane. L in first column = L in second column + feed = 5,580 + 360 + 240 = 6,180 lbmol/h. V in first column = L in first column - B = 6,180 - (600 - 351) = 5,931 lbmol/h From Perry's Handbook, at 136oF = 596oR = 331 K, vapor density = V = 2.93 lb/ft3 and liquid density = L = 27.0 lb/ft3. Analysis: (continued) FLV Exercise 7.44 (continued) LM L V = VMV L 1/ 2 = (6,180)(44.1) 2.93 (5,931)(44.1) 27.0 0.5 = 0.343 Assume 24-inch tray spacing. From Fig. 6.24, CF = 0.22 From below Eq. (6-44), Ad/A = 0.1+ (0.343 - 0.1)/9 = 0.127 Assume surface tension, , = 10 dyne/cm. From below Eq. (6-42), FST = (10/20)0.2 = 0.87 Assume that the foaming factor, FF = 1.0 and take FHA = 1.0. From Eq. (6-42), C = 0.87(1.0)(1.0)(0.22) = 0.19 ft/s - V From Eq. (6-40), U f = C L V From Eq. (6-44), 4VM V DT = fU f (1 - Ad / A)V 1/ 2 1/ 2 27.0 - 2.93 = 019 . 2.93 0.5 = 0.55 ft / s 4(5, 931/ 3600)(44.1) = 0.85(0.55)(3.14)(1 - 0.127)(2.93) 0.5 = 8.8 ft Tray Efficiencies: Use Fig. 7.32. Need relative volatility and average liquid viscosity. From Fig. 2.8, K-values and relative volatilities are: Component K-value at 116oF, 280 psia K-value at 136oF, 300 psia Propylene 1.00 1.14 Propane 0.87 1.00 At the top, = 1.00/0.87 = 1.15. At the bottom, = 1.14/1.00 = 1.14 At the top, propylene viscosity = 0.45 cP. At the bottom, propane viscosity = 0.44 cP. At the top, = 1.15(0.45) = 0.52. From Fig. 7.32, Eo = 60%. At the bottom, = 1.14(0.44) = 0.50. From Fig. 7.32, Eo = 62%. Actual trays and column heights: Actual tray requirements are 90/0.62 = 145 trays for the bottom column and 90/0.60 = 150 trays for top column. Column heights are: Bottom column = 144(2) + 4 + 10 = 302 ft Top column = 149(2) + 4 + 10 = 312 ft Exercise 7.45 Subject: Sizing of a vertical flash drum. Given: Flash temperature and pressure, with product flow rates in Fig. 7.48. Find: Drum height and diameter. Analysis: To determine minimum drum diameter, use Eq. (6-44) with f = 0.85 and Ad = 0. The flooding velocity is obtained from Eqs. (6-40) and (6-42), using FST = 1.0, FHA = 1.0 and FF = 1.0, with CF from the 24-inch plate spacing curve in Fig. 6.24. The value of FLV is obtained first, using the conditions in Fig. 7.48. MV = [187.6(58.1) + 176.4(72.2) + 82.5(86.2)]/446.5 = 68.9 ML = [112.4(58.1) + 223.6(72.2) + 217.5(86.2)]/553.5 = 74.8 From a simulation program, V = 1.13 lb/ft3, compared to 0.965 from the ideal gas law and L = 33.2 lb/ft3. Also given, V = 446.5 lbmol/h and L = 553.5 lbmol/h In Fig. 6.24, FLV LM L V = VMV L 1/ 2 = 553.6(74.8) 113 . 446.5(68.9) 33.2 1/ 2 = 0.248 From Fig. 6.24, CF = 0.27 ft/s. From Eq. (6.42), C = (1)(1)(1)0.27 = 0.27 ft/s. From Eq. (6-40), - V Uf =C L V From Eq. (6-44), 4VMV DT = fU f V 1/ 2 1/ 2 33.2 - 113 . = 0.27 113 . 1/ 2 = 144 ft/s . 4(446.5 / 3600)(68.9) = 0.85(1.44)(314)(113) . . 1/ 2 = 2.81 ft Assume a liquid residence time, t, of 5 minutes (0.0833 h), half full . 2 LM L t 2(553.5)(74.8)(0.0833) = = 208 ft3 From Eq. (7-44), vessel volume = VV = L 33.2 4V 4(208) Now compute the height, H, from Eq. (7-45). H = V2 = = 33.6 ft DT 314(2.81) 2 . Thus, H/D = 33.6/2.81 = 12.0. This is too large, so re-dimension the volume to give H/D = 4. From Eq. (7-46), V DT = V 1/ 3 208 = 3.14 1/ 3 = 4.05 ft From Eq. (7-46), H = 4DT = 4(4.05) = 16.2 ft Exercise 7.46 Subject: Sizing of a horizontal reflux drum. Given: Flow conditions leaving reflux drum in Fig. 7.49. Find: Drum length and diameter. Analysis: Liquid rate leaving drum = (3 + 1)120 = 480 lbmol/h Assume properties of pure n-hexane. From Fig. 2.4, boiling point of nC6 at 1 atm = 150oF. From Fig. 2.3, L = 0.63 g/cm3 = 39 lb/ft3 Assume a liquid residence time in the reflux drum of 5 minutes (0.0833 h), half full. From Eq. (7-44), 2 LM L t 2(480)(86.2)(0.0833) VV = = = 177 ft3 L 39 From Eq. (7-46), V DT = V 1/ 3 177 = 3.14 1/ 3 = 3.83 ft From Eq. (7-46), H = 4DT = 4(3.83) = 15.3 ft Might use a drum of 4 ft diameter and 16 ft long. Exercise 7.47 Subject: Sizing of a sieve-tray distillation column separating methanol and water. Given: Column top and bottom conditions in Fig. 7.50. Assumptions: Column feed is a bubble-point liquid. Constant molar overflow. Find: (a) Column diameters at top and bottom. Should column be swaged? (b) Length and diameter of horizontal reflux drum. Analysis: First, compute the liquid and vapor flow rates at the top and bottom of the column. Molecular weight of distillate = 0.9905(32.04) + 0.0095(18.02) = 31.91 Distillate flow rate = D = 462,385/31.91 = 14,490 lbmol/h Molecular weight of bottoms = 0.0101(32.04) + 0.9899(18.02) = 18.16 Bottoms flow rate = B = 188,975/18.16 = 10,410 lbmol/h From these values and the given compositions, the overall material balance is: Lbmol/h: Component Feed Distillate Bottoms Methanol 14,457 14,352 105 Water 10,443 138 10,305 Total: 24,900 14,490 10,410 From the reboiler duty, assuming the bottoms is pure water, can compute the boilup rate. Heat of vaporization of water at 262.5oF = 937 Btu/lb = 16,885 Btu/lbmol Boilup rate = V = 442,900,000/16,885 = 26,230 lbmol/h Assume constant molar overflow with saturated liquid feed. Therefore, overhead vapor rate also is equal to 26,230 lbmol/h. The reflux rate = V - D = 26,230 - 14,490 = 11,740 lbmol/h A more accurate calculation by CHEMCAD with an overall energy balance gives a reflux rate of 15,800 lbmol/h and an overhead vapor rate of 30,300 lbmol/h. But use the constant molar overflow values. (a) Column diameter at the bottom: From a simulation program, at 262.5oF and 40 psia, V = 0.0994 lb/ft3 and L =57.9 lb/ft3. V = 26,230 lbmol/h and L = 10,410 + 26,230 = 36,640 lbmol/h In Fig. 6.24, FLV LM L V = VMV L 1/ 2 = 36,640(18.2) 0.0994 26,230(18.29) 57.9 1/ 2 = 0.058 From Fig. 6.24, for 24-inch plate spacing, CF = 0.27 ft/s. For Eqs. (6-40) and (6-42), use FHA = 1.0 and FF = 1.0. For water, = 51 dyne/cm Therefore, FST = (51/20)0.2 = 1.21. From Eq. (6.42), C = 1.21(1)(1)0.37 = 0.45 ft/s. Analysis: (a) (continued) Exercise 7.47 (continued) 1/ 2 - V From Eq. (6-40), U f = C L V 57.9 - 0.0994 = 0.45 0.0994 1/ 2 1/ 2 = 10.9 ft/s Since FLV < 0.1, from below Eq. (6-44), Ad /A = 0.1. Assume f = 0.80. 4VM V From Eq. (6-44), DT = fU f (1 - Ad / A)V 4(26, 230 / 3600)(18.2) = 0.80(10.9)(3.14)(1 - 0.1)(0.0994) 1/ 2 = 14.7 ft Column diameter at the top: From a simulation program, at 189oF and 33 psia, V = 0.1566 lb/ft3 and L =45.3 lb/ft3. V = 26,230 lbmol/h and L = 11,740 lbmol/h In Fig. 6.24, FLV LM L V = VMV L 1/ 2 = 11,740(34) 01566 . 26,230(34) 45.3 1/ 2 = 0.026 From Fig. 6.24, for 24-inch plate spacing, CF = 0.38 ft/s For Eqs. (6-40) and (6-42), use FHA = 1.0 and FF = 1.0. For methanol, = 17 dyne/cm Therefore, FST = (17/20)0.2 = 0.97 From Eq. (6.42), C = 0.97(1)(1)0.38 = 0.37 ft/s. From Eq. (6-40), - V Uf =C L V 1/ 2 = 0.37 45.3 - 01566 . 0.1566 1/ 2 = 6.28 ft/s Since FLV < 0.1, from below Eq. (6-44), Ad /A = 0.1. Assume f = 0.80. From Eq. (6-44), 4VM V DT = fU f (1 - Ad / A)V 1/ 2 4(26, 230 / 3600)(34) = 0.80(6.28)(3.14)(1 - 0.1)(0.1556) 1/ 2 = 21.2 ft Therefore, column would be swaged. (b) Sizing of reflux drum. Assume a liquid residence time in the reflux drum of 5 minutes (0.0833 h), half full. From Eq. (7-44), VV = From Eq. (7-46), V DT = V 1/ 3 2 LM L t 2(26,230)(34)(0.0833) = = 3,280 ft3 L 45.3 3, 280 = 3.14 1/ 3 = 10.1 ft From Eq. (7-46), H = 4DT = 4(10.1) = 40.4 ft Exercise 7.48 Subject: Given: Tray hydraulics for methanol-water separation' Data from Exercise 7.41. Percent of flooding. Tray pressure drop in psi. Weeping potential. Entrainment rate. Downcomer froth height. Find: (a) (b) (c) (d) (e) Analysis: Use the material balance from Exercise 7.41. Make calculations at the top and bottom trays. (a) At the top of the column, the reflux rate is 0.947(768) = 727 lbmol/h. The overhead vapor rate = 727 + 768 = 1,495 lbmol/h. Use the entrainment flooding correlation of Fig. 6.24, with densities from a simulation program, where the abscissa is, FLV LM L V = VM G L 1/ 2 = 727(30.9) 0.0703 1,495(30.9) 47.2 1/ 2 = 0.0188 (1) From Fig. 6.24, for 24-inch tray spacing, CF = 0.38 ft/s. Because FLV < 1, Ad /A = 0.1. FHA = 1.0, FF = 1.0, and since = 20 dynes/cm, FST = 1.0. From Eq. (6-24), C = FSTFFFHACF = (1)(1)(1)(0.38) = 0.38 ft/s From Eq. (6-40), U f = C L - V / V 1/ 2 = 0.38 47.2 - 0.0703 / 0.0703 1/ 2 = 9.84 ft / s From a rearrangement of Eq. (6-44), solving for the fraction of flooding, 4VMV 4(1,495 / 3,600)(30.9) f = 2 = 2 = 0.73 or 73% . . D U f 1 - Ad / A V 6 (9.84)(314)(1 - 01)(0.0703) At the bottom of the column, the boilup rate from Exercise 7.41 is 1.138(1,209) = 1,376 lbmol/h. The liquid rate leaving the bottom tray = 1,376 + 1,201.9 = 2,578 lbmol/h. Use the entrainment flooding correlation of Fig. 6.24, with densities and molecular weights from a simulation program, where the abscissa is, FLV LM L V = VM G L 1/ 2 = 2,578(181) 0.0408 . 1,376(18.8) 59.5 1/ 2 = 0.0473 From Fig. 6.24, for 24-inch tray spacing, CF = 0.37 ft/s. Because FLV < 1, Ad /A = 0.1. FHA = 1.0, FF = 1.0, and since = 20 dynes/cm, FST =(58/20)0.2 = 1.24 From Eq. (6-24), C = FSTFFFHACF = (1.24)(1)(1)(0.37) = 0.46 ft/s From Eq. (6-40), U f = C L - V / V 1/ 2 = 0.46 59.5 - 0.04 / 0.0408 1/ 2 = 17.6 ft / s From a rearrangement of Eq. (6-44), solving for the fraction of flooding, 4VMV 4(1,376 / 3,600)(18.8) f = 2 = 2 = 0.39 or 39% D U f 1 - Ad / A V 6 (17.6)(314)(1 - 01)(0.0408) . . Exercise 7.48 (continued) Analysis: (continued) (b) Vapor pressure drop per tray is given by Eq. (6-49), ht = hd + hl + h 2 uo V From Eq. (6-50), hd = 0186 2 . Co L Column area = A = D2/4 = 3.14(6)2/4 = 28.3 ft2 Take downcomer area = Ad = 0.1(28.3) = 2.83 ft2 Bubbling area = Aa = A - 2Ad = 28.3 - 2(2.83) = 22.6 ft2 Hole area = Ah = 0.1(22.6) = 2.26 ft2 (2) (3) Consider first, the conditions at the top of the column. From the continuity equation, uo = m/AhV uo = (1,495/3,600)(30.9)/[(2.26)(0.0703)] = 80.7 ft/s and Co = 0.73 80.7 2 0.0703 From Eq. (3), hd = 0186 . = 3.39 inches of liquid 0.732 47.2 Superficial velocity based on bubbling area = Ua = 80.7(2.26/22.6) = 7.97 ft/s From Eq. (6-53), Ks = U a V L - V 1/ 2 (4) = 7.97 0.0703 47.2 - 0.0703 1/ 2 = 0.31 ft/s From Eq. (6-52), e = exp(-4.257Ks0.91) = exp[-4.257(0.31)0.91] = 0.23 From Eq. (6-54), Cl = 0.362 + 0.317 exp(-3.5hw ) = 0.362 + 0.317 exp[-3.5(2.0)] = 0.362 Weir length = Lw = 42.5 in. Volumetric liquid rate = qL = 727(30.9)/[60(8.33)(47.2/62.4)] = 59.4 gpm From Eq. (6-51), hl = e hw + Cl qL Lw e 2/3 59.4 = 0.23 2 + 0.362 ( 42.5 ) 0.23 2/3 = 0.74 in. liquid From Eq. (6-55), with maximum bubble size = DH = 3/16 inch = 0.00476 m and liquid density = 47.2 lb/ft3 or 756 kg/m3, h = 6 / g L DBmax = 6(20 / 1000) / (9.8)(756)(0.00476) = 0.0034 m = 0.13 in. liquid From Eq. (2), ht = hd + hl + h = 3.39 + 0.74 + 0.13 = 4.26 inches liquid = 0.115 psi/tray. Consider next, the conditions at the bottom tray of the column. From the continuity equation, uo = m/AhV (4) uo = (1,376/3,600)(18.1)/[(2.26)(0.0408)] = 75.0 ft/s and Co = 0.73 75.02 0.0408 From Eq. (3), hd = 0186 . = 135 inches of liquid . 0.732 59.5 Superficial velocity based on bubbling area = Ua = 75.0(2.26/22.6) = 7.5 ft/s From Eq. (6-53), Ks = U a V L - V 1/ 2 = 7.5 0.0408 59.5 - 0.0408 1/ 2 = 0.196 ft/s From Eq. (6-52), e = exp(-4.257Ks0.91) = exp[-4.257(0.196)0.91] = 0.381 Exercise 7.48 (continued) Analysis: (b) (continued) From Eq. (6-54), Cl = 0.362 + 0.317 exp(-3.5hw ) = 0.362 + 0.317 exp[-3.5(2.0)] = 0.362 Weir length = Lw = 42.5 in. Volumetric liquid rate = qL = 2,578(18.1)/[60(8.33)(59.5/62.4)] = 97.9 gpm From Eq. (6-51), hl = e hw + Cl qL Lw e 2/3 97.9 = 0.381 2 + 0.362 ( 42.5 ) 0.381 2/3 = 1.22 in. liq From Eq. (6-55), with maximum bubble size = DH = 3/16 inch = 0.00476 m and liquid density = 59.5 lb/ft3 or 954 kg/m3, and = 58 dyne/cm = 0.058 kg/s2, h = 6 / g L DBmax = 6(0.058) / (9.8)(954)(0.00476) = 0.0078 m = 0.31 in. liquid From Eq. (2), ht = hd + hl + h = 1.35 + 1.22 + 0.31 = 2.88 inches liquid = 0.098 psi/tray. (c) Applying the criterion of Eq. (6-68) to the top tray, hd + h = 3.39 + 0.13 = 3.52 in., which is greater than hl = 0.74 in. Therefore, weeping will not occur. Applying the criterion of Eq. (6-68) to the bottom tray, hd + h = 1.35 + 0.31 = 1.66 in., which is greater than hl = 1.22 in. Therefore, weeping will not occur. (d) From Fig. 6.28 for the top tray, with FLV = 0.0192 and % flood = 73%, fractional entrainment = = 0.11, which is fairly high. From Fig. 6.28 for the bottom tray, with FLV = 0.0473 and % flood = 39%, fractional entrainment = = 0.0065, which is very low. The fractional entrainment is defined as, = e/(L + e). Rearranging, entrainment rate = e = L/(1 - ). At the top tray, e = 0.11(727)/(1 - 0.11) = 90 lbmol/h At the bottom tray, e = 0.0065(2,578)/(1 - 0.0065) = 16.9 lbmol/h (e) From Eqs. (6-70) and (6-72), the froth height in the downcomer = hdf =(ht + hl + hda) / 0.5 Estimate the head loss for flow under the downcomer from Eq. (6-71). Area for flow under the downcomer apron = Ada = Lwha = 42.5(2 - 0.5) = 63.8 in2 = 0.433 ft2 59.4 For the top tray, qL = 59.4 gpm. hda = 0.03 100(0.433) hdf = (4.26 + 0.74 + 0.056) / 0.5 = 10.1 in. liquid 2 = 0.056 in. liquid 2 98.3 For the bottom tray, qL = 98.3 gpm. hda = 0.03 100(0.433) hdf = (2.88 + 1.22 + 0.155) / 0.5 = 8.5 in. liquid = 0155 in. liquid . Exercise 7.49 Subject: Tray hydraulics for methanol-water separation. Given: Data from Exercise 7.41, except that feed rate is increased by 30%. Find: (a) Percent of flooding. (b) Tray pressure drop in psi. (c) Entrainment rate. (d) Downcomer froth height. Is operation acceptable or should packing be used in one or both sections? Analysis: Use the material balance from Exercise 7.41. Make calculations at the top and bottom trays. But take into account the 30% increase in all flow rates. (a) At the top of the column, the reflux rate is 1.3(0.947)(768) = 945 lbmol/h. The overhead vapor rate = 1.3[0.947(768) + 768] = 1,944 lbmol/h. Use the entrainment flooding correlation of Fig. 6.24, with densities from a simulation program, where the abscissa is, FLV LM L V = VM G L 1/ 2 = 945(30.9) 0.0703 1,944(30.9) 47.2 1/ 2 = 0.0192 (1) From Fig. 6.24, for 24-inch tray spacing, CF = 0.38 ft/s. Because FLV < 1, Ad /A = 0.1. FHA = 1.0, FF = 1.0, and since = 20 dynes/cm, FST = 1.0. From Eq. (6-24), C = FSTFFFHACF = (1)(1)(1)(0.38) = 0.38 ft/s From Eq. (6-40), U f = C L - V / V 1/ 2 = 0.38 47.2 - 0.0703 / 0.0703 1/ 2 = 9.84 ft / s From a rearrangement of Eq. (6-44), solving for the fraction of flooding, 4VMV 4(1,944 / 3,600)(30.9) f = 2 = 2 = 0.95 or 95% Unacceptable! D U f 1 - Ad / A V 6 (9.84)(314)(1 - 01)(0.0703) . . At the bottom of the column, the boilup rate from Exercise 7.41 is 1.3(1.138)(1,209) = 1,789 lbmol/h. The liquid rate leaving the bottom tray = 1.3(1,376 + 1,201.9) = 3,351 lbmol/h. Use the entrainment flooding correlation of Fig. 6.24, with densities and molecular weights from a simulation program, where the abscissa is, FLV LM L V = VM G L 1/ 2 = 3,351(181) 0.0408 . 1,789(18.8) 59.5 1/ 2 = 0.0473 From Fig. 6.24, for 24-inch tray spacing, CF = 0.37 ft/s. Because FLV < 1, Ad /A = 0.1. FHA = 1.0, FF = 1.0, and since = 20 dynes/cm, FST =(58/20)0.2 = 1.24 From Eq. (6-24), C = FSTFFFHACF = (1.24)(1)(1)(0.37) = 0.46 ft/s From Eq. (6-40), U f = C L - V / V 1/ 2 = 0.46 59.5 - 0.04 / 0.0408 1/ 2 = 17.6 ft / s From a rearrangement of Eq. (6-44), solving for the fraction of flooding, 4VMV 4(1,789 / 3,600)(18.8) f = 2 = 2 = 0.51 or 51% D U f 1 - Ad / A V 6 (17.6)(314)(1 - 01)(0.0408) . . Analysis: (continued) (b) Vapor pressure drop per tray is given by Eq. (6-49), ht = hd + hl + h u 2 V From Eq. (6-50), hd = 0186 o2 . Co L Column area = A = D2/4 = 3.14(6)2/4 = 28.3 ft2 Take downcomer area = Ad = 0.1(28.3) = 2.83 ft2 Bubbling area = Aa = A - 2Ad = 28.3 - 2(2.83) = 22.6 ft2 Hole area = Ah = 0.1(22.6) = 2.26 ft2 Consider first, the conditions at the top of the column. From the continuity equation, uo = m/AhV uo = (1,944/3,600)(30.9)/[(2.26)(0.0703)] = 105 ft/s and Co = 0.73 1052 0.0703 From Eq. (3), hd = 0186 . = 5.73 inches of liquid 0.732 47.2 Superficial velocity based on bubbling area = Ua = 105(2.26/22.6) = 10.4 ft/s From Eq. (6-53), Ks = U a V L - V 1/ 2 Exercise 7.49 (continued) (2) (3) (4) = 10.4 0.0703 47.2 - 0.0703 1/ 2 = 0.40 ft/s From Eq. (6-52), e = exp(-4.257Ks0.91) = exp[-4.257(0.40)0.91] = 0.16 From Eq. (6-54), Cl = 0.362 + 0.317 exp(-3.5hw ) = 0.362 + 0.317 exp[-3.5(2.0)] = 0.362 Weir length = Lw = 42.5 in. Volumetric liquid rate = qL = 945(30.9)/[60(8.33)(47.2/62.4)] = 77.2 gpm From Eq. (6-51), hl = e hw + Cl qL Lw e 2/3 77.2 = 0.16 2 + 0.362 ( 42.5 ) 0.16 2/3 = 0.61 in. liquid From Eq. (6-55), with maximum bubble size = DH = 3/16 inch = 0.00476 m and liquid density = 47.2 lb/ft3 or 756 kg/m3, h = 6 / g L DBmax = 6(20 / 1000) / (9.8)(756)(0.00476) = 0.0034 m = 0.13 in. liquid From Eq. (2), ht = hd + hl + h =5.73 + 0.61 + 0.13 = 6.47 inches liquid = 0.175 psi/tray. This is a very high pressure drop/tray. Consider next, the conditions at the bottom tray of the column. From the continuity equation, uo = m/AhV uo = (1,789/3,600)(18.1)/[(2.26)(0.0408)] = 97.9 ft/s and Co = 0.73 97.9 2 0.0408 From Eq. (3), hd = 0186 . = 2.28 inches of liquid 0.732 59.5 From Eq. (6-53), Ks = U a V L - V 1/ 2 (4) Superficial velocity based on bubbling area = Ua = 97.9(2.26/22.6) = 9.8 ft/s = 9.8 0.0408 59.5 - 0.0408 1/ 2 = 0.255 ft/s From Eq. (6-52), e = exp(-4.257Ks0.91) = exp[-4.257(0.255)0.91] = 0.293 Analysis: (b) (continued) Exercise 7.49 (continued) From Eq. (6-54), Cl = 0.362 + 0.317 exp(-3.5hw ) = 0.362 + 0.317 exp[-3.5(2.0)] = 0.362 Weir length = Lw = 42.5 in. Volumetric liquid rate = qL = 3,351(18.1)/[60(8.33)(59.5/62.4)] =127 gpm From Eq. (6-51), hl = e hw + Cl qL Lw e 2/3 127 = 0.293 2 + 0.362 ( 42.5 ) 0.293 2/3 = 1.08 in. liq From Eq. (6-55), with maximum bubble size = DH = 3/16 inch = 0.00476 m and liquid density = 59.5 lb/ft3 or 954 kg/m3, and = 58 dyne/cm = 0.058 kg/s2, h = 6 / g L DBmax = 6(0.058) / (9.8)(954)(0.00476) = 0.0078 m = 0.31 in. liquid From Eq. (2), ht = hd + hl + h =2.28 + 1.08 + 0.31 = 3.67 inches liquid = 0.125 psi/tray. This pressure drop is a slightly on the high side. (c) From Fig. 6.28 for the top tray, with FLV = 0.0192 and % flood = 95%, fractional entrainment = = 0.50, which is fairly high. From Fig. 6.28 for the bottom tray, with FLV = 0.0473 and % flood = 51%, fractional entrainment = = 0.017, which is low. The fractional entrainment is defined as, = e/(L + e). Rearranging, entrainment rate = e = L/(1 - ). At the top tray, e = 0.50(945)/(1 - 0.50) = 945 lbmol/h, which is very excessive. At the bottom tray, e = 0.017(3,351)/(1 - 0.017) = 58 lbmol/h (d) From Eqs. (6-70) and (6-72), the froth height in the downcomer = hdf =(ht + hl + hda) / 0.5 Estimate the head loss for flow under the downcomer from Eq. (6-71). Area for flow under the downcomer apron = Ada = Lwha = 42.5(2 - 0.5) = 63.8 in2 = 0.433 ft2 77.2 For the top tray, qL = 77.2 gpm. hda = 0.03 100(0.433) hdf = (6.47 + 0.61 + 0.095) / 0.5 = 14.4 in. liquid 2 = 0.095 in. liquid 128 For the bottom tray, qL = 128 gpm. hda = 0.03 100(0.433) hdf = (3.67 + 1.08 + 0.262) / 0.5 = 10.0 in. liquid 2 = 0.262 in. liquid The new operation is not acceptable in the rectifying section because % flooding is 95%, causing very high entrainment. Should retrofit that section with structured packing. Bottom section is operable with the trays if the higher pressure drop is acceptable. Exercise 7.50 Subject: Determination of HETP from experimental data for a packed column separating benzene and dichloroethane at total reflux. Given: 10 feet of packing, with operation at 1 atm. Vapor-liquid equilibrium data in terms of benzene mole fractions. Measured benzene mole fractions are xD = 0.653 and xB = 0.298. Find: HETP in inches. Limitations of the use of the HETP for design. Analysis: In the plot below, the y-x equilibrium data and the operating line (45o line for total reflux condition) are plotted and equilibrium stages are stepped off between x = 0.298 and 0.653. The number of stages = Nt = 14.4. From Eq. (6-73), HETP = 10/14.4 = 0.7 ft. This value must be used with caution for design, because the HETP is a function of the L/V ratio. The value of 0.7 ft applies only to L/V =1. Exercise 7.51 Subject: atm. Design of a packed column for separation of a mixture of ethanol and water at 1 Given: Bubble-point feed of 10 mol% ethanol. Bottoms of 1 mol% ethanol and distillate of 80 mol% ethanol. Reflux ratio at 1.5 times minimum. Equilibrium data from Exercise 7.29. Assumptions: Constant molar overflow. Find: (a) (b) (c) (d) (e) Equilibrium stages above and below feed. Number of transfer units above and below the feed. Height of plate column for 18-inch tray spacing and 80% tray efficiency. Height of packing for HOG = 1.2 ft. HTU from HTU for benzene-toluene system. Analysis: By material balance, with a basis of 100 moles of feed, Total material balance, F = 100 = D + B Benzene material balance, 0.10F = 10 = 0.80D + 0.01B Solving these two equations, D = 11.39 moles and B = 88.61 moles. (a) Below is the McCabe-Thiele diagram showing the equilibrium curve from the data of Exercise 7.29, a vertical q-line for the saturated liquid feed at xF = 0.10, and the construction of the rectification section operating line for determining the minimum reflux ratio. The slope of that line is, (0.80 - 0.44)/(0.80 - 0.10) = 0.514 = Lmin/V. From a rearrangement of Eq. (7-7), Rmin = Lmin / V 0.514 = = 1.058 and R = 1.5 Rmin = 1.5(1.058) = 1.59 1 - Lmin / V 1 - 0.514 Slope of the rectification section operating line is, from Eq. (7-7), L/V = 1.59/(1 + 1.59) = 0.614. In the McCabe-Thiele diagrams below, one for the higher concentration region and one for the lower concentration region, a rectification section operating line of this slope is drawn that passes through the 45o line at x = xD = 0.8. A stripping section operating line is drawn that extends from the intersection of the rectifying operating line and the q-line to the 45o line at x = xB = 0.01. The stages are stepped off, with the optimal feed stage location. The result is 3 equilibrium stages in the stripping section, including one for the partial reboiler, and 13.8 equilibrium stages in the rectifying section. Analysis: (a) (continued) Exercise 7.51 (continued) Exercise 7.51 (continued) Analysis: (a) (continued) Exercise 7.51 (continued) Analysis: (a) (continued) Exercise 7.51 (continued) Analysis: (continued) (b) Because the value of HOG is given in part (d), calculate NOG for each section, using from Table 6.7 for EM diffusion, N OG = dy y -y * For the stripping section, the first stage is the partial reboiler. Therefore, from the McCabeThiele diagram, starting from Stage 2, the limits on y are y = 0.09 at x = 0.03 to y = 0.37 at x = 0.10. Using the trapezoidal method to solve the integral, ( NOG )SS = y (0.23 - 0.09) (0.37 - 0.23) = + = 2.4 0.10 0.14 ( y *- y) For the rectifying section, the limits on y are y = 0.37 at x = 0.1 to y = 0.80. Using the trapezoidal method to solve the integral, ( NOG ) RS = y (0.42 - 0.37) (0.665 - 0.42) ( 0.80 - 0.665 ) = + + = 12.4 0.084 0.063 0.017 ( y *-y) (c) For 15.8 equilibrium stages at 80% tray efficiency, need 15.8/0.8 = 19.8 or 20 trays. If they are on 18-inch spacing, column height = 19(1.5) + 4 + 10 = 42.5 ft. (d) For HOG = 1.2 ft, packed height = 1.2(2.4 + 12.4) = 18 ft. Add 14 ft for disengagement at the top and sump at the bottom. Need 18 + 14 = 32 ft column. (e) We would need data on individual values of HL and HG. From Eq. (6-132), 1 H L is proportional to DL 1/ 2 . Therefore, can ratio with the liquid diffusivities. V V -3/ 4 From Eq. (6-133), H G is proportional to V V DV -1/ 3 . Therefore, we can ratio off these properties. Then from Eq. (7.52), HOG = HG + mH L , where m = dy / dx , which varies over the equilibrium curve. Exercise 7.52 Subject: Doubling plant capacity for distillation of methanol-water by adding a packed column. Given: Conditions to duplicate those of Exercise 7.41. Design for 70% of flooding. Two packings to be considered: 1. 50-mm plastic NOR PAC rings and 2. Montz metal B1-300. Assumptions: Constant molar overflow. Partial reboiler is an equilibrium stage. Find: (a) Liquid holdup. (b) Column diameter. (c) HOG. (d) Packed height. (e) Pressure drop. Advantages of packed column over trayed column. Preferred packing. Analysis: First, convert the performance data for feed and product flow rates and compositions from mass units into mole units, using molecular weights of 32.04 for methanol and 18.02 for water. The results are as follows: Component Methanol Water Total: Flow rate, lbmol/h: Feed Distillate 709.1 702.7 1260.8 65.3 1969.9 768.0 Bottoms 6.4 1195.5 1201.9 Mole fraction: Feed Distillate 0.360 0.915 0.640 0.085 1.000 1.000 Bottoms 0.0053 0.9947 1.0000 The results of Exercise of 7.41 show that the observed reflux and boilup ratios (0.947 and 1.138, respectively) are consistent with the separation achieved. Therefore, at the top of the column, L = RD = 0.947(768) = 727 lbmol/h, and V = L + D = 727 + 768 = 1,495 lbmol/h. At the bottom of the column, V = VB B = 1138(1,2019) = 1,367.8 lbmol / h, and L = V + B = 1,367.8 + 1,2019 = 2,569.7 lbmol / h . . . Before computing liquid holdup, we need to compute column diameter. (b) Column diameter based on conditions at the top of the column, where T = 154oF and P = 14.7 psia = 1 atm. Use Figs. 6.36a, b, and c to compute flooding velocity, following Example 6.13. From the ideal gas law, V = PMV/RT = (1)(30.9)/[0.730(154 + 460)] = 0.0689 lb/ft3 For the liquid, which is mainly methanol, use L = 0.77 g/cm3 = 48 lb/ft3 and L = 0.35 cP. LM L The abscissa in Fig. 6.36a = X = VM V V L 1/ 2 727(30.9) 0.0689 = 1495(30.9) 48 1/ 2 = 0.0184 From Fig. 6.36a, Y at flooding = 0.21; from Fig. 6.36b, for water/ L = 1/0.77 = 1.3, f{L } = 1.6; and from Fig. 6.36c, for L = 0.35 cP, f{L } = 0.8. From Table 6.8, FP = 14 for NOR PAC and FP = 33 for Montz. Analysis: (b) (continued) Exercise 7.52 (continued) g H2 O (L) FP V 1 f { L } f { L } 2 From a rearrangement of the ordinate of Fig. 6.36a, uo = Y 32.2 62.4 1 = 341 (ft/s)2 and uo = 18.5 ft/s 14 0.0689 16(0.8) . 32.2 62.4 1 2 For Montz, uo = 0.21 = 145 (ft/s)2 and uo = 12.0 ft/s 33 0.0689 16(0.8) . For fraction of flooding = f = 0.7, uV = uo f = 18.5(0.7) = 13.0 ft/s for NOR PAC, uV = uo f = 12.0(0.7) = 8.4 ft/s for Montz. 2 For NOR PAC, uo = 0.21 4VM V From Eq. (6-103), column diameter = DT = fu0 V 4(1495 / 3600)(30.9) For NORPAC, DT = 13(3.14)(0.0689) 1/ 2 1/ 2 4VM V = uV V 1/ 2 = 4.3 ft. 1/ 2 4(1495 / 3600)(30.9) For Montz, DT = = 5.3 ft 8.4(3.14)(0.0689) Column diameter based on conditions at the bottom of the column, where T = 207oF and P = 15.5 psia = 1.05 atm. Use Figs. 6.36a, b, and c to compute flooding velocity, following Example 6.13. From the ideal gas law, V = PMV/RT = (1.05)(18.1)/[0.730(207 + 460)] = 0.039 lb/ft3 For the liquid, which is mainly water, use L = 1.00 g/cm3 = 62.4 lb/ft3 and L = 0.27 cP. LM L The abscissa in Fig. 6.36a = X = VM V V L 1/ 2 2569.7(18.1) 0.039 = 1367.8(18.1) 62.4 1/ 2 = 0.047 From Fig. 6.36a, Y at flooding = 0.17; from Fig. 6.36b, for water/ L = 1/1 = 1, f{L }= 1; and from Fig. 6.36c, for L = 0.27 cP, f{L } = 0.77. From Table 6.8, FP = 14 for NOR PAC and FP = 33 for Montz. g H2 O (L) 1 2 From a rearrangement of the ordinate of Fig. 6.36a, uo = Y FP V f { L } f { L } 32.2 62.4 1 = 812 (ft/s)2 and uo = 28.5 ft/s 14 0.039 1(0.77) 32.2 62.4 1 2 For Montz, uo = 0.17 = 345 (ft/s)2 and uo = 18.6 ft/s 33 0.039 1(0.77) For fraction of flooding = f = 0.7, uV = uo f = 28.5(0.7) = 20.0 ft/s for NOR PAC, uV = uo f = 18.6(0.7) = 13.0 ft/s for Montz. 2 For NOR PAC, uo = 0.17 4VM V From Eq. (6-103), column diameter = DT = fu0 V 1/ 2 4VM V = uV V 1/ 2 Analysis: (b) (continued) Exercise 7.52 (continued) 1/ 2 4(1367.8 / 3600)(18.1) For NOR PAC, DT = 20(3.14)(0.039) = 3.4 ft. 1/ 2 4(1367.8 / 3600)(18.1) For Montz, DT = = 4.2 ft 13.0(3.14)(0.039) (a) For liquid holdup estimates, assume the column operates in the preloading region. Therefore, the holdup is independent of the gas rate. Follow Example 6.12. Pertinent packing characteristics from Table 6.8: Ch Packing a, m2/m3 a, ft2/ft3 NOR PAC 86.8 26.5 0.947 0.651 Montz 300 91.4 0.930 0.482 Use Eqs. (6-97) to (6-101), which requires calculating the liquid Reynolds and Froude numbers. Liquid holdup based on conditions at the top of the column. By the continuity equation, superficial liquid velocity is, uL = LML /LS = (727/3600)(30.9)/ [48(3.14)(DT)2/4] = 0.166/(DT)2 ft/s 2 0166 / DT 48 . 33900 u L L From Eq. (6-98), N Re L = = = 2 L a (0.35)(0.000672)a DT a From Eq. (6-99), N FrL 2 2 u L a ( 0.166 / DT ) a a = = = 0.000856 4 g 32.2 DT 2 Packing DT, ft uL, ft/s a, ft2/ft3 NOR PAC 4.3 0.00898 26.5 69.2 6.63 x 10-5 Montz 5.3 0.00591 91.4 13.2 9.91 x 10-5 0. 0 Since both values of N Re L > 5, use Eq. (6-101), ah / a = 0.85Ch N Re25 N Fr.1 L L For NOR PAC, ah / a = 0.85(0.651)(69.2) 0.25 (6.63 10 -5 ) 0.1 = 0.61 For Montz, ah / a = 0.85(0.482)(13.2) 0.25 (9.91 10 -5 ) 0.1 = 0.31 From Eq. (6-97), the fractional liquid holdup is hL = 12 6.63 10 -5 For NOR PAC, hL = (12) 69.2 1/ 3 N Re L N FrL N FrL N Re L 1/ 3 ah a 2/3 ( 0.61) 2/3 = 0.0162 m3/m3 9.91 10-5 2/3 For Montz, hL = (12) ( 0.31) = 0.0241 m3/m3 13.2 These liquid holdups at the top of the column are quite small. 1/ 3 Exercise 7.52 (continued) Analysis: (a) (continued) Liquid holdup based on conditions at the bottom of the column. By the continuity equation, superficial liquid velocity is, uL = LML /LS = (2570/3600)(18.1)/ [62.4(3.14)(DT)2/4] = 0.264/(DT)2 ft/s 2 0.264 / DT 62.4 90800 u L L From Eq. (6-98), N Re L = = = 2 L a (0.27)(0.000672)a DT a From Eq. (6-99), N FrL 2 2 u L a ( 0.264 / DT ) a a = = = 0.00216 4 g DT 32.2 2 Packing NOR PAC Montz DT, ft 3.4 4.2 uL, ft/s 0.0228 0.0150 a, ft2/ft3 26.5 91.4 N Re L 296 56.3 4.28 x 10-4 6.39 x 10-4 N FrL 0. 0 Since both values of N Re L are > 5, use Eq. (6-101), ah / a = 0.85Ch N Re25 N Fr.1 L L . For NOR PAC, ah / a = 0.85(0.651)(296) 0.25 (4.28 10 -4 ) 0.1 = 1057 For Montz, ah / a = 0.85(0.482)(56.3) 0.25 (6.39 10 -4 ) 0.1 = 0.54 From Eq. (6-97), the fractional liquid holdup is hL = 12 4.28 10 -4 For NOR PAC, hL = (12) 69.2 1/ 3 N FrL N Re L 1/ 3 ah a 2/3 (1.057 ) 2/3 = 0.0269 m3/m3 6.39 10 -4 2/3 For Montz, hL = (12) ( 0.54 ) = 0.0341 m3/m3 56.3 These liquid holdups at the bottom of the column are quite small. mV (c) From Eq. (7.52), HOG = HG + H L , where we must use the slope, m, of the equilibrium L curve instead of a K-value because the equilibrium curve is curved. At the top of the column: Near a mole fraction, xD , of 0.9, m = (0.9780 - 0.9359)/0.10 = 0.42 Estimate HL from Eq. (6-132), using the following properties and parameters: NOR PAC Montz CL 1.080 1.165 3 3 hL , m /m 0.0162 0.0241 0.947 0.930 , m3/m3 uL, ft/s 0.00898 0.00591 2 3 a, m /m 86.8 300 ah, m2/m3 52.9 93 1/ 3 Analysis: (c) at the top of the column (continued) Exercise 7.52 (continued) Need an estimate of the diffusivity of methanol in water at high concentrations of methanol. From Example 3.7, the diffusivity of methanol in water at a mole fraction of 0.8 and 25oC is 1.5 x 10-5 cm2/s. Use Eq. (3-39) to correct this for temperature to 170oF or 350 K and viscosity. For methanol, viscosity at 25oC (298 K) = 0.55 cP and at 350 K, viscosity = 0.32 cP. Therefore, 350 0.55 DMeOH = 150 10-5 . = 3.0 10-5 cm2/s 298 0.32 From Eq. (6-132), for NOR PACK, using SI (instead of American Engineering) units, 1 1 HL = CL 12 = 0.174 1/ 6 4hL DL au L 1/ 2 uL a a 1 1 = 1.08 12 aPh 1/ 6 4(0.0162)(0.947) (3.0 10-9 )(86.8)(0.00274) 1/ 2 0.00274 86.8 a aPh a m aPh To compute (aPh /a), we need for the liquid phase, the Reynolds, Weber, and Froude numbers from equations (6-138), (6-139), and (6-140), respectively, based on the packing hydraulic diameter from (6-137). d h = packing hydraulic diameter = 4 0.947 =4 = 0.0436 m = 0.143 ft 86.8 a Reynolds number = N ReL , h = u L d h L (0.00898)(0.143)(48) = = 262 [(0.35)(0.000672)] L Take the surface tension of methanol at 154oF as = 17 dynes/cm = 0.001165 lbf/ft or 0.001165(32.2) = 0.0375 lbm/s2 Weber number = N WeL ,h 2 u L L d h ( 0.00898 ) (48)(0.143) = = = 0.0148 0.0375 2 2 Froude number = N FrL , h ( 0.00898) = 0.0000175 u2 = L = gd h 32.2(0.143) From (6-136), aPh -1/ 2 = 1.5 ( ad h ) N Re L ,h a = 1.5 [ (86.8)(0.0436) ] ( ) (N -0.2 -1/ 2 We L , h ) (N ) 0.75 FrL , h -0.45 ( 262 ) ( 0.0148) ( 0.0000175) -0.2 0.75 -0.45 = 1.485 Exercise 7.52 (continued) Analysis: (c) at the top of the column (continued) HL = 0.174 1 = 0.117 m 1.485 From Eq. (6-132), for Montz, using SI units, 1 1 HL = CL 12 = 0.0253 1/ 6 4hL DL au L m 1/ 2 uL a a 1 1 = 1.165 12 aPh 1/ 6 4(0.0241)(0.930) (3.0 10-9 )(300)(0.00180) 1/ 2 0.00180 300 a aPh a aPh To compute (aPh /a), we need for the liquid phase, the Reynolds, Weber, and Froude numbers from equations (6-138), (6-139), and (6-140), respectively, based on the packing hydraulic diameter from (6-137). d h = packing hydraulic diameter = 4 0.930 =4 = 0.0124 m = 0.0407 ft a 300 Reynolds number = N ReL ,h = uL d h L (0.00591)(0.0407)(48) = = 49.1 L [(0.35)(0.000672)] Weber number = N WeL ,h = 2 u L L d h ( 0.00898 ) (48)(0.407) = = 0.00421 0.0375 2 2 Froude number = N FrL ,h ( 0.00898) = 0.0000615 u2 = L = gd h 32.2(0.0407) From (6-136), aPh -1/ 2 = 1.5 ( ad h ) N Re L ,h a = 1.5 [ (300)(0.0124)] ( ) (N -0.2 We L , h ) (N ) 0.75 FrL , h -0.45 -1/ 2 ( 49.1) ( 0.00421) ( 0.0000615 ) -0.2 0.75 -0.45 = 0.47 HL = 0.0253 1 = 0.054 m 0.47 Exercise 7.52 (continued) Analysis: (c) at the top of the column (continued) Now estimate the value of HG from Eq. (6-133). Estimate the gas diffusivity from Eq. (3-36), where from Table 3.1, . . . V = 131 for water, and V = 15.9 + 4( 2.31) + 611 = 312 for methanol and molecular weights are 32 for methanol and 18 for water. 0.00143T 1.75 DV = DAB = 1/ 2 1/ 3 1/ 3 2 2 P ( V ) A + ( V )B (1/ M A ) + (1/ M B ) 2 1/ 3 1/ 3 2 (13.1) + ( 31.2 ) (1/18) + (1/ 32) Use the following additional properties and parameters, together with = 1.1 x 10-2 cP and V u = 0.0688 lb/ft3 , where from Eqs. (6-134) and (6-135), N ReV = V V and N ScV = V aV V DV (1) CV uV , m/s (NSc)V (NRe)V NOR PAC 0.322 3.96 0.36 4,560 Montz 0.422 2.56 0.36 850 = 0.00143(350)1.75 1/ 2 = 0.28 cm 2 /s From Eq. (6-133), for NOR PAC, using SI units, 1 1/ 2 4 HG = ( - hL ) CV a4 1/ 2 (N ) (N ) -3/ 4 -1/ 3 ReV ScV uV a DV aPh -3/ 4 1 1/ 2 4(0.947) = ( 0.947 - 0.0162 ) 0.322 86.84 For Montz, 1/ 2 ( 4, 560 ) ( 0.36 ) -1/ 3 (3.96)(86.8) = 0.34 m (0.28 10-4 )(71.2) 1 1/ 2 4(0.930) HG = ( 0.947 - 0.0241) 0.422 3004 1/ 2 ( 850 ) -3/ 4 ( 0.36 ) -1/ 3 (2.56)(300) = 0.16 m (0.28 10-4 )(78) At the top of the column, vapor rate = V =1,495 lbmol/h and liquid rate = L = 727 lbmol/h. Therefore, mV/L = (0.42)(1,495)/(727) = 0.864 mV From Eq. (7.51), for NOR PAC, H OG = H G + H L = 0.34 + (0.864)(0.117) = 0.44 m = 1.5 ft L Exercise 7.52 (continued) Analysis: (c) at the top of the column (continued) mV For Montz, H OG = H G + H L = 0.16 + (0.864)(0.054) = 0.21 m = 0.68 ft L At the bottom of the column: Near a mole fraction, xB , of 0.005, m = 2.5 Estimate HL from Eq. (6-123), using the following properties and parameters: CL hL , m3/m3 , m3/m3 uL, ft/s a, m2/m3 ah, m2/m3 NOR PAC 1.080 0.0269 0.947 0.0228 86.8 91.7 Montz 1.165 0.0341 0.930 0.0150 300 162 Need an estimate of the diffusivity of methanol in water at low concentrations of methanol. From Example 3.7, the diffusivity of methanol in water at infinite dilution at 25oC is 1.5 x 10-5 cm2/s. Use Eq. (3-39) to correct this for temperature to 207oF or 370 K and viscosity. For water, viscosity at 25oC (298 K) = 0.90 cP and at 370 K, viscosity = 0.26 cP. Therefore, 370 0.90 DMeOH = 150 10 -5 . = 6.4 10-5 cm2/s 298 0.26 From Eq. (6-132), for NOR PACK, using SI (instead of American Engineering) units, 1 1 HL = CL 12 = 0.456 1/ 6 4hL DL au L 1/ 2 uL a a 1 1 = aPh 1.08 12 1/ 6 4(0.0269)(0.947) (6.4 10 -9 )(86.8)(0.0228) 1/ 2 0.0228 86.8 a aPh a m aPh To compute (aPh /a), we need for the liquid phase, the Reynolds, Weber, and Froude numbers from equations (6-138), (6-139), and (6-140), respectively, based on the packing hydraulic diameter from (6-137). d h = packing hydraulic diameter = 4 0.947 =4 = 0.0436 m = 0.143 ft a 86.8 Reynolds number = N Re L , h = u L d h L (0.0228)(0.143)(62.4) = = 1120 L [(0.27)(0.000672)] Exercise 7.52 (continued) Analysis: (c) at the bottom of the column (continued) Take the surface tension of water at 207oF as = 60 dynes/cm = 0.00411 lbf/ft or 0.00411(32.2) = 0.132 lbm/s2 Weber number = N WeL ,h ( 0.0228 ) (62.4)(0.143) = 0.035 u 2 d = L L h = 0.132 2 2 Froude number = N FrL , h ( 0.0228 ) = 0.000113 u2 = L = gd h 32.2(0.143) From (6-136), aPh -1/ 2 = 1.5 ( ad h ) N Re L ,h a = 1.5 [ (86.8)(0.0436)] ( ) (N -0.2 -1/ 2 We L , h ) (N ) 0.75 FrL , h -0.45 (1120 ) ( 0.035 ) ( 0.000113) -0.2 0.75 -0.45 = 0.92 HL = 0.456 1 = 0.50 m 0.92 From Eq. (6-132), for Montz, using SI units, 1 1 HL = CL 12 = 0.0595 1/ 6 4hL DL au L m 1/ 2 uL a a 1 1 = aPh 1.165 12 1/ 6 4(0.0341)(0.930) (6.4 10 -9 )(300)(0.0150) 1/ 2 0.0150 300 a aPh a aPh To compute (aPh /a), we need for the liquid phase, the Reynolds, Weber, and Froude numbers from equations (6-138), (6-139), and (6-140), respectively, based on the packing hydraulic diameter from (6-137). d h = packing hydraulic diameter = 4 0.930 =4 = 0.0124 m = 0.0407 ft a 300 Reynolds number = N ReL , h = u L d h L (0.0150)(0.0407)(62.4) = = 210 L [(0.27)(0.000672)] Exercise 7.52 (continued) Analysis: (c) at the bottom of the column (continued) Weber number = N WeL ,h ( 0.0150 ) (62.4)(0.407) = 0.0433 u 2 d = L L h = 0.132 2 2 ( 0.0150 ) = 0.000172 uL = = gd h 32.2(0.0407) 2 Froude number = N FrL , h From (6-136), aPh -1/ 2 = 1.5 ( ad h ) N Re L ,h a = 1.5 [ (300)(0.0124)] ( ) (N -0.2 We L , h ) (N ) 0.75 FrL , h -0.45 -1/ 2 ( 210 ) ( 0.0433) ( 0.000172 ) -0.2 0.75 -0.45 = 1.25 HL = 0.0595 1 = 0.048 m 1.25 Now estimate the value of HG from Eq. (6-133). Estimate the gas diffusivity from Eq. (3-36), where from Table 3.1, . . . V = 131 for water, and V = 15.9 + 4( 2.31) + 611 = 312 for methanol and molecular weights are 32 for methanol and 18 for water. DV = DAB = 0.00143T 1.75 2 P (1/ M A ) + (1/ M B ) 1/ 2 ( V ) 1/ 3 A +( V ) 1/ 3 B 2 2 1/ 3 1/ 3 2 (13.1) + ( 31.2 ) (1/18) + (1/ 32) Use the following additional properties and parameters, together with V = 1.2 x 10-2 cP and V u = 0.0370 lb/ft3 , where from Eqs. (6-134) and (6-135), N ReV = V V and N ScV = V aV V DV (1) CV uV , m/s (NSc)V (NRe)V From Eq. (6-133), for NOR PAC, = 0.00143(370)1.75 1/ 2 = 0.31 cm 2 /s NOR PAC 0.322 6.10 0.65 3,470 Montz 0.422 3.96 0.65 652 Analysis: (c) at the bottom of the column (continued) From Eq. (6-133), for NOR PAC, using SI units, 1 1/ 2 4 HG = ( - hL ) CV a4 = 1/ 2 Exercise 7.52 (continued) (N ) (N ) -3/ 4 ReV ScV -1/ 3 uV a DV aPh -3/ 4 1 1/ 2 4(0.947) ( 0.947 - 0.0269 ) 0.322 86.84 1/ 2 ( 3, 470 ) ( 0.65 ) -1/ 3 (6.10)(86.8) = 0.69 m (0.34 10-4 )(44.2) For Montz, HG = 1 1/ 2 4(0.930) ( 0.947 - 0.0341) 0.422 3004 1/ 2 ( 652 ) -3/ 4 ( 0.65) -1/ 3 (3.96)(300) = 0.073 m (0.34 10-4 )(207) At the bottom of the column, vapor rate = V =1,368 lbmol/h and liquid rate = L = 2,570 lbmol/h. Therefore, mV/L = (2.50)(1,368)/(2,570) = 1.33 mV From Eq. (7.51), for NOR PAC, H OG = H G + H L = 0.69 + (1.33)(0.50) = 1.36 m = 4.4 ft L mV For Montz, H OG = H G + H L = 0.073 + (1.33)(0.048) = 0.14 m = 0.45 ft L (d) The packed height is given by Eq. (6-127), lT = HOGNOG Calculate NOG for each section, using from Table 6.7 for EM diffusion, N OG = dy y -y Apply this equation to the rectifying section and to the stripping section, using the McCabeThiele diagrams developed in Exercise 7.41. * For the rectifying section, the limits on y from the feed to the distillate are 0.64 to 0.915. The difference, y*-y is almost constant at an average value of approximately 0.065. Therefore, (NOG)RS = y/( y*-y) = (0.915-0.64)/0.065 = 4.2. Therefore, Packed height for NOR PAC = (1.5)(4.2) = 6.3 ft Packed height for Montz = (0.68)(4.2) = 2.9 ft For the stripping section, after accounting for the partial reboiler, the limits on y from the bottom to the feed are 0.05 to 0.68. The difference, y*-y varies from 0.12 at y = 0.05, to 0.23 at y = 0.17, to 0.19 at y = 0.40, to 0.09 at y = 0.59, to 0.03 at y = 0.68 Using the trapezoidal method to solve the integral, y (0.17 - 0.05) (0.40 - 017) 0.59 - 0.40 (0.68 - 0.59) . N OG RS = = + + + = 4.6 y *-y 0175 . 0.21 014 . 0.06 Analysis: (d) (continued) Exercise 7.52 (continued) Packed height for NOR PAC = (4.4)(4.6) = 20.2 ft Packed height for Montz = (0.45)(4.6) = 2.1 ft The total packed height for both sections are: 6.3 + 20.2 = 26.5 ft for NOR PAC 2.9 + 2.1 = 5.0 ft for Montz However, these heights are approximate because they are based on values of HOG that are assumed constant over the packed sections, which is not the case because m varies. (e) Compute pressure drops for the packed sections from the correlation of Billet and Schultes. The pressure drop per unit height of packed bed is given by Eq. (6-115), 1/ 2 Po - hL P 13300 = exp ( N FrL ) lT lT a 3/ 2 This equation will be used for each packing in each section. 3/ 2 (1) Po a u2 1 = o 3 V V lT 2 g c KW The required packing parameters are as follows from Table 6.8 and above: NOR PAC Montz 2 3 FP , ft /ft 14 33 2 3 a, m /m 86.8 300 0.947 0.930 , m3/m3 Ch 0.651 0.482 0.350 0.295 Cp From Eq. (6-110), From Eq. (6-112), DP = 6 (2) 1- 1 - 0.947 =6 = 0.00366 m or 0.012 ft for NOR PAC a 86.8 1- 1 - 0.930 DP = 6 =6 = 0.0014 m or 0.00459 ft for Montz a 300 Rectifying section of the column: For NOR PAC , column diameter = DT = 4.3 ft From Eq. (6-111), 1 2 1 DP 2 1 0.012 = 1+ = 1+ = 1035 . KW 3 1 - DT 3 1 - 0.947 4.3 Therefore, KW = 1/1.035 = 0.966 Analysis: (e) (continued) Exercise 7.52 (continued) For Montz, column diameter = DT = 5.3 ft 1 2 1 DP 2 1 0.00459 From Eq. (6-111), = 1+ = 1+ = 1008 . KW 3 1 - DT 3 1 - 0.930 5.3 Therefore, KW = 1/1.008 = 0.992 Superficial gas velocity at 70% of flooding from above = 13 ft/s for NOR PAC. Superficial gas velocity at 70% of flooding from above = 8.4 ft/s for Montz Gas viscosity = 0.011 cP and gas density = 0.0688 lb/ft3. From Eq. (6-114), for NOR PAC, u D (13)(0.012)(0.0688) N ReV = V P V KW = (0.966) = 26,500 (1 - ) V (1 - 0.947 ) (0.011)(0.000672) For Montz, N ReV = (8.4)(0.00459)(0.0688) (0.992) = 5, 090 (1 - 0.930 ) (0.011)(0.000672) From Eq. (6-113), for NOR PAC, o = Cp 64 1.8 64 1.8 + 0.08 = 0.350 + = 0.280 N ReV N ReV 26, 500 26,5000.08 64 18 . + = 0.272 5,090 5,0900.08 For Montz, o = 0.295 From Eq. (2), for NOR PAC, 2 Po a uV V 1 86.8(0.3048) 132 (0.0688) 1 = o 3 = 0.280 = 1.63 lbf/ft3 or 0.0113 psi/ft 3 lT 2 g c KW 0.947 2(32.2) 0.966 For Montz, Po 300(0.3048) 8.4 2 (0.0688) 1 = 0.272 = 2.35 lbf/ft3 or 0.0163 psi/ft 3 lT 0.930 2(32.2) 0.992 From Eq. (1), for NOR PAC, with hL = 0.0162 m3/m3 and liquid Froude number = 6.63x10-5 P 0.947 - 0.0162 = 0.0113 lT 0.947 1.5 exp 1/ 2 13300 6.63 10 -5 ) = 0.0126 psi/ft 3/ 2 ( 86.8 Exercise 7.52 (continued) Analysis: (e) Rectifying section (continued) From Eq. (1), for Montz, with hL = 0.0241 m3/m3 and liquid Froude number =9.91x10-5 P 0.930 - 0.0241 = 0.0163 lT 0.930 1.5 exp 1/ 2 13300 9.91 10-5 ) = 0.0161 psi/ft 3/ 2 ( 300 Stripping section of the column: For NOR PAC , column diameter = DT = 3.4 ft 1 2 1 DP 2 1 0.012 From Eq. (6-111), = 1+ = 1+ = 1044 . KW 3 1 - DT 3 1 - 0.947 3.4 Therefore, KW = 1/1.044 = 0.957 For Montz, column diameter = DT = 4.2 ft 1 2 1 DP 2 1 0.00459 From Eq. (6-102), = 1+ = 1+ = 1010 . KW 3 1 - DT 3 1 - 0.930 4.2 Therefore, KW = 1/1.010 = 0.990 Superficial gas velocity at 70% of flooding from above = 20 ft/s for NOR PAC. Superficial gas velocity at 70% of flooding from above = 13 ft/s for Montz Gas viscosity = 0.012 cP and gas density = 0.0370 lb/ft3. From Eq. (6-114), for NOR PAC, N ReV = uV DP V (20)(0.012)(0.0370) KW = (0.957) = 19, 900 (1 - ) V (1 - 0.947 ) (0.012)(0.000672) For Montz, (13)(0.00459)(0.0370) N ReV = (0.990) = 3,870 (1 - 0.930 ) (0.012)(0.000672) From Eq. (6-113), for NOR PAC, o = Cp 64 1.8 64 1.8 + 0.08 = 0.350 + = 0.287 N ReV N ReV 19,900 19,9000.08 64 18 . + = 0.279 3,870 3,8700.08 For Montz, o = 0.295 Exercise 7.52 (continued) Analysis: (e) Stripping section (continued) From Eq. (2), for NOR PAC, 2 Po a uV V 1 86.8(0.3048) 202 (0.0370) 1 = o 3 = 0.287 = 2.15 lbf/ft3 or 0.0149 psi/ft 3 lT 2 g c KW 0.947 2(32.2) 0.957 For Montz, Po 300(0.3048) 8.4 2 (0.0370) 1 = 0.279 = 311 lbf/ft3 or 0.0216 psi/ft . 3 lT 0.930 2(32.2) 0.990 From Eq. (1), for NOR PAC, with hL = 0.0269 m3/m3 and liquid Froude number = 4.28x10-4 P 0.947 - 0.0269 = 0.0149 lT 0.947 3/ 2 exp 1/ 2 13300 ( 4.28 10-4 ) = 0.0150 psi/ft 3003/ 2 From Eq. (1), for Montz, with hL = 0.0341 m3/m3 and liquid Froude number = 6.39x10-4 P 0.930 - 0.0341 = 0.0216 lT 0.930 Summary of Results: 1.5 exp 1/ 2 13300 6.39 10-4 ) = 0.0218 psi/ft 3/ 2 ( 300 Column diameter, ft: Rectifying section Stripping section Packed height, ft Rectifying section Stripping section Total Pressure drop, psi NOR PAC 4.3 3.4 6.3 20.2 26.5 0.38 Montz 5.3 4.2 2.9 2.1 5.0 0.093 From these results, the column diameter for NOR PAC would be 4.5 ft without swaging. The column diameter for Montz would be 5.5 ft without swaging. The column height for NOR PAC, allowing 4 ft for disengagement, 2 ft for feed entry, and 10 ft for sump at the bottom would be 42.5 ft or say 43 ft. Similarly, the column height for Montz would be 21 ft. Exercise 7.52 (continued) Analysis: Summary (continued) The above may be compared with the existing plate column of Exercise 7.41: Column diameter, ft Column height, ft Pressure drop, psi Plate column 6.0 36 0.8 NOR PAC 4.5 43 0.38 Montz 5.5 21 0.093 The packed columns are smaller in diameter, the NOR PAC column significantly smaller. The Montz structure packing column is shorter in height and has much less pressure drop. The NOR PAC rings random packing column is taller and has significantly more pressure drop than the column with Montz packing. The pressure drop is not critical. Because of the much higher cost for Montz packing and the much smaller diameter for the NOR PAC column, it is probable that economics would favor the NOR PAC column. Exercise 7.53 Subject: Distillation of a mixture of benzene and toluene in a packed column. Given: Conditions in Example 7.1, but using 50-mm Hiflow rings at 70% of flooding. Assumptions: Constant molar overflow. Partial reboiler is an equilibrium stage. Find: (a) Liquid holdup. (b) Column diameter. (c) HOG. (d) Packed height. (e) Pressure drop. Advantages of packed column over trayed column. Analysis: At the top of the column, L = RD =1.59(275) = 437 lbmol/h, and V = L + D = 437 + 275 = 712 lbmol/h. At the bottom of the column, for constant molar overflow and 61.1 mol% feed vapor, V = 712 - 0.611(450) = 437 lbmol / h, and L = V + B = 437 + 175 = 612 lbmol / h Before computing liquid holdup, we need to compute column diameter. (b) Column diameter based on conditions at the top of the column, where T = 180oF and P = 14.7 psia = 1 atm. Use Figs. 6.36a, b, and c to compute flooding velocity, following Example 6.13. From the ideal gas law, V = PMV/RT = (1)(78.8)/[0.730(180 + 460)] = 0.171 lb/ft3 For the liquid, which is mainly benzene, use L = 50.6 lb/ft3 and L = 0.31 cP. LM L The abscissa in Fig. 6.36a = X = VM V V L 1/ 2 437(79.4) 0.171 = 712(78.8) 50.6 1/ 2 = 0.0376 From Fig. 6.36a, Y at flooding = 0.17; from Fig. 6.36b, for water/ L = 62.4/50.6 = 1.23, f{L } = 1.5; and from Fig. 6.36c, for L = 0.31 cP, f{L } = 0.8. From Table 6.8, FP = 16 for 50-mm Hiflow rings. From a rearrangement of the ordinate of Fig. 6.36a, g H2 O (L) 1 32.2 62.4 1 2 uo = Y = 0.17 = 127 (ft/s)2 and uo = 11.2 ft/s FP V f { L } f { L } 16 0171 123(0.8) . . For fraction of flooding = f = 0.7, uV = uo f = 11.2(0.7) = 7.84 ft/s. From Eq. (6-103), column diameter is, 4VM G DT = fu0 G 1/ 2 4VM G = uG G 1/ 2 4(712 / 3600)(78.8) = 7.84(3.14)(0.171) 1/ 2 = 3.8 ft. Analysis: (b) (continued) Column diameter based on conditions at the bottom of the column, where T = 227oF and P = 14.7 psia = 1 atm. Use Figs. 6.36a, b, and c to compute flooding velocity, following Example 6.13. From the ideal gas law, V = PMV/RT = (1)(90.6)/[0.730(227 + 460)] = 0.181 lb/ft3 For the liquid, which is mainly benzene, use L = 48.7 lb/ft3 and L = 0.25 cP. LM L The abscissa in Fig. 6.36a = X = VM V V L 1/ 2 Exercise 7.53 (continued) 612(90.6) 0.181 = 437(91.5) 48.7 1/ 2 = 0.085 From Fig. 6.36a, Y at flooding = 0.12; from Fig. 6.36b, for water/ L = 62.4/48.7 = 1.28, f{L } = 1.5; and from Fig. 6.36c, for L = 0.25 cP, f{L } = 0.78. From Table 6.8, FP = 16 for 50-mm Hiflow rings. From a rearrangement of the ordinate of Fig. 6.36a, g H2 O (L) 1 32.2 62.4 1 2 uo = Y = 0.12 = 71 (ft/s)2 and uo = 8.43 ft/s . . V 16 0181 150(0.78) FP f { L } f { L } For fraction of flooding = f = 0.7, uV = uo f = 8.43(0.7) = 5.9 ft/s. From Eq. (6-103), column diameter is, 4VM V DT = fu0 V 1/ 2 4VM V = uV V 1/ 2 4(437 / 3600)(90.6) = 5.9(3.14)(0.181) 1/ 2 = 3.6 ft. (a) For liquid holdup estimates, assume the column operates in the preloading region. Therefore, the holdup is independent of the gas rate. Follow Example 6.12. Pertinent packing characteristics from Table 6.8 are a = 92.3 m2/m3 = 28.1 ft2/ft3, = 0.977, and Ch = 0.876 Use Eqs. (6-97) to (6-101), which requires calculating the liquid Reynolds and Froude numbers. Liquid holdup based on conditions at the top of the column. By the continuity equation, superficial liquid velocity is, uL = LML /LS = (437/3600)(79.9)/ [50.6(3.14)(3.8)2/4] =0.0168 ft/s 0.0168 50.6 u From Eq. (6-98), N Re L = L L = = 145 L a (0.31)(0.000672)281 . u 2 a ( 0.0168 ) 28.1 From Eq. (6-99), N FrL = L = = 2.46 10 -4 g 32.2 Since NRe > 5, use Eq. (6-101), 2 0. 0 ah / a = 0.85Ch N Re25 N Fr.1 = 0.85(0.876)(145) 0.25 (2.46 10-4 ) 0.1 = 112 . L L Exercise 7.53 (continued) Analysis: (a) top of the column (continued) From Eq. (6-97), the fractional liquid holdup is, hL = 12 N FrL N Re L 1/ 3 ah a 2/3 2.46 10-4 = (12) 145 1/ 3 (1.12 ) 2/3 = 0.0294 m3/m3 Liquid holdup based on conditions at the bottom of the column. By the continuity equation, superficial liquid velocity is, uL = LML /LS = (612/3600)(90.6)/ [50.6(3.14)(3.6)2/4] =0.0299 ft/s 0.0299 48.7 u From Eq. (6-98), N Re L = L L = = 308 L a (0.25)(0.000672)281 . u 2 a ( 0.0299 ) 28.1 = 7.80 10 -4 From Eq. (6-99), N FrL = L = 32.2 g Since NRe > 5, use Eq. (6-101), 0. 0 ah / a = 0.85Ch N Re25 N Fr.1 = 0.85(0.876)(308) 0.25 (7.80 10 -4 ) 0.1 = 152 . L L 2 From Eq. (6-97), the fractional liquid holdup is, hL = 12 N FrL N ReL 1/ 3 ah a 2/3 7.80 10-4 = (12) 308 1/ 3 (1.52 ) 2/3 = 0.0412 m3/m3 mV H L , where we must use the slope, m, of the L equilibrium curve instead of a K-value because the equilibrium curve is curved. At the top of the column: (c) From Eq. (7.51), HOG = HG + Estimate HL from Eq. (6-132), using the following properties and parameters: CL = 1.168, hL = 0.0294 m3/m3, = 0.977, and uL = 0.00512 m/s. We need an estimate of the diffusivity of benzene in toluene at high concentrations of benzene. Because these two components form a nearly ideal solution, assume that this diffusivity is equal to that of toluene (A) at infinite dilution in benzene (B), as estimated from the Eq. (3-42), using Table 3.3 and Table 3.2. In this equation, T = 180oF = 356 K, B = 205.3, A = 245.5, = 0.32 cP, and = 6(14.8)+6(3.7) -15 = 96 DL = DAB = 155 10-8 . T 1.29 (PB0.5 / PA0.5 ) 3551.29 (205.30.5 / 24550.42 ) . = 155 10-8 . = 4.28 10-5 cm2/s 0.92 0.23 0.92 0.23 B B 0.32 96 Analysis: (c) top of the column (continued) From Eq. (6-132), 1 1 HL = CL 12 = 0.24 a aPh 1/ 6 Exercise 7.53 (continued) 4hL DL au L m 1/ 2 uL a a 1 1 = aPh 1.168 12 1/ 6 4(0.0294)(0.977) (4.28 10 -9 )(92.3)(0.00512) 1/ 2 0.00512 92.3 a aPh To compute (aPh /a), we need for the liquid phase, the Reynolds, Weber, and Froude numbers from equations (6-138), (6-139), and (6-140), respectively, based on the packing hydraulic diameter from (6-137). d h = packing hydraulic diameter = 4 0.977 =4 = 0.0423 m = 0.139 ft a 92.3 Reynolds number = N ReL , h = u L d h L (0.0168)(0.139)(50.6) = = 567 [(0.31)(0.000672)] L Take the surface tension of benzene at 180oF as = 21 dynes/cm = 0.00144 lbf/ft or 0.00144(32.2) = 0.0464 lbm/s2 Weber number = N WeL ,h ( 0.0168 ) (50.6)(0.139) = 0.0428 u 2 d = L L h = 0.0464 2 2 Froude number = N FrL , h ( 0.0168) = 0.0000631 u2 = L = gd h 32.2(0.139) From (6-136), aPh -1/ 2 = 1.5 ( ad h ) N Re L ,h a = 1.5 [ (92.3)(0.0423) ] ( ) (N -0.2 -1/ 2 We L , h ) (N ) 0.75 FrL , h -0.45 ( 567 ) ( 0.0428 ) ( 0.0000631) -0.2 0.75 -0.45 = 1.56 HL = 0.24 1 = 0.154 m 1.56 Exercise 7.53 (continued) Analysis: (c) at the top of the column (continued) Now estimate the value of HG from Eq. (6-133). Estimate the gas diffusivity from Eq. (3-36), where from Table 3.1, . V = 6(15.9 + 2.31) - 18.3 = 90.96 for benzene, and V = 7(15.9 + 2.31) + 2.31 - 18.3 = 1115 for toluene and molecular weights are 787 for benzene and 92 for toluene. 0.00143T 1.75 DG = DAB = 1/ 2 1/ 3 2 P V A + (1 / M A ) + (1 / M B ) = 0.00143(355)1.75 1/ 2 1/ 3 2 V B 2 1/ 3 90.96 + 1115 . (1 / 78) + (1 / 92) Use the following additional properties and parameters, together with V = 0.0091 cP and V = u (7.84)(0.171) 0.171 lb/ft3 , where from Eq. (6-134), N ReV = V V = = 7,800 aV (28.1)(0.0091)(0.000672) (0.0091)(0.000672) From (6-135) NScV = V = = 0.639 V DV 0.052 (0.171) (2.54) 2 (12) 2 From Eq. (6-133), with CV = 0.408 from Table 6.8, (1) 1/ 3 2 = 0.052 cm2 / s 1 1/ 2 4 HG = ( - hL ) CV a4 1/ 2 (N ) (N ) -3/ 4 ReV ScV -1/ 3 uV a DV aPh -3/ 4 1 1/ 2 4(0.977) = ( 0.977 - 0.0273) 0.408 92.34 1/ 2 ( 7,800 ) ( 0.639 ) -1/ 3 (2.39)(92.3) = 0.23 m (0.052 10-4 )(144) At the top of the column, vapor rate = V =712 lbmol/h and liquid rate = L = 437 lbmol/h. Therefore, V/L = (712)/(437) = 1.63 At the bottom of the column: Near a mole fraction, xB , of 0.05, m = 2.1 Need an estimate of the liquid diffusivity of benzene in toluene at low concentrations of benzene. At the bottom, temperature = 227oF = 382 K and liquid viscosity = 0.25 cP. This time, benzene is component A and toluene is B. From Eq. (3-42), T 1.29 (PB0.5 / PA0.5 ) 3821.29 (24550.5 / 205.30.42 ) . DL = DAB = 155 10-8 . = 155 10-8 . = 6.64 10-5 0.92 0.23 0.92 0.23 B B 0.25 118.2 Analysis: (c) bottom of the column (continued) From Eq. (6-132), using SI units Exercise 7.53 (continued) 1 1 HL = CL 12 1/ 6 4hL DL au L 1/ 2 uL a a aPh 1/ 2 1 1 = 1.168 12 1/ 6 4(0.0312)(0.977) (6.64 10 -9 )(92.3)(0.00911) 0.00911 92.3 a a = 0.26 aPh aPh m To compute (aPh /a), we need for the liquid phase, the Reynolds, Weber, and Froude numbers from equations (6-138), (6-139), and (6-140), respectively, based on the packing hydraulic diameter from (6-137). d h = packing hydraulic diameter = 4 0.977 =4 = 0.0423 m = 0.139 ft a 92.3 Reynolds number = N ReL ,h = uL d h L (0.0299)(0.139)(48.7) = = 1205 [(0.25)(0.000672)] L Take the surface tension of toluene at 227oF as = 18 dynes/cm = 0.00123 lbf/ft or 0.00123(32.2) = 0.0397 lbm/s2 Weber number = N WeL ,h = 2 uL L d h ( 0.0299 ) (48.7)(0.139) = = 0.152 0.0397 2 2 Froude number = N FrL ,h ( 0.0299 ) = 0.000200 u2 = L = gd h 32.2(0.139) From (6-136), aPh -1/ 2 = 1.5 ( ad h ) N Re L ,h a = 1.5 [ (92.3)(0.0423) ] ( ) (N -0.2 -1/ 2 We L , h ) (N ) 0.75 FrL , h -0.45 (1205 ) ( 0.152 ) ( 0.000200 ) -0.2 0.75 -0.45 = 2.07 HL = 0.26 1 = 0.126 m 2.07 Analysis: (c) bottom of the column (continued) Exercise 7.53 (continued) Now estimate the value of HG from Eq. (6-133). Estimate the gas diffusivity from Eq. (3-36), by ratio from the value above for the top of the column, DV = 0.052(382/354)1.75 = 0.059 cm2/s Use the following additional properties and parameters, together with V = 0.0092 cP and V = 0.179 lb/ft3 , where from Eq. (6-134), N ReV = From (6-135),) NScV = uV V (5.9)(0.181) = = 6,150 aV (28.1)(0.0092)(0.000672) V (0.0092)(0.000672) = = 0.54 V DV 0.059 (0.181) (2.54) 2 (12) 2 From Eq. (6-133), 1 1/ 2 4 HG = ( - hL ) CV a4 1/ 2 (N ) (N ) -3/ 4 ReV ScV -1/ 3 uV a DV aPh -3/ 4 1 1/ 2 4(0.977) = ( 0.977 - 0.0412 ) 0.408 92.34 1/ 2 ( 6,150 ) ( 0.54 ) -1/ 3 (1.80)(92.3) = 0.14 m (0.059 10 -4 )(191) At the bottom of the column, vapor rate = V =437 lbmol/h and liquid rate = L = 612 lbmol/h. Therefore, V/L = (437)/(612) = 0.714 mV From Eq. (7.51), HOG = HG + HL L Therefore, values of HOG depend on the slope of the equilibrium line, which varies widely from the top of the column to the bottom. At the top, assuming HG and HL are constant, mV H L = 0.23 + 1.63(0.154) m = 0.23 + 0.25 m L At the bottom, assuming HG and HL are constant, H OG = H G + H OG = H G + mV H L = 0.14 + 0.714(0.126) m = 0.14 + 0.09 m L (d) The packed height is given by Eq. (6-127), lT = HOGNOG = HETP Nt Apply these equations in a stage-by-stage fashion, using the equilibrium stages in Fig. 7.15. For each stage, convert the value of HOG to HETP with Eq. (7.53) as in Table 7.7 for Example 7.9, HETP = HOG ln/(-1) Analysis: (d) (continued) Exercise 7.53 (continued) Using a spreadsheet, the following results are obtained: Rectifying Section: Stage m 1 0.47 2 0.53 3 0.61 4 0.67 5 0.72 6 0.80 Total height = = mV/L HOG , m 0.77 0.35 0.86 0.36 0.99 0.38 1.09 0.40 1.17 0.41 1.30 0.43 2.31 m = mV/L 0.64 0.70 0.82 1.00 1.21 1.36 1.57 1.67 m HOG , m 0.22 0.23 0.24 0.27 0.29 0.31 0.34 HETP, m 0.39 0.39 0.38 0.38 0.38 0.38 Stripping Section: Stage m 7 0.90 8 0.98 9 1.15 10 1.40 11 1.70 12 1.90 0.2 of 2.20 13 Total height = HETP, m 0.27 0.27 0.27 0.27 0.27 0.27 0.2*0.27 = 0.05 Note that although HOG varies significantly with stage in both sections, HETP does not. The total packed height = 2.31 + 1.67 = 3.98 m or 13.1 ft. (e) Compute pressure drops for the packed sections from the correlation of Billet and Schultes. The pressure drop per unit height of packed bed is given by Eq. (6-115), 1/ 2 Po - hL P 13300 = exp ( N FrL ) lT lT a 3/ 2 This equation will be used in each section. 1.5 (1) Po a u 2 1 = o 3 o G lT 2 gc KW The required packing parameters are as follows from Table 6.8 and above: FP = 16 ft2/ft3, a = 92.3 m2/m3, = 0.977 m3/m3, Ch = 0.876, and Cp = 0.421 From Eq. (6-110), (2) Analysis: (e) (continued) 1- 1 - 0.977 From Eq. (6-112), DP = 6 =6 = 0.00150 m or 0.0049 ft a 92.3 Rectifying section of the column: Column diameter = DT = 3.8 ft 1 2 1 DP 2 1 0.0049 From Eq. (6-111), = 1+ = 1+ = 1037 . KW 3 1 - DT 3 1 - 0.977 3.8 Therefore, KW = 1/1.037 = 0.964 Superficial gas velocity at 70% of flooding from above = 7.84 ft/s Gas viscosity = 0.0091 cP and gas density = 0.171 lb/ft3. From Eq. (6-114), N ReV = Exercise 7.53 (continued) uV DP V (7.84)(0.0049)(0.171) KW = (0.964) = 45, 000 (1 - ) V (1 - 0.977 ) (0.0091)(0.000672) 64 1.8 64 1.8 + 0.08 = 0.421 + = 0.322 N ReV N ReV 45, 000 45, 0000.08 From Eq. (6-113), o = C p From Eq. (2), 92.3(0.3048) 7.84 2 (0171) . 1 Po = 0.322 = 162 lbf/ft3 or 0.0113 psi/ft . 3 0.977 2(32.2) 0.977 lT From Eq. (1), with hL = 0.0294 m3/m3 and liquid Froude number = 2.46 x 10-4 P 0.977 - 0.0294 = 0.0113 0.977 lT Stripping section of the column: Column diameter = DT = 3.6 ft From Eq. (6-102), 3/ 2 exp 1/ 2 13300 ( 2.46 10-4 ) = 0.0137 psi/ft 92.33/ 2 1 2 1 DP 2 1 0.0049 = 1+ = 1+ = 1039 . KW 3 1 - DT 3 1 - 0.977 3.6 Therefore, KW = 1/1.039 = 0.962 Superficial gas velocity at 70% of flooding from above = 5.9 ft/s Gas viscosity = 0.0092 cP and gas density = 0.179 lb/ft3. Exercise 7.53 (continued) Analysis: (e) Stripping section (continued) From Eq. (6-105), N ReV = uV DP V (5.9)(0.0049)(0.179) (0.962) = 35, 000 KW = (1 - ) V (1 - 0.977 ) (0.0092)(0.000672) 64 1.8 64 1.8 + 0.08 = 0.421 + = 0.329 35, 000 35, 0000.08 N ReV N ReV From Eq. (6-104), o = C p From Eq. (2), Po 92.3(0.3048) 5.9 2 (0.179) 1 = 0.329 = 1.0 lbf/ft3 or 0.0069 psi/ft 3 0.977 2(32.2) 0.962 lT From Eq. (1), with hL = 0.0412 m3/m3 and liquid Froude number = 7.80 x 10-4 P 0.977 - 0.0412 = 0.0069 lT 0.977 3/ 2 exp 1/ 2 13300 7.80 10-4 ) = 0.0098 psi/ft 3/ 2 ( 92.3 From these results, the column diameter would be 4.0 ft without swaging. Take the packed height as 14 ft. The column height, allowing 4 ft for disengagement, 2 ft for feed entry, and 10 ft for sump at the bottom would be 30 ft. The total pressure drop is: (2.31/0.3048)(0.0069) + (1.67/0.3048)(0.0098) = 0.11 psi. For the plate column of Example 7.1, with 12.2 equilibrium stages, take tray efficiency = 80%. Therefore, need 16 trays, with a column height of 44 ft. Take a pressure drop of 0.08 psi/tray, giving a total pressure drop of 18(0.08) = 1.44 psi. The column diameter is estimated to be 4.5 ft. The following summarizes the comparison. with the plate column of Exercise 7.1: Plate column Column diameter, ft Column height, ft Pressure drop, psi 4.5 44 1.44 Packed column 4.0 30 0.11 The packed column is very competitive with the plate column. Exercise 7.54 Subject: Determining compositions and energy requirements for a mixture of nC6 and nC8 with an enthalpy-concentration diagram for 101 kPa. Given: Enthalpy-concentration diagram in Fig. 7.37. Find: (a) Composition of vapor at bubble-point for 30 mol% nC6. (b) Energy to vaporize 60 mol% of a mixture initially at 100oF with 20 mol% nC6 . (c) Vapor and liquid compositions resulting from part (b). Analysis: (a) In Fig. 7.37, below, the bubble point for 30 mol% nC6 in nC8 is shown as point B, which corresponds to a temperature of 211oF. An equilibrium tie line is drawn from point B to point M at the intersection with the saturated vapor line. The vapor composition at point M is 69 mol% nC6 and 31 mol% nC8. (b) From Fig. 7.37, reproduced below, the enthalpy of the initial mixture of 20 mol% nC6 at 100oF, shown at point G, is 6,500 Btu/lbmol = HG . A vertical line is drawn upward from point G until it reaches the point E, located on the dashed tie line between the saturated liquid and saturated vapor curves, at a point that is 60% of the way from the saturated liquid curve to the saturated vapor curve. The enthalpy of the two-phase mixture at point E is 23,000 Btu/lbmol = HE. Therefore, the energy required = Q = HE - HG = 23,000 - 6,500 = 16,500 Btu/lbmol. (c) From the reproduction of Fig. 7.37 below, the equilibrium liquid phase composition is at point D, which is 7.5 mol% nC6 , while the equilibrium vapor composition is at point F, which is 28.5 mol% nC6 . Checking this by material balance on nC6 , for a basis of 100 lbmoles of mixture: 0.20(100) = 20 lbmoles 0.075(40) + 0.285(60) = 20.1 lbmoles This is a reasonably good check. Analysis: (continued) Exercise 7.54 (continued) Exercise 7.55 Subject: Determining compositions and energy requirements for a mixture of nC6 and nC8 with an enthalpy-concentration diagram for 101 kPa. Given: Enthalpy-concentration diagram in Fig. 7.37. Find: (a) Temperature and compositions of liquid and vapor, where F1 = 950 lb/h of 30 mol% nC6 at 180oF is adiabatically mixed with F2 = 1,125 lb/h of 80 mol% nC6 at 240oF. (b) Energy and resulting phase compositions when a mixture of 60 mol% nC6 at 260oF is cooled and partially condensed to 200oF. (c) Compositions and relative amounts of the two phases when the equilibrium vapor from part (b) is further cooled to 180oF. Analysis: (a) In Fig. 7.37, reproduced on the next page, the two feeds before mixing are shown at F1 and F2. When these two feeds are adiabatically mixed, the resulting two-phase mixture contains the following, where it is necessary to convert mass to moles. Component nC6 nC8 Total MW 86.11 114.14 Feed F1 : lb 232 718 950 lbmol 2.69 6.29 8.98 Feed F2 : lb 845 280 1,125 lbmol 9.81 2.45 12.26 Total feeds: lbmol 12.5 8.74 21.24 mole fraction 0.5885 0.4115 100 Based on the mole fractions of the combined feed, point M, the adiabatic mixing point is located in Fig.7.37 on the next page. By coincidence the mixing line connecting the two feeds is also a tie line for a temperature of 204oF. The equilibrium vapor is located at point V, while the equilibrium liquid is located at point L. The equilibrium vapor has a mole fraction of 0.75 for nC6 and 0.25 for nC8. The equilibrium liquid has a mole fraction of 0.35 for nC6 and 0.65 for nC8. Although not requested, the moles of equilibrium vapor and liquid can be determined by material balances. By total material balance, F1 + F2 = 21.24 = V + L (1) By material balance on n-hexane, 12.5 = 0.75V + 0.35L (2) Solving Eqs. (1) and (2), V = 12.68 lbmol and L = 8.56 lbmol The molar % vaporized = 12.69/21.24 x 100% = 59.7%. This is consistent with the value obtained by the inverse lever arm rule on the diagram on the next page. Analysis: (a) (continued) Exercise 7.55 (continued) Exercise 7.55 (continued) Analysis: (continued) (b) From the diagram below, the enthalpy of the feed, F , of 60 mol% nC6 at 260oF is equal to HF = 27,700 Btu/lbmol. Upon cooling, the point P, in the two-phase region, is reached at 200oF. From the tie line through point P, the equilibrium vapor and liquid phases are shown at points V and L. The enthalpy of the two-phase mixture at point P = HP = 17,700 Btu/lbmol. Therefore, the energy that must be removed = Q = 27,700 - 17,700 = 10,000 Btu/lbmol mixture. The composition of the equilibrium vapor at point V is 77 mol% nC6. The composition of equilibrium liquid at point L is 38 mol% nC6 . The amounts of vapor and liquid can be obtained by material balances or by the inverse lever arm rule using the plot below. By the latter, the molar percent vapor = 53%. (c) In the diagram below, the equilibrium vapor from part (b), shown at F, is cooled to 180oF to point M in the two-phase region. This mixture splits into equilibrium vapor, V, and equilibrium liquid, L. The composition of V is 92 mol% nC6. The composition of L is 60 mol% nC6. By the lever arm rule, the mole % of equilibrium vapor = 60%. Analysis: (b) (continued) Exercise 7.55 (continued) Analysis: (c) (continued) Exercise 7.55 (continued) Exercise 7.56 Subject: Development of an enthalpy-concentration diagram for methanol-water at 1 atm. Use of diagram to obtain minimum stages. Given: Enthalpy and vapor-liquid equilibrium data in Table 7.8. Feed of 100 lbmol/h containing 60 mol% methanol. Liquid distillate to contain 98 mol% methanol. Liquid bottoms to contain 96 mol% water. Find: (a) Enthalpy-concentration diagram. (b) Minimum number of stages. (c) Temperatures of distillate and bottoms. Analysis: (a) Use a spreadsheet to plot an enthalpy-concentration diagram, using the supplied data. The result is shown below, where the dashed lines are tie lines for the vapor-liquid equilibrium compositions at the temperatures indicated and a pressure of 1 atm. Exercise 7.56 (continued) Analysis: (continued) (b) and (c) For the condition of total reflux, passing vapor and liquid streams have the same compositions and are saturated vapor and liquid, respectively. Therefore, while vapor and liquid streams at equilibrium are related by tie lines extending from the saturated liquid curve to the saturated vapor curve, passing vapor and liquid streams are related by vertical operating lines extending from the saturated liquid curve to the saturated vapor curve. By alternating between the equilibrium and operating lines, from the bottoms composition to the distillate composition, the minimum number of stages is determined by the number of tie lines used, as shown in the diagram below for the example of a methanol bottoms mole fraction of 0.04 and a methanol distillate mole fraction of 0.98, where the construction lines are solid lines and the equilibrium lines are numbered from the top. The minimum number of stages is seen to be 6. The value is independent of the feed condition because at total reflux, there is no feed. The temperature of the bottoms is 93.5oC. The temperature of the distillate is 64.7oC Exercise 8.1 Subject: Extraction vs. Distillation. Given: Dilute mixture of benzoic acid in water. Find: Why liquid-liquid extraction is the preferred separation method. Analysis: At 1 atm, the boiling point of benzoic acid is 249.2OC. Thus, if a dilute solution of benzoic acid and water is distilled, a large amount of water, with its very high heat of vaporization (about 1,000 Btu/lb) must be vaporized, even though the relative volatility, , is large and the separation is easy. L/L extraction can be used with a solvent like methyl isobutyl ketone, MIBK, which has a normal boiling point of 119oC and can be easily recovered from the benzoic acid by distillation and then recycled [see Trans. AIChE, 42, 331 (1946). Exercise 8.2 Subject: Liquid-Liquid Extraction vs. Distillation. Given: Dilute mixture of formic acid in water. Find: Why liquid-liquid extraction is the preferred separation method. Analysis: At 1 atm, the boiling point of formic acid is 100.6oC, almost identical to the normal boiling point of water. A maximum-boiling azeotrope is formed at 107.1oC with 43.3 mol% water. Thus, it is impossible to separate formic acid from water at 1 atm by ordinary distillation. L/L extraction can be used with chloroform as the solvent to extract formic acid [see J. Chem. Eng. Data, 5, 301 (1960) and J. Chem. Eng. Data, 4, 42 (1959)]. Tetrahydrofuran can also be used as the solvent. Exercise 8.3 Subject: Selection of extraction device. Given: Liquid-liquid extraction of acetic acid from water by ethyl acetate. Fig. 8.8 for selecting extraction equipment. Find: Adequacy of the RDC. Other types of adequate extractors. Analysis: In the process of Fig. 8.1 for the L/L extraction of acetic acid from water by ethyl acetate, 6 equilibrium stages are needed. Pertinent properties are: < 1 cP, = 0.08 g/cm3, I > 30 dyne/cm Need stainless steel equipment. Assume no emulsion formation. Have large flow rates. Assume water-rich phase has a density of 1,000 kg/m3. Assume ethyl acetate-rich phase has a density of 900 kg/m3. Feed rate = 30,300 lb/h or 13,700 kg/h Solvent rate = 71,100 lb/h or 32,300 kg/h Total volumetric throughput = 13,700/1,000 + 32,300/900 = 49.6 m3/h In Fig. 8.8, this throughput is on the border line. Use a mixer-settler battery or any kind of a mechanically assisted column. Exercise 8.4 Subject: Extraction devices. Given: ARD and RDC. Find: Advantages and disadvantages of each device. RDC: A great deal of experience exists. Hundreds of units have been installed. Can easily remove agitator assembly shaft. Easily fabricated and assembled. ARD: Get reduced back-mixing. Can not easily remove agitator assembly shaft. More complex unit to fabricate and assemble. Analysis: Exercise 8.5 Subject: Selection of extraction devices. Given: Cascade of mixer-settler units. Find: Conditions for selection of this device. Analysis: Mixer-settler units are preferred for: 1. Less than 5 equilibrium stages. 2. Low head room, if much floor space is available. 3. Easy scale-up and certain performance. 4. Wide ratio of feed-to-solvent flow rates. Exercise 8.6 Subject: Selection of extraction device Given: 4,000 bbl/day of petroleum reformate. 5 volumes of diethyleneglycol to extract aromatics from the paraffins in one volume of reformate. 8 theoretical stages. Find: Preferred type of extractor using Tables 8.2 and 8.3, with Fig. 8.8. Analysis: Total throughput is 4,000 + 5(4,000) = 24,000 bbl/day. Assume 42 gal/bbl, which is equivalent to 0.159 m3/bbl. Then throughput = 24,000(0.159)/24 = 159 m3/h From Fig. 8.8, consider RDC, ARD, Kuhni, or Lurgi extractors. Exercise 8.7 Subject: Selection of extraction solvents. Given: The following mixtures: (a) water - ethyl alcohol (b) water - aniline (c) water - acetic acid Find: A possible solvent for each mixture and the resulting solute. Analysis: Use Table 8.4, where the minus sign designates a suitable solvent group. (a) water - ethyl alcohol: Solute Solute group Desired solvent group Possible solvent water Group 2 Group 1 no solvent to extract water ethyl alcohol Group 2 Group 1 n-butanol Solute water aniline (b) water - aniline: Solute group Group 2 Assume Group 5 with amine group controlling Desired solvent group Group 1 Groups 1 and 6 Possible solvent no solvent to extract water trichloroethane (c) water - acetic acid: Solute Solute group Desired solvent group water Group 2 Group 1 acetic acid Group 1 Groups 2,3.4.5 Possible solvent no solvent to extract water 2 - 1-butanol 3 - diisobutyl ketone 4 - methyl acetate 5 - isopropyl ether Note that it is very difficult to find a solvent to extract water from organic compounds. For that separation, adsorption may be preferred. Also, distillation is possible if the water is the more volatile component. Exercise 8.8 Subject: Selection of extraction solvents. Given: The following mixtures with solute first: (a) acetone - ethylene glycol (b) toluene - n-heptane (c) ethyl alcohol - glycerine Find: Suitable solvents Analysis: Use Table 8.4, where the minus sign designates a suitable solvent group. (a) acetone - ethylene glycol: Component Group Desired solvent group Possible solvent acetone (solute) 3 1 and 6 trichloroethane ethylene glycol 2 1 (b) toluene - n-heptane: Component Group Desired solvent group toluene (solute) 7 maybe 3,5,6 n-heptane 8 none (c) ethyl alcohol - glycerine: Component Group Desired solvent group ethyl alcohol 2 1 (solute) glycerine 1 Possible solvent aniline Possible solvent Table 8.4 does not give a choice. From Section 15 of Perry's Handbook, possible solvents are benzene and carbon tetrachloride. Exercise 8.9 Subject: Characteristics of the extraction of acetic acid (A) from a dilute solution in water (C), with ethyl acetate (S) at 25oC. Given: Values or estimates of values of (KA)D , (KC)D , (KS)D , and AC. Find: If the system exhibits (a) high selectivity, (b) high solvent capacity, and (c) ease of recovery of solvent. A better solvent. Analysis: Use CHEMCAD with UNIFAC for LLE to obtain typical mass fraction compositions for two liquid phases in equilibrium at 25oC and 1 atm. The LLVF three-phase flash model gives the following results: Mass fractions: Component Feed Extract Raffinate acetic acid (A) 0.0588 0.0595 0.0564 water (C) 0.2353 0.0332 0.9242 ethyl acetate (S) 0.7059 0.9073 0.0194 Define KD = mass fraction in extract phase/mass fraction in raffinate phase. (KA)D = 0.0595/0.0564 = 1.055 (KC)D = 0.0332/0.9242 = 0.036 (KS)D = 0.9073/0.0194 = 46.8 AC = (KA)D /(KC)D = 1.055/0.036 = 29.3 (a) The value of AC is high, indicating a high selectivity. (b) The value of (KA)D is not high, indicating a solvent capacity that is not high. (c) Since (KS)D is high and (KC)D is low, recovery of solvent is relatively easy. A better solvent should be sought because ethyl acetate does not have a high capacity. From the results of Exercise 8.7, four other solvents are selected. Again using CHEMCAD, the following results are obtained for distribution coefficients and relative selectivity: Solvent (KS)D (KC)D (KA)D AC ethyl acetate 29.3 1.055 46.8 0.036 1-butanol 4.43 1.187 8.76 0.268 diisobutyl ketone 94.8 0.365 1710 0.00385 methyl acetate 12.8 1.470 6.50 0.115 diisopropyl ether 102.7 0.453 516 0.00441 No solvent meets all four criteria. Ethyl acetate may be the compromise. No solvent has a high capacity. Exercise 8.10 Subject: Estimation of interfacial tension. Given: Composition of a ternary mixture. Find: Method for estimating the interfacial tension from the values of surface tension in air for each component in the mixture. Analysis: Antonoff's rule can be used to compute the interfacial tension for two liquid phases, I and II, in equilibrium, i = ( II )in air - ( I )in air This equation, which is discussed in, 1. J. Russ. Phys. Chem. Soc., 39, 342 (1907). 2. Hart, Duga, Res./Devel., 16 (9), 46 (1965). requires values of the multicomponent surface tensions of each phase [see Reid, Prausnitz, and Poling, "The Properties of Liquids and Gases", 4th ed. (1986)]. Exercise 8.11 Subject: Extraction of acetone (A) from water (C) by 1,1,2-trichloroethane (S) at 25oC Given: 1,000 kg/h of 45 wt% A in C. Liquid-liquid equilibrium data. 10 wt% A in raffinate. Find: Using Hunter-Nash method with an equilateral triangle diagram. (a) Minimum flow rate of S. (b) Number of equilibrium stages for solvent rate of 1.5 times minimum. (c) Flow rate and composition of each stream leaving each stage. Analysis: Using the given equilibrium data in weight fractions, the triangular diagram is shown on the next page, where a solid line is used for the equilibrium curve and dashed lines are used for the tie lines, only three of which are given. Additional tie-line locations can be made using either of the two techniques illustrated in Fig. 8.16. (a) The minimum solvent flow rate corresponds to an infinite number of equilibrium stages. To determine this minimum: (1) Points are plotted on the triangular diagram for the given compositions of the solvent, S (pure S), feed, F (45 wt% A, 55 wt% C), and raffinate, RN (10 wt% A on the equilibrium curve), followed by drawing an operating line through the points S and RN and extending it only to the right because the tie lines slope down from left to right; (2) Because the pinch region, where equilibrium stages crowd together, may occur anywhere, additional potential operating lines are drawn through the tie lines extended to the right until they cross the operating line at the solvent end, drawn in (1), where these intersections for four tie lines (the three given and one other that extends through the feed point) are denoted P1, P2, P3, and P4; (3) The intersection point farthest from the triangular diagram is the one used to determine the minimum solvent rate, where here it is P1 = Pmin , corresponding to a pinch point at the feed end of the cascade; (4) An operating line is drawn from P1 through F to the determine the extract E1, followed by drawing lines from E! to RN and from S to F, with the intersection of these two lines determining the mixing point, M. The wt% A at the mixing point is 36.5%. From Eq. (8-10), xA F - xA M 0.45 - 0.365 S min = = = 0.233 F xA M - xA S 0.365 - 0.0 Therefore, Smin = 0.233F = 0.233(1,000) = 233 kg/h (b) Solvent rate = 1.5 Smin = 1.5(233) = 350 kg/h With this rate, the total (feed + solvent) component flow rates becomes: A = 450 kg/h, C = 550 kg/h, and S = 350 kg/h, giving a total of 1,350 kg/h. Thus, the mixing point is at (xA)M = 450/1,350 = 0.333, (xC)M = 550/1,350 = 0.408, and (xS)M = 350/1,350 = 0.259. This point is plotted as M on the second triangular diagram on a following page. A line from RN through M to an intersection with the equilibrium curve determines the extract point E1, as illustrated in Fig. 8.14. The operating point P is now determined, as in Fig. 8.17, by finding the intersection of lines drawn through S and RN, and through F and E1. The number of equilibrium stages is now stepped off, as in Fig. 8.17, giving a value of Nt slightly greater than 5. Exercise 8.11 (continued) Analysis: (continued) (c) First compute the two product flow rates, using the following product compositions read from the plot: Mass fractions: Component Acetone Water Feed 0.450 0.550 Solvent 0.000 0.000 Raffinate 0.100 0.895 Extract 0.509 0.040 1,1,2 TCE 0.000 1.000 0.005 0.451 By overall total material balance, F + S = 1,000 + 350 = 1,350 = RN + E1 1) By overall acetone balance, 0.45F = 450 = 0.100RN + 0.509E1 (2) Solving Eqs. (1) and (2) simultaneously, E1 = 770 kg/h and RN = 580 kg/h The composition of each stream leaving each stage can be read from the right-triangle plot prepared in part (b). Let x be mass fraction in the raffinate and y be mass fraction in the extract. The stages are arbitrarily numbered from the feed end. The flow rate of each stream leaving each stage is best obtained by total material balances around groups of stages and the inverse lever arm rule, using the operating lines as. A total balance for the first n-1 stages is: or F + En= 1,000 + En = Rn-1 + E1 = Rn-1 + 770 Rn-1 - En = 1,000 - 770 = 230 (3) By the inverse lever arm rule with the right-triangle diagram, Rn -1 E P = n En Rn -1 P (4) where the lines are measured from the right-triangle diagram. Eqs. (3) and (4) are solved simultaneously for Rn-1 and En for n = 2 to 6. The following results are obtained: Raffinate: Stage xA xC 1 0.390 0.581 2 0.325 0.651 3 0.260 0.724 4 0.190 0.800 5 0.100 0.895 xS 0.029 0.024 0.016 0.010 0.005 Extract: yA 0.509 0.440 0.360 0.270 0.165 yC 0.040 0.033 0.024 0.017 0.011 kg/h: yS R 0.451 935 0.537 822 0.616 738 0.713 656 0.824 580 E 770 706 592 508 426 Exercise 8.12 Subject: Extraction of acetone (A) from water (C) by 1,1,2-trichloroethane (S) at 25oC Given: 1,000 kg/h of 45 wt% A in C. Liquid-liquid equilibrium data. 10 wt% A in raffinate. Find: Using a right triangle diagram. (a) Minimum flow rate of S. (b) Number of equilibrium stages for solvent rate of 1.5 times minimum. (c) Flow rate and composition of each stream leaving each stage. Analysis: Using the given equilibrium data in weight fractions, the right-triangle diagram is shown on the next page, where a solid line is used for the equilibrium curve and dashed lines are used for the tie lines, only three of which are given. Additional tie-line locations can be made using either of the two techniques illustrated in Fig. 8.16. (a) The minimum solvent flow rate corresponds to an infinite number of equilibrium stages. To determine this minimum: (1) Points are plotted on the triangular diagram for the given compositions of the solvent, S (pure S), feed, F (45 wt% A, 55 wt% C), and raffinate, RN (10 wt% A on the equilibrium curve), followed by drawing an operating line through the points S and RN and extending it only to the right because the tie lines slope down from left to right; (2) Because the pinch region, where equilibrium stages crowd together, may occur anywhere, additional potential operating lines are drawn through the tie lines extended to the right until they cross the operating line at the solvent end, drawn in (1), where these intersections for four tie lines (the three given and one other that extends through the feed point) are denoted P1, P2, P3, and P4; (3) The intersection point farthest from the triangular diagram is the one used to determine the minimum solvent rate, where here it is P1 = Pmin , corresponding to a pinch point at the feed end of the cascade; (4) An operating line is drawn from P1 through F to the determine the extract E1, followed by drawing lines from E! to RN and from S to F, with the intersection of these two lines determining the mixing point, M. The wt% A at the mixing point is 36.5%. From Eq. (8-10), xA F - xA M 0.45 - 0.365 S min = = = 0.233 F xA M - xA S 0.365 - 0.0 Therefore, Smin = 0.233F = 0.233(1,000) = 233 kg/h (b) Solvent rate = 1.5 Smin = 1.5(233) = 350 kg/h With this rate, the total (feed + solvent) component flow rates becomes: A = 450 kg/h, C = 550 kg/h, and S = 350 kg/h, giving a total of 1,350 kg/h. Thus, the mixing point is at (xA)M = 450/1,350 = 0.333, (xC)M = 550/1,350 = 0.408, and (xS)M = 350/1,350 = 0.259. This point is plotted as M on the second triangular diagram on a following page. A line from RN through M to an intersection with the equilibrium curve determines the extract point E1, as illustrated in Fig. 8.14. The operating point P is now determined, as in Fig. 8.17, by finding the intersection of lines drawn through S and RN, and through F and E1. The number of equilibrium stages is now stepped off, as in Fig. 8.17, giving a value of just slightly greater than Nt = 5. Exercise 8.12 (continued) Analysis: (continued) (c) First compute the two product flow rates, using the following product compositions read from the plot: Component Feed Solvent Raffinate Extract Mass fractions: Acetone Water 0.450 0.550 0.000 0.000 0.100 0.895 0.509 0.040 1,1,2 TCE 0.000 1.000 0.005 0.451 1) By overall total material balance, F + S = 1,000 + 350 = 1,350 = RN + E1 By overall acetone balance, 0.45F = 450 = 0.100RN + 0.509E1 (2) Solving Eqs. (1) and (2) simultaneously, E1 = 770 kg/h and RN = 580 kg/h The composition of each stream leaving each stage can be read from the right-triangle plot prepared in part (b). Let x be mass fraction in the raffinate and y be mass fraction in the extract. The stages are arbitrarily numbered from the feed end. The flow rate of each stream leaving each stage is best obtained by total material balances around groups of stages and the inverse lever arm rule, using the operating lines. A total balance for the first n-1 stages is: or F + En= 1,000 + En = Rn-1 + E1 = Rn-1 + 770 Rn-1 - En = 1,000 - 770 = 230 (3) By the inverse lever arm rule with the right-triangle diagram, Rn -1 E P = n En Rn -1 P (4) where the lines are measured from the right-triangle diagram. Eqs. (3) and (4) are solved simultaneously for Rn-1 and En for n = 2 to 6. The following results are obtained: Stage 1 2 3 4 5 Raffinate: xA xC 0.390 0.581 0.325 0.651 0.260 0.724 0.190 0.800 0.100 0.895 xS 0.029 0.024 0.016 0.010 0.005 Extract: yA 0.509 0.440 0.360 0.270 0.165 yC 0.040 0.033 0.024 0.017 0.011 yS 0.451 0.537 0.616 0.713 0.824 kg/h: R 935 822 738 656 580 E 770 706 592 508 426 Exercise 8.13 Subject: Extraction of isopropanol (A) from diisopropylether (C) by water (S) at 25oC. Given: Feed containing 45 wt% A, 50 wt% C, and 5 wt% S. Raffinate to contain 2.5 wt% A. Extract to contain 20 wt% A. Liquid-liquid equilibrium data. Find: Using Varteressian-Fenske method with a McCabe-Thiele diagram, find the number of theoretical stages. Can an extract of 25 wt% A be obtained? Analysis: Because the wt% of the solute, A, in both the raffinate and extract are given, the solvent rate must be determined. Do this first by constructing a right-triangle diagram from the equilibrium tie line and phase boundary data. As shown on the next page, for a diagram that does not show the tie lines, points on the diagram are placed for: the feed, F, (45 wt% A, 50 wt% C, and 5 wt% S); the solvent, S, (100% S); the raffinate, R, (2.5 wt% A on the phase boundary line); and the extract, E, (20 wt% A on the phase boundary line). The mixing point, M, is determined by the intersection of lines drawn from: (1) F to S and (2) R to E. The ratio of solvent to feed is given by the inverse lever arm rule as the ratio of line length MF to line length MS, which equals 1.67. Take a basis of 100 kg/h of feed. Then the solvent rate is 1.67(100) = 167 kg/h. The flow rates of raffinate and extract are obtained from an overall total material balance and an overall material balance for A: F + S = 100 + 167 = 267 = R + E xA F (1) R F + xA S S = 0.45(100) + 0 = 45 = xA R + xA E E = 0.025R + 0.20 E (2) Solving Eqs. (1) and (2) simultaneously, R = 48 kg/h and E = 219 kg/h. The resulting overall material balance is as follows, where raffinate and extract compositions are read from the phase boundary curve. Component Isopropyl alcohol (A) Diisopropyl ether (C) Water (S) Total Feed 45.0 50.0 5.0 100.0 Solvent 0.0 0.0 167.0 167.0 Raffinate 1.20 46.28 0.52 48.00 Extract 43.80 3.72 171.48 219.00 To determine the number of equilibrium stages for the above solvent rate and material balance, using the Varteressian-Fenske method with a McCabe-Thiele diagram, an equilibrium curve and an operating curve are needed in a McCabe-Thiele plot of mass fraction of alcohol in the extract against the mass fraction of alcohol in the raffinate, as illustrated for a different system in Fig. 8.25. The equilibrium curve is determined from the equilibrium tie-line data, shown in a right triangle diagram below and obtained directly from the statement of the exercise. These equilibria data are tabulated below. The operating curve is determined by first determining the operating point, P, as illustrated in Fig. 8.15, followed by drawing arbitrary straight lines through that point Analysis: (continued) Exercise 8.13 (continued) Exercise 8.13 (continued) Analysis: (continued) to give corresponding extract and raffinate compositions from intersections with the phase boundary curve, as illustrated in Fig. 8.21c. Several such operating lines are shown on the righttriangle diagram below. The operating points used to draw the operating curve on the McCabeThiele are tabulated below. The points for the equilibrium curve are: Mass fraction alcohol in Extract 0.081 0.102 0.117 0.175 0.217 0.268 Mass fraction alcohol in Raffinate 0.024 0.050 0.093 0.249 0.380 0.452 Analysis: (continued) Exercise 8.13 (continued) The determined points for the operating curve are: Mass fraction alcohol in Extract 0.000 0.009 0.044 0.095 0.150 0.200 Mass fraction alcohol in Raffinate 0.025 0.050 0.150 0.250 0.350 0.450 These points are plotted in the McCabe-Thiele diagram below, where the stages are stepped off as illustrated in Fig. 8.25. The result is 2.5 equilibrium stages. Analysis: (continued) To determine if an extract of 25 wt% A can be obtained, as shown in the diagram below, the feed and desired extract points are plotted as F and E, respectively, with a straight operating line drawn between the two points. Since that line appears to be coincident with an equilibrium tie line, the extract could be obtained, but only with an infinite number of stages. An extract with an alcohol content of greater than 25 wt% A is impossible because the operating line would have a slope less than an overlapping tie line, making it impossible to step off stages in the right direction. Exercise 8.13 (continued) Exercise 8.14 Subject: Extraction of trimethylamine (TMA) from benzene (C) with water (S). Given: Three equilibrium stages. Solvent-free extract to contain 70 wt% TMA. Solvent-free raffinate to contain 3 wt% TMA. Liquid-liquid equilibrium data. Find: Using the Hunter-Nash method with a right-triangle diagram, find feed composition and water-to-feed ratio. Analysis: The given phase boundary liquid compositions and the equilibrium liquid-liquid compositions (as dashed tie lines) are plotted on the right-triangle diagram below. Included on the diagram are the final extract and raffinate compositions. The final extract composition is obtained by locating a point, P, for 70 wt% TMA, 30 wt% benzene, and 0% water (given solventfree composition), and drawing a straight line from P toward the point S (pure water) to where the line intersects the phase boundary. This is point E for the extract. In a similar manner, the raffinate composition, R, is determined. Exercise 8.14 (continued) On a second diagram, shown below, a trial and error procedure is used to find the operating point, P' , that will result in the stepping off of three equilibrium stages, as illustrated in Fig. 8.19, to obtain the specified final extract and final raffinate. The final trial is shown on the diagram, where M is the mixing point for extract + raffinate, and for feed + solvent. Assuming water-free feed, the resulting feed composition is at F, with 57.5 wt% TMA and 42.5 wt% benzene. The ratio of mass of solvent to mass of feed is given by the ratio of line lengths = line FM/line MS = 0.56. Exercise 8.15 Subject: Extraction of diphenylhexane (DPH) from docosane (C) with furfural (U) at 45 and 80oC Given: Feed, F, of 500 kg/h of 40 wt% DPH in C. 500 kg/h of solvent, S, containing 98 wt% U and 2 wt% DPH. Raffinate to contain 5 wt% DPH. Liquid-liquid equilibrium data. Find: Number of theoretical stages. DPH in kg/h in the extract. Analysis: Case of 45oC: The given liquid-liquid equilibrium data are plotted in the right-triangle diagram below. Included on the diagram are composition points F for the feed, R for the raffinate on the equilibrium curve, and S for the solvent. A straight line extends from point F to point S. Because the mass flow rates of the feed and solvent are equal, the mixing point, M, is located at the midpoint of this line. Another straight line extends from point R to point M, and then to an intersection with the equilibrium curve at point E, which is the final extract. Using the inverse lever-arm rule on line RME, the mass ratio of R to E is 0.445. Combining this with an overall material balance: F + S = 500 + 500 = 1,000 = R +E gives R = 308 kg/h and E = 692 kg/h. From the diagram, the mass fraction of DPH in the extract is 0.281. Therefore, the DPH in the extract is 0.281(692) = 194.5 kg/h, which is 92.6% of the total DPH entering the extractor. On the following page, the equilibrium stages are stepped on another right triangle diagram, as in Fig. 8.17, by determining the operating point P from extensions of lines drawn through points F and E, and S and R, followed by alternating between operating lines and equilibrium tie lines. The result is 5 equilibrium stages. Case of 80oC: The given liquid-liquid equilibrium data are plotted in the right-triangle diagram below. Included on the diagram are composition points F for the feed, R for the raffinate on the equilibrium curve, and S for the solvent. A straight line extends from point F to point S. Because the mass flow rates of the feed and solvent are equal, the mixing point, M, is located at the midpoint of this line. Another straight line extends from point R to point M, and then to an intersection with the equilibrium curve at point E, which is the final extract. Using the inverse lever-arm rule on line RME, the mass ratio of R to E is 0.383. Combining this with an overall material balance: F + S = 500 + 500 = 1,000 = R +E gives R = 277 kg/h and E = 723 kg/h. From the diagram, the mass fraction of DPH in the extract is 0.271. Therefore, the DPH in the extract is 0.271(723) = 195.9 kg/h, which is 93.3% of the total DPH entering the extractor. On the following page, the equilibrium stages are stepped on another right triangle diagram, as in Fig. 8.17, by determining the operating point P from extensions of lines drawn through points F and E, and S and R, followed by alternating between operating lines and equilibrium tie lines. The result is 4+ equilibrium stages. Analysis: Exercise 8.15 (continued) Analysis: (case of 80 C continued) o Exercise 8.15 (continued) Exercise 8.16 Subject: Selection of extraction method. Given: Four ternary systems in Fig. 8.43. Diagrams 1, 2, 4 are Type I. Diagram 3 is Type II. Find: Method for most economical process for each system. Methods are: (a) Countercurrent extraction (CE). (b) CE with extract reflux (ER). (c) CE with raffinate reflux (RR). (d) CE with ER and RR. Analysis: Note that y1 is the composition of the extract. Raffinate reflux (RR) is of little value, so don't use it. For Type I diagrams, extract reflux is rarely useful. All three Type I diagrams exhibit solutropy, making it almost impossible to obtain a good separation. The Type II diagram uses a poor solvent, making the use of extract reflux questionable. Summary: Use CE for all four systems. However, better solvents should be sought for all four systems. Exercise 8.17 Subject: Stage requirements for extraction of acetone (A) from two feeds, of different composition of acetone (A) and water (C), with a solvent of 1,1,2 - trichloroethane (S). Given: Feed F of 7,500 kg/h containing 50 wt% A in C. Second feed F' of 7,500 kg/h containing 25 wt% A in C. Solvent of 5,000 kg/h of S. Raffinate to contain 10 wt% A. Liquidliquid equilibrium data from Exercise 8.11. Find: Number of equilibrium stages and feed locations, using a right-triangle diagram. Analysis: Because the flow rates and compositions of both feeds and the solvent are given, the mixing point, M, can be computed, with the following results: kg/h: Component Feed, F Feed, F' Solvent, S Mixing point, M Acetone 3,750 1,875 0 5,625 Water 3,750 5,625 0 9,375 Trichloroethane 0 0 5,000 5,000 Total: 7,500 7,500 5,000 20,000 Thus, the composition of the mixing point is 28.125 wt% A, 46.875 wt% C, and 25.0 wt% S. This is shown below on the right-triangle equilibrium diagram. Also shown is the composition of the raffinate, R, on the equilibrium curve and corresponding to 10 wt% A. A straight line extending from R, through M, to its intersection with the equilibrium curve, establishes the extract composition at point E, as shown in the diagram below. The compositions of the raffinate and extract are thus determined from the diagram to be: Composition Acetone Water Trichloroethane Total: Raffinate, wt% 10.0 89.5 0.5 100.0 Extract, wt% 46.5 3.7 49.8 100.0 By component material balances or the inverse lever-arm rule, using the mixing point, the raffinate rate is 10,060 kg/h, the extract rate is 9,940 kg/h. Because the two feeds differ in composition, they should not be mixed before entering the cascade, but should enter separately at different stages. Note that feed F' has a composition closer to the final raffinate, leaving from stage N, than feed F. Therefore, feed F' should be fed to an intermediate stage, m, and feed F should be fed to stage 1, from which the final extract leaves. Feed F' should enter at that stage, m, where it matches most closely the intermediate raffinate entering stage m. In the diagram on the next page, two operating points are shown, P1 and P2. Analysis: (continued) Exercise 8.17 (continued) Point P1 is used to step off stages from m to N, based on a material balance for the entire cascade, because both F and F' move in the direction toward the final raffinate. Thus, the material balance is: F + F' + S = E1 + RN = M Rearranging for passing streams, F + F' - E1 = RN - S = P1 (1) Straight lines drawn through (F + F1) and E1, and through RN and S, intersect at the operating point P1. For stages 1 through m-1, the material balance is: F + Em = Rm-1 +E1 Rearranging for passing streams, F - E1 = Rm-1 - Em = P2 (2) Substituting Eq. (2) into Eq. (1), P2 + F' = RN - S = P1 (3) Exercise 8.17 (continued) Analysis: (continued) Therefore, operating point P2 is obtained by the intersection of straight lines drawn through F and E1, and through P1 and F'. The constructions to obtain P1 and P2 are shown on the next page. Note that for this system, P1 and P2 are on opposite sides of the triangular diagram. The equilibrium stages are stepped off starting with P2 from stage 1 at the end where the final extract, E1, is produced. By a tie line, E1 is in equilibrium with R1. An operating line connecting R1 with P2 gives E2. The stepping off, using P2, is continued until R2 is reached, which is very close to the composition of F'. Therefore, F' is fed to stage 2. The remaining stages are stepped off using P1 until final raffinate RN is reached. The total number of stages is almost 5. Thus, feed F is fed to stage 1 and feed F' is fed to stage 2 of a cascade with almost 5 stages. Analysis: (continued) Exercise 8.17 (continued) Exercise 8.18 Subject: Extraction of trimethylamine (T) from benzene (B) with water (W) in a threeequilibrium stage liquid-liquid extractor shown in Fig. 8.44. Given: Feed of 10,000 kg/h of 40 wt% B and 60 wt% T. Fresh solvent is sent to each of the stages. The solvent flow rate to stage 3, call it S3, is 5,185 kg/h. In Fig. 8.44, V1 is to contain 76 wt% T on a solvent-free basis, and L3 is to contain 3 wt% T on a solvent-free basis. Equilibrium data are given in Exercise 8.14. Find: Using an equilateral-triangle diagram, determine the solvent flow rates to stages 1 and 2. Analysis: The triangular diagram is shown on a following page. Since the extract, V1, is given on a solvent-free basis, its complete composition is obtained by the intersection with the equilibrium curve of a line drawn from the point (76 wt% T and 24 wt% B) to the 100% W point. Similarly, the composition of the raffinate, L3, is obtained by the intersection with the equilibrium curve of a line drawn from the point (3 wt% T and 97 wt% B) to the 100% W point. The resulting compositions of the final raffinate, L3, and the final extract, V1, are as follows: Component T B W Raffinate, L3, wt% 3.0 97.0 0.0 100.0 Extract, V1, wt% 39.5 12.5 48.0 100.0 Total Using these compositions, determine the flow rates of final raffinate and final extract by solving material balances for T and B: T balance: 0.03 L3 + 0.395 V1 = 6,000 B balance: 0.97 L3 + 0.125 V1 = 4,000 Solving, L3 = 2,188 kg/h and V1 = 15,020 kg/h By an overall total material balance, the total solvent water rate is: S3 + S1 + S2 = L3 + V1 - F = 2,188 + 15,020 - 10,000 = 7,208 kg/h where S3 = Solvent in Fig. 8.44, which enters from the right into stage 3. Because S3 is given as 5,185 kg/h, S1 + S2 = 7,208 - 5,185 = 2,023 kg/h From the above results, the overall mass balance is as follows in flow rates in kg/h: Component T B W Total Feed 6,000 4,000 0 10,000 Solvent, S3 0 0 5,185 5,185 L3 66 2,122 0 2,188 V1 5,934 1,878 7,208 15,020 S1 + S2 0 0 2,023 2,023 Exercise 8.18 (continued) Analysis: (continued) The next step is to determine the separate values for S1 and S2. Because fresh solvent enters all three stages, instead of one operating point, P, there will be three operating points, designated here as P1, P2, and P3. Referring to Fig. 8.44, let the operating point to the right of stage 3 be P3. Therefore, from (8-5), P3 = S3 - L3 = 5,185 - 2,188 = 2,997 kg/h This operating point is located along a line from L3 through solvent, S3 to the right of the triangular diagram, as shown on the next page. This operating point is also common to the passing streams, V3 and L2 to the left of Stage 3 in Fig. 8.44. That is, P3 = V3 L2 = 2,997 kg/h. Using the following compositions from Fig. 8.44 and material balances, the following values are obtained for V3 and L2: Component T B W Total Wt % in V3 4.05 0.35 95.60 100.00 Flow in V3, kg/h 219 19 5,186 5,424 Wt% in L2 11.75 88.22 0.03 100.00 Flow in L2, kg/h 285 2,141 1 2,427 The location of P3 is obtained by measurement, using a ratio, similar to (8-6), L3 P3 L3 +P3 S3 5,185 L S 2.37 - 1.00 = = = = 2.37 or 3 = = 0.578 L3 L3 2,188 2.37 SP3 L3 P3 Using this ratio, with the measured length of L3 S , the point P3 is located as shown on the triangular plot below. In Fig. 8.44, the point V3 is in equilibrium with L3 and, therefore, is located at the end of a tie line drawn from L3. V3 operates with L2 , where L2 is the intersection on the equilibrium curve of an extended line through V3 and, again, P3. From the triangular diagram, the composition of V2 is obtained from a tie line with L2, and the composition of L1 is obtained from a tie line with V1. The results are as follows, including the results of material balances around stage 1, with flows in kg/h Component T B W Total L1 , wt% 32 59 9 100 Flow in L1 , kg/h 1,199 2,210 337 3,746 V2 , wt% 16.7 1.3 82.0 100.0 Flow in V2 , kg/h 1,133 88 5,562 6,783 Flow in S1 , kg/h 0 0 1,983 1,983 Since S1 = 1,983 kg/h and, from above, S1 + S2 = 2,023, S2 = 2,023 1,983 = 40 kg/h. Exercise 8.19 Subject: Analysis of a multiple-feed countercurrent extraction cascade shown in Fig. 8.45. Given: solute. Feed F' contains solvent and solute. Feed F " contains unextracted raffinate and Find: Equations to establish the three reference (operating) points on a right-triangle graph. Analysis: Referring to Fig. 8.45, let: P' be the operating point from the left side of stage 1 to between stages m-1 and m. P" be the operating point from between stages m and m+1 to between stages p-1 and p. P''' be the operating point from between stages p and p+1 to the right side of stage n. The equations for locating these three operating points are obtained by combining total material balance equations as follows, using the nomenclature of Fig. 8.45: Overall material balance: Mixing point = M = L0 + F '+ F "+Vn +1 = V1 + Ln (1) (2) Material balance around stage m: F '+Vm+ 1 + Lm- 1 = Vm + Lm Material balance around stage p: (3) F "+Vp + 1 + Lp -1 = Vp + Lp Operating point P' : P' = V1 - L0 = ....... = Vm - Lm-1 (4) (5) Operating point P'' : P'' = Vm+1 - Lm = ....... = Vp - Lp-1 Operating point P''' : P''' = Vp+1 - Lp = ....... = Vn+1 - Ln (6) (7) From Eqs. (1) and (4): V1 - L0 = M - (L0 + Ln) = M - J = P' where J = (L0 + Ln) JLn L0 Locate J by the inverse lever arm rule, = (8) JL0 Ln Then operating point P' is the intersection of lines through V1 and L0 , and through M and J Operating point P'' is obtained by combining Eqs. (2) and (5) to give: (9) P " = Vm+1 - Lm = Vm - Lm-1 - F ' = P ' - F ' Operating point P''' is obtained by combining Eqs. (3) and (6) to give: P ''' = V p - L p -1 - F '' = P '' - F '' = P ' - ( F ' + F '' ) = P ' - K (10) where K = (F' + F") Locate K by the inverse lever arm rule, F '' (11) ' KF '' F Then operating point P''' is the intersection of lines through Vn+1 and Ln , and through P' and K Operating point P''' is the intersection of lines through P' and F' , and through P''' and F'' = These constructions are illustrated in the right-triangle diagram on the next page. KF ' Exercise 8.19 (continued) Analysis: (continued) Exercise 8.20 Subject: Extraction of methylcyclohexane (MCH) from n-heptane (C) with aniline (S) at 25oC in a countercurrent-stage extractor. Given: Feed of 50 wt% MCH in C. On a solvent-free basis, extract contains 95 wt% MCH and raffinate contains 5 wt% MCH. Reflux at both ends as in Fig. 8.26a. Minimum extract reflux ratio = 3.49. Equilibrium data in Exercise 8.22. Find: With a right-triangle diagram: (a) Raffinate reflux ratio. (b) Amount of aniline that must be removed by solvent removal at the top of the extractor, as shown in Fig. 8.26a. (c) Amount of solvent that must be added to the mixer at the bottom of the extractor, as shown in Fig. 8.26a. Analysis: The liquid-liquid phase equilibrium data are given, for each phase, as mass % MCH on an aniline-free basis, and mass of aniline per mass of aniline-free mixture. To plot the data on a triangular diagram, it is necessary to convert the equilibrium data to mass % for each of the three components. For example, when the HC-rich layer contains 9.9 wt% MCH on an anilinefree basis and 0.0836 lb aniline per lb aniline-free mixture, we have for 1 lb of aniline-free mixture: 0.099 lb MCH, 0.901 lb n-heptane, and 0.0836 lb aniline, for a total of 1.0836 lb. The corresponding mass % values are: MCH 9.14 wt%, n-heptane 83.15 wt%, and aniline 7.71 wt%. Calculations for the other equilibrium compositions given in the table accompanying Exercise 8.22 are summarized in the following table: Hydrocarbon-rich layer: Wt% Wt% MCH nC7 0.00 92.60 9.14 83.15 18.58 73.41 21.94 69.85 33.73 57.68 40.63 50.68 45.96 45.05 59.66 30.74 67.14 22.86 71.58 18.23 73.57 16.04 83.30 5.41 88.11 0.00 Aniline-rich layer: Wt% Wt% Aniline MCH 7.40 0.00 7.71 0.80 8.01 2.70 8.21 3.00 8.59 4.61 8.69 6.00 8.99 7.40 9.60 9.80 10.00 11.32 10.19 12.69 10.39 13.10 11.29 15.59 11.89 16.89 Wt% nC7 6.20 5.99 5.30 5.11 4.50 4.00 3.60 2.98 2.11 1.60 1.39 0.62 0.00 Wt% Aniline 93.80 93.21 92.00 91.89 90.89 90.00 89.00 87.22 86.57 85.71 85.51 83.79 83.11 The triangular liquid-liquid equilibrium diagram is given on the next page. Referring to Fig. 8.26a, compositions of feed, F, solvent-free extract, D, and solvent-free raffinate, B, are plotted. Exercise 8.20 (continued) Analysis: (continued) Analysis: (continued) Exercise 8.20 (continued) By construction on the triangular diagram on the previous page, the compositions of the solvent-containing extract, VN , and the solvent-containing raffinate, L1 , are determined as intersections of straight lines (drawn from the solvent-free points toward the 100% solvent point, SB ) with the equilibrium curve. The wt% compositions of these points are read from the diagram to be: Wt%: Component Extract, VN Raffinate, L1 Aniline, S 84.0 7.5 n-Heptane, C 1.0 88.0 Methylcyclohexane, MCH 15.0 4.5 Total: 100.0 100.0 Take a basis of 1,000 kg/h of feed and make material balances for the minimum extract reflux condition. Because the feed is 50 wt% MCH and 50 wt% C, and the two solvent-free product compositions are symmetrical with respect to MCH and C, the solvent-free material balance is quickly determined to be: kg/h: Component Feed, F Extract, D Raffinate, B n-Heptane, C 500 25 475 Methylcyclohexane, MCH 500 475 25 Total: 1,000 500 500 (b) Assume the given minimum extract reflux ratio = 3.49 = LR/D in Fig. 8.36a. Then, LR = 3.49(500) = 1,745 kg/h. Therefore, the flow rate of solvent-free extract in stream VN = 1,745 + 500 = 2,245 kg/h. But, from above, this stream contains 84 wt% aniline solvent. Therefore, the flow rate of aniline in stream VN = (84/16)(2,245) = 11,790 kg/h. This is the flow rate of aniline that must be removed by the separator at the top of the cascade. (c) From above, the raffinate L1 leaving at the bottom of the cascade is 7.5 wt% aniline. Therefore, aniline flow rate in B is (7.5/92.5)(500) = 40 kg/h and B = 540 kg/h The entering solvent rate SB is, therefore, 40 + 11,790 = 11,830 kg/h. This is the amount of fresh solvent that must be added to the mixer at the bottom of the column. (a) Referring to Fig. 8.36a, define the raffinate reflux ratio as (L1 - B)/B. Because the cascade is operating at minimum reflux, assume infinite stages so that VB is in equilibrium with L1. From the triangular diagram on the previous page, the composition of VB is obtained by the dashed tie line from L1, which gives 93.5 wt% aniline. A total material balance around the bottom Mixer gives: SB - B = 11,830 - 540 = 11, 290 = VB - L1 . An aniline balance around the bottom Mixer gives: 11,830 - 40 = 0.935VB - 0.075L1 . Solving these two equations gives L1 = 1,435 kg/h. Therefore, the raffinate reflux ratio = (1,435 - 540)/540 = 1.66. Exercise 8.21 Subject: Liquid-liquid extraction of hafnium from zirconium in an aqueous nitric acid solution, using tributyl phosphate (TBP) as the solvent Given: A flowsheet in Fig. 8.46 for a process comprised of 14 equilibrium stages and a solvent stripping unit. Feed consisting of 1.0 L/h of 5.10 N HNO3, containing 127 g of dissolved Hf and Zr oxides per liter, including 22,000 ppm by wt. Hf, enters Stage 5. Stages 5 to 14 constitute an extraction section. TBP enters Stage 14. Scrubbing water enters Stage 1, with Stages 1 to 4 constituting a scrubbing section. Aqueous raffinate is removed from Stage 14. Organic-rich extract leaving Stage 1 enters the Stripping Unit, which recovers the TBP for recycle by stripping with fresh water. The aqueous product from stripping leaves the process. Stagewise data for the distribution of both Hf and Zr between TBP (containing nitric acid) and aqueous nitric acid. Assumptions: The 22,000 ppm of Hf refers to a basis of Hf plus Zr, not to the total feed. Find: (a) (b) (c) Analysis: (a) and (b) A material balance. A check on consistency of the given data. Advantages of the process. Need for all stages. The molecular weights of the components involved are: Component Molecular weight Hf 178.6 Zr 91.22 HfO2 210.60 123.22 ZrO2 63.01 HNO3 From the CRC Handbook, the density of aqueous 5.1 N HNO3 = 1.166 g/cm3 Neglect the effect of dissolved oxides on this density. 1.0 L/h = 1,000 cm3/L. Therefore, in the feed have 1,166 g/h of combined HNO3 and H2O. For 5.1 N HNO3 = 5.1 mol/L of HNO3, have 5.1(63.01) = 321.4 g/h HNO3 in the feed. Therefore, the H2O in the feed = 1,166 - 321.4 = 844.6 g/h. Since the feed includes 127 g/h of combined Hf and Zr oxides, the total feed flow rate = 1,166 + 127 = 1,293 g/h. Since the feed contains 22,000 parts of Hf per 1,000,000 parts of Zr, we have: (22,000/1,000,000 = 0.022 g Hf/g Zr. This corresponds to 0.022(210.6/178.6)/(123.22/91.22) = 0.0192 g HfO2/g ZrO2. Thus, of the 127 g/h of oxides, we have 127(0.0192/1.0192) = 2.39 g/h of HfO2 and 127 - 2.39 = 124.61 g/h of ZrO2. In summary, the feed contains: Component g/h g/L mol/L HfO2 2.39 2.39 0.0194 ZrO2 124.61 124.61 0.5917 HNO3 321.4 321.4 5.1 H2 O 884.6 884.6 49.1 Total: 1,293.0 1,293.0 Exercise 8.21 (continued) Analysis: (continued) HfO2 balance: From the given Stagewise Analyses table, the aqueous phase leaving Stage 14 is the final raffinate and contains 3.54 g total oxides /L and g Hf/g Zr = 0.72. Therefore, g HfO2/g ZrO2 = 0.72(210.6/178.6)/(123.22/91.22) = 0.6285. Then, g HfO2/L = 3.54(0.6285/1.6285) = 1.366 g HfO2/L and g ZrO2/L = 3.54 - 1.366 = 2.174. From the Stagewise Analyses table, it appears that the concentration of Hf in the extract leaving Stage 1 is negligible. Therefore, all of the HfO2 entering in the feed exits in the raffinate. Therefore, with a feed of , QF, of 1 L/h with 2.39 g/L of HfO2 and a raffinate with 1.366 g/L HfO2, the volumetric flow rate of raffinate = 1.0(2.39/1.366) = 1.75 L/h = QR . Total oxide balance: All of the oxides in the feed leave in either the raffinate from Stage 14 or the aqueous product (stripped extract, E) from the Stripper. Using the Stagewise Analyses table, a total oxide material balance gives: 127QF = 3.54QR + 76.4QE , where QF = 1.0 L/h and QR = 1.75 L/h. Therefore, solving, QE = [127 - 3.54(1.75)]/76.4 = 1.58 L/h of stripped extract, which contains 76.4 g ZrO2/L. Total balance: Assume that the total volume flows in = total volume flows out. Let QST = volumetric flow rate of stripping water, and QSC = volumetric flow rate of scrubbing water. Then a volumetric flow balance gives: QST + QSC = QE + QR - QF = 1.58 + 1.75 - 1.0 = 2.33 L/h. Nitric acid balance: The Stagewise Analyses data around Stage 1 indicate that the scrubbing water contains HNO3. Let x = mol/L of HNO3 in the scrubbing water, but assume no HNO3 in the stripping water. An overall nitric acid balance gives: 5.1(1) +xQSC = 2.56QR + 3.96QE . Therefore, xQSC = [2.56(1.75) + 3.96(1.58) - 5.1] = 5.64 mol/h = HNO3 entering in scrub water. Total balance around stripper: QST + Q1 = QE + QS = 1.58 + QS (1) HNO3 balance around stripper: Using the Stagewise Analyses table, 1.95Q1 + 0 = 3.96(1.58) + 0.65Qs = 6.26 + 0.65Qs (2) Oxide balance around stripper: Using the Stagewise Analyses table, 22.2Q1 = 76.4(1.58) = 120.7 g/h of ZrO2 Therefore, solving, Q1 = 120.7/22.2 = 5.44 L/h. From Eq. (2), QS = [1.95(5.44) - 6.26]/0.65 = 6.69 L/h From Eq. (1), QST = 1.58 + 6.69 - 5.44 = 2.83 L/h From above, QST + QSC = 2.33. Therefore, this results in an impossible negative QSC . The data are inconsistent, at least for the Hf/Zr ratio in the final raffinate. This error has a significant effect on the volumetric flow rate of the raffinate, which then affects other flow rates. (c) The extractor as shown removes almost all of the ZrO2 and is, therefore, effective. However, the Hf/Zr ratio in the aqueous phase, as shown in the Stagewise Analyses table, decreases, rather than increases, in Stages 13 and 14. This indicates the need for more nitric acid in the recycle solvent. Also, Stages 1 and 2 appear to do little extraction and might be eliminated. Exercise 8.22 Subject: Extraction of diphenylhexane (DPH) from docosane (C) with furfural (U) at 45oC in a continuous, countercurrent, multistage liquid-liquid extraction system. Given: Feed, F, of 5,000 kg/h of 65 wt% DPH, 28 wt% C, and 7 wt% U. Solvent is pure U. Raffinate to contain 12 wt% DPH. Liquid-liquid equilibrium data from Exercise 8.15. Find: (a) Minimum solvent flow rate. (b) Flow rate and composition of extract at minimum solvent rate. (c) Number of equilibrium stages at a solvent rate of 1.5 times the minimum. The given liquid-liquid equilibrium data at 45oC are plotted in the right-triangle diagram below. Included on the diagram are composition points F for the feed, RN for the raffinate on the equilibrium curve, and S for the solvent. (a) To find the minimum solvent flow rate, Smin , corresponding to an infinite number of equilibrium stages, A straight line is drawn on the diagram below that joins points RN and S and extends to either side of the diagram. Similar lines are extended from the tie lines to where they intersect the first line. The construction is similar to Figure 8.18. For this exercise the intersection points are closest to the raffinate side as they are in Figure 8.18. Therefore the intersection farthest away corresponds to Smin. and locates the composition of the final extract, E1. From either the inverse lever-arm rule or a mass balance, the minimum solvent rate = 2,190 kg/h. b) From the diagram, the composition of the extract is: 2 wt% docosane 28 wt% DPH 70 wt% furfural The composition of the raffinate is: 83 wt% docosane 12 wt% DPH 5 wt% furfural The total flow rate of extract + raffinate = feed + min. solvent = 5,000 + 2,190 = 7,190 kg/h Using the inverse lever-arm rule or mass balances, E1 = 3,355 kg/h and RN = 3,835 kg/h (c) The solvent rate for finite stages = 1.5 Smin = 1.5(2,190) = 3,285 kg/h. The mixing point for this rate is at a composition of: DPH: 0.28(5,000)/(5,000 + 3,285) = 0.169 or 16.9 wt% Furfural: [0.07(5,000) + 3,285]/(5,000 + 3,285) = 0.439 or 43.9 wt% The second figure below shows the new extract point for this solvent rate and a mixing point. The operating point, P, is the intersection of the two lines, from through RN and S; the other through F and E1. The point is not shown because far to the left off the page. However, using that point and alternating between operating lines and equilibrium tie lines, as in Figure 8.17, only 2 equilibrium stages are required to move from E1 to RN . Analysis: Analysis: (a) and (b) Exercise 8.22 (continued) S E1 F RN Pmin is further down Analysis: (a) and (b) Exercise 8.22 (continued) S E1 F RN P is further down Exercise 8.23 Subject: Extraction of diphenylhexane (DPH) from docosane (C) with furfural (U) at 45oC in a single-stage and a continuous, countercurrent, multistage liquid-liquid extraction system. Given: Feed, F, of 1,000 kg/h of 20 wt% DPH and 80 wt% docosane. Solvent is pure furfural. . Liquid-liquid equilibrium data from Exercise 8.15. Find: (a) For a single equilibrium stage, the compositions and flow rates of extract and raffinate for solvent rates of 100, 1,000, and 10,000 kg/h. (b) Minimum solvent flow rate to form two liquid phases. (c) Maximum solvent flow rate to form two liquid phases. (d) For two countercurrent-flow equilibrium stages, the compositions and flow rates of extract and raffinate for a solvent flow rate of 2,000 kg/h. Analysis: (a) First, calculate the composition of the mixing point for each solvent flow rate, noting that the feed is 800 kg/h of docosane and 200 kg/h of DPH. Solvent flow rate, kg/h Feed + Solvent, kg/h Wt% DPH Wt% Furfural 100 1,100 0.182 0.091 1,000 2,000 0.100 0.500 10,000 11,000 0.0182 0.910 These three mixing points can be plotted on the triangular phase diagram shown below. Then a tie line is drawn through each mixing point. The intersections of the tie lines with the bimodal curve give the following compositions from a single equilibrium stage, in the manner of Figure 4.16, for each of the three cases: Solvent flow rate, kg/h Raffinate composition, wt%: DPH Furfural Docosane Extract composition, wt% DPH Furfural Docosane 100 18.2 5.7 76.1 18.0 80.5 1.5 1,000 10.1 4.9 85.0 9.9 1.1 89.0 10,000 2.1 4.2 93.7 1.9 98.0 0.1 Analysis: (a) continued: Exercise 8.23 (continued) The following diagram shows the construction for a solvent rate of 1,000 kg/h. R R F R M E S Analysis: (a) continued: Exercise 8.23 (continued) Next, mass balances can be used to compute the flow rates of the raffinate and the extract for each of the three cases, as follows. Case 1. Solvent flow rate = 100 kg/h DPH mass balance: 0.182 R + 0.180 E = 200 Total mass balance: R + E = F + S = 1,000 + 100 = 1,100 Solving, R = 1,000 kg/h and E = 100 kg/h Case 2 Solvent flow rate = 1,000 kg/h DPH mass balance: 0.101 R + 0.099 E = 200 Total mass balance: R + E = F + S = 1,000 + 1,000 = 2,000 Solving, R = 1,000 kg/h and E = 1,000 kg/h Case 3. Solvent flow rate = 10,000 kg/h Docosane mass balance: 0.937 R + 0.000 E = 800 Total mass balance: R + E = F + S = 1,000 + 10,000 = 11,000 Solving, R = 850 kg/h and E = 10,150 kg/h (b) and (c) The minimum and maximum solvent flow rates for the formation of two liquid phases is obtained from the intersection with the bimodal curve of a straight line between the feed and solvent compositions. The construction is shown on a separate page below. The solvent rates are obtained using the inverse lever-arm rule with that diagram, resulting in the following results: Smin = 64 kg/h and Smax = 142,000 kg/h (d) The determination of the raffinate and extract obtained from a two-equilibrium-stage, countercurrent-flow extraction system requires a trial-and-error procedure. One approach is to assume a sequence of equilibrium raffinate compositions and step off the corresponding stages until 2 stages result. A good way to start the start sequence is to solve the single-stage problem in a manner similar to that in Part (a) of this exercise. That construction is shown below and results in a raffinate containing 6.7 wt% DPH. For two stages, the wt% DPH will be less. In this manner, it is found that the raffinate will contain 2.5 wt% DPH for 2 stages. The construction is shown below and the compositions and flow rates of the corresponding raffinate and extract are as follows for the solvent flow rate of 2,000 kg/h: Flow rate, kg/h DPH, wt % Docosane, wt% Furfural, wt% Raffinate 855 2.5 93.3 4.2 Extract 2,145 8.0 0.5 91.5 Analysis: (b, c) continued: Exercise 8.23 (continued) The following diagram shows the construction to determine minimum and maximum solvent rates to produce two liquid phases. R R F S min S max S Analysis: (d) continued: Exercise 8.23 (continued) The diagram bellows shows construction for an assumed raffinate containing 2.5 wt% DPH and two stages. The mixing point, M, corresponds to 200 kg/h of DPH and 2,000 kg/h of furfural in a total of 3,000 kg/h sent to the extraction unit. This gives 6.7 wt% DPH and 66.7 wt% furfural. R R Analysis: (d) continued Exercise 8.23 (continued) Exercise 8.24 Subject: Extraction of acetone (A) from water (C) by 1,1,2-trichloroethane (S) at 25oC Given: Liquid mixture of 27 wt% A and 73 wt% C. Liquid-liquid equilibrium data. Raffinate essentially of acetone Find: (a) Minimum solvent-to-feed ratio. (b) Composition of the extract at the minimum solvent-to-feed ratio. (c) Composition of the extract stream leaving Stage 2 from the feed end. Analysis: Using the given equilibrium data in weight fractions, the right-triangle diagram is shown below, where a solid line is used for the equilibrium curve and dashed lines are used for the tie lines, only three of which are given. Additional tie-line locations can be made using either of the two techniques illustrated in Fig. 8.16. (a) The minimum solvent flow rate corresponds to an infinite number of equilibrium stages. To determine this minimum: (1) Points are plotted on the triangular diagram for the given compositions of the solvent, S (pure S), feed, F (27 wt% A, 73 wt% C), and raffinate, RN (0 wt% A on the equilibrium curve), followed by drawing an operating line through the points S and RN and extending it only to the right because the tie lines slope down from left to right; (2) Because the pinch region, where equilibrium stages crowd together, may occur anywhere, additional potential operating lines may be drawn through the tie lines extended to the right until they cross the operating line at the solvent end, drawn in (1) The intersections are not shown in the diagram because they are far off the page. (3) The intersection point farthest from the triangular diagram is the one used to determine the minimum solvent rate. For this exercise, that point is quite sensitive. Assume the controlling line is the one through F, corresponding to a pinch point at the feed end of the cascade; (4) An operating line is drawn from Pmin through F to the determine the extract E1, followed by drawing lines from E! to RN and from S to F, with the intersection of these two lines determining the mixing point, M. The wt% A at the mixing point is approximately 17 wt% From Eq. (8-10), S min ( xA ) F - ( xA ) M 0.27 - 0.17 = 0.59 = = F ( xA )M - ( xA )S 0.17 - 0.0 (b) The composition of the extract at the minimum solvent rate is: 34 wt% acetone, 1.5 wt% water, and 64.5 wt% solvent (c) If the pinch point is at the feed end of the system as assumed above, then the extract leaving Stage 2 is essentially the same composition as the final extract. Analysis: (continued) Exercise 8.24 (continued) Exercise 8.25 Subject: Liquid-liquid extraction of methylcyclohexane (A) from a mixture with n-heptane (C) by aniline (S) as the solvent, using extract reflux. Given: Feed of 50 wt% C and 50 wt% A. Extract to contain 92.5 wt% A in the extract, and 7.5 wt% A in the raffinate, both on a solvent-free basis. Equilibrium data at 25oC and 1 atm. Assumptions: Perfect solvent separator. Find: Using the graphical method of Maloney and Schubert: (a) Minimum stages. (b) Minimum extract reflux ratio. (c) Theoretical stages for an extract reflux ratio = 7. Analysis: The cascade is like that shown in Fig. 8.26b or 8.32, where extract reflux is used. The Maloney-Schubert method for the use of extract reflux is conducted on a Janecke diagram like that in Figs. 8.31 and 8.33. For the system here, the diagram is shown on the next page, where the composition on the ordinate, Y, is mass of solvent (S) per mass of solvent-free extract (upper curve), and per mass of solvent-free raffinate (lower curve). The abscissa composition, X, is mass of solute (A) per mass of solvent-free extract or raffinate. The dashed lines are the liquidliquid equilibrium tie lines. (a) For the minimum number of equilibrium stages, the construction is that is Fig. 8.31a, where all operating lines on the Y - X diagram are vertical lines. As shown on the following page, the number of stages stepped off by alternating between equilibrium tie lines (many of which are added by interpolation) and vertical operating lines starting from the raffinate, L1, and ending at the extract, VN , is just less than 10. Call the minimum number of stages 10. (b) For the minimum extract reflux ratio, the construction is that shown in Figs. 8.31a and b and is given here on the third page of this exercise. A tie line through the feed point, F, establishes the point P' on a vertical line through X = 0.975, which is the solvent-free composition of the solute in the extract. No other line drawn through a tie line to the right of the feed tie line or drawn from P'' points (established by tie lines to the right of the feed tie line) through the feed point, intersects the vertical line through X = 0.975 at a higher P' point. Therefore, as shown, P'min = 20.5. Using Eq. (8-18) with measurements from the diagram, the minimum reflux ratio = line P'VN / line VNLR = LR/D = 2.5 (c) With an extract reflux ratio, LR/D = 7 and line VNLR = 5.0, P' = 5(7) + 5 = 40. The construction follows that in Fig. 8.33, as given on the fourth page of this exercise, where 16 stages are stepped off with the feed stage at 11 from the raffinate end. Analysis: (a) (continued) Exercise 8.25 (continued) Analysis: (b) (continued) Exercise 8.25 (continued) Analysis: (c) (continued) Exercise 8.25 (continued) Exercise 8.26 Subject: Liquid-liquid extraction of A from a mixture with B, by solvent C, using extract reflux. Given: Feed of 1,000 lb/h containing 35 wt% A and 65 wt% B on a solvent-free basis, but as a saturated raffinate. Extract to contain 83 wt% A and 17 wt% B on a solvent-free basis. Raffinate to contain 10 wt% A and 90 wt% B on a solvent-free basis. Liquid-liquid equilibrium data for A,B,C system. Assumptions: Solvent is completely removed from the extract leaving Stage N in Fig. 8.26b. Thus, the extract reflux, LR , contains no solvent. Find: Minimum mass of reflux. Number of equilibrium stages for twice the minimum reflux ratio. Masses of streams on a solvent-free basis. Analysis: Solve the problem using solvent-free coordinates with a Y - X diagram for the solute, A. The same results should be obtained using triangular coordinates. The configuration of the extraction cascade is that in Fig. 8.26b, where the feed is F , the solvent-free extract is D, and the raffinate is B or say B' on a solvent-free basis. The equilibrium data are plotted on the next page as: wt% A in extract layer YA = (wt% A + wt% B) in extract layer wt% A in raffinate layer XA = (wt% A + wt% B) in raffinate layer where these values are computed from the given equilibrium data as follows: Extract: wt% A 0.0 1.0 1.8 3.7 6.2 9.2 13.0 18.3 24.5 31.2 wt% B 7.0 6.1 5.5 4.4 3.3 2.4 1.8 1.8 3.0 5.6 Raffinate: wt% A wt%B 0.0 92.0 9.0 81.7 14.9 75.0 25.3 63.0 35.0 51.5 42.0 41.0 48.1 29.3 52.0 20.0 47.1 12.9 31.2 5.6 XA 0.000 0.099 0.166 0.287 0.405 0.506 0.621 0.722 0.785 0.848 YA 0.000 0.141 0.247 0.457 0.653 0.793 0.878 0.910 0.891 0.848 Included on the plot are the compositions of the feed, the final extract, and the final raffinate. Analysis: (continued) Exercise 8.26 (continued) Exercise 8.26 (continued) Analysis: (continued) An overall component material balance for A, using solvent-free flow rates, gives: 1,000(0.35) = 350 = YDD + XBB' = 0.83D = 0.10B' (1) An overall total material balance (on a solvent-free basis) gives: 1,000 = D + B' (2) Solving Eqs. (1) and (2), D = 342.5 lb/h and B' = 657.5 lb/h. To obtain the total flows, including the solvent contributions, the above table of Y-y and X-x compositions can be interpolated, from plots given on the next page. For the feed as a raffinate, saturated with solvent, with XA = 0.35, the wt% C = 12.5%. Therefore, the total feed rate = 1,000/(1 - 0.125) = 1,140 lb/h, including 140 lb/h of solvent. The total flow rate of the final raffinate, including the solvent, is found in a similar manner. For XA = 0.10, the wt% C for a saturated raffinate = 9.3. Therefore, the total final raffinate, L1 = B = 657.5(1 - 0.093) = 724.9 lb/h. The solvent in the final total raffinate = 724.9 - 657.5 = 67.4 lb/h. The total flow rate of the exiting solvent, SD , can not be found yet because the flow rate, VN , can not be determined until the extract reflux rate is known. However, for, YA = 0.83 in the extract leaving Stage N, the wt% C for an extract saturated with solvent = 87.5. Similarly, the entering fresh solvent rate, SB can not be determined yet. From the shape of the equilibrium curve on the previous page, at minimum extract reflux, the intersection of the operating lines for the extractor sections above and below the feed would appear to be at the intersection with a vertical line for the wt% A (solvent-free) in the feed, and with the equilibrium curve. Assume this is so. Then at minimum extract reflux, the pinch point will be at the feed stage. Accordingly, the composition of raffinate stream, LF+1, entering the feed stage F (in Fig. 8.26b) is identical to the composition of the feed. From the Y - X plot on the previous page, the solvent-free composition of the extract, VF , is that in equilibrium with the X = 0.35 of the feed, or Y = 0.56. From the figures on the next page, the wt% C in this extract is 91.3. Now, the entering and leaving flow rates can be calculated by material balances for the section of the extractor above the feed stage in Fig. 8.26b. An overall solvent-free material balance gives the following, where the prime superscript indicates a flow rate on a solvent-free basis: VF' = L'F +1 + D = L'F +1 + 342.5 (3) A similar balance on the solute, A, gives: YVF VF' = 0.56VF' = X LF +1 L'F +1 + YD D = 0.35 L'F +1 + 0.83(342.5) = 0.35 L'F +1 + 284.3 (4) Solving Eqs. (3) and (4), gives: VF' = 782.9 lb / h and L'F +1 = 440.4 lb / h The corresponding total flow rates (including solvent) are: VF = 782.9/(1 - 0.913) = 8,998.9 lb/h of which, 8,998.9 - 782.9 = 8,216.0 is solvent LF+1 = 440.4/(1 - 0.125) = 503.3 lb/h of which 503.3 - 440.4 = 62.9 is solvent Therefore, the exiting solvent rate, SD = 8,216.0 - 62.9 = 8,152.1 lb/h Analysis: (continued) Exercise 8.26 (continued) Exercise 8.26 (continued) Analysis: (continued) The extract, VN , contains 83 wt% A on a solvent-free basis. Therefore, From the equilibrium plots on a previous page, the corresponding wt% C in this extract is 87.5 wt%. Therefore, the total flow rate of VN = 8,152.1/(0.875) = 9,316.7 lb/h By total material balance around the solvent removal and stream divider above the top of the cascade (as shown in Fig. 8.26b), LD = VN - SD - D = 9,316.7 - 8,152.1 - 342.5 = 822.1 lb/h. Therefore, the minimum extract reflux rate = 822.1 lb/h, and the minimum extract reflux ratio = 822.1/342.5 = 2.40 To compute the number of equilibrium stages, take twice the minimum extract reflux flow rate or 2(822.1) = 1,644.2 lb/h with the same composition of 83 wt% A and 17 wt% B. The solvent-free extract is the same flow rate as before of D = 342.5 lb/h. Therefore, the solvent-free flow rate of V'N = 1,644.2 + 342.5 = 1,986.7 lb/h. Since, this flow is 87.5 wt% solvent, total flow rate VN = 1,986.7/(1 - 0.875) = 15,893.6 lb/h. The corresponding flow rate of exiting solvent is SD = 15,893.6 - 1,986.7 = 13,906.9 lb/h The slope of the solvent-free operating line at the top of the cascade = LD/V'N = 1,644.2/1,986.7 = 0.828. However, this slope is not constant because as we move down from the top of the cascade, all three components undergo mass transfer between the two liquid phases in an overall non-equimolar fashion. At the bottom of the cascade, by overall solvent material balance, the entering solvent flow rate = SB = 13,906.9 + 67.4 - 140 = 13,834.3 lb/h Next, determine two Y - X points on the operating line above the feed entry by writing the following balances around a section of stages from Stage n above the feed to Stage N at the top, in terms of total stream mass flows, V and L, and mass fractions (including solvent), y and x: Total mass balance: Solute mass balance: Solvent mass balance: Vn-1 = Ln + 14,249.4 (yn-1)A Vn+1 - (xn)A Ln = 284.3 (yn-1)C Vn+1 - (xn)C Ln = 13,906.9 (5) (6) (7) Eqs. (5) to (7) are solved by trial and error for a selected value of Xn. To do this, combine Eq. (5) with (6) to eliminate Vn-1, and combine Eq. (5) with (7) to eliminate Vn-1. We now have two equations for Ln. The selected value of Xn fixes the mass fractions of A and C in the raffinate. Then find a value of Yn-1 (which fixes the mass fractions of A and C in the extract) that gives identical values of Ln. The results are as follows, with these two points for Y and X, together with Y = X = 0.83 establishing the upper operating line, as shown in the Y - X diagram on the second page of this exercise: Xn 0.621 0.405 Yn-1 0.670 0.520 Vn-1, lb/h 15,790 15,400 Ln, lb/h 1,541 1,151 Exercise 8.26 (continued) Analysis: (continued) Next, determine two Y - X points on the operating line below the feed entry by solving the following balances around a section of stages from Stage m below the feed to bottom Stage1: Total mass balance: Solute mass balance: Solvent mass balance: Vm = Lm+1 - 13,109.4 (ym)A Vm - (xm+1)A Lm+1 = -65.7 (ym)C Vm - (xm+1)C Lm+1 = 13,766.9 (8) (10) (9) . The results are as follows, with these two points for Y and X, together with Y = X = 0.10 establishing the lower operating line, as shown in the Y - X diagram on the second page of this exercise: Xm+1 0.287 0.166 Ym 0.390 0.205 Vm, lb/h 15,182 15,064 Lm+1,lb/h 2,073 1,955 From the Y - X plot on the second page of this exercise, 10 equilibrium stages are stepped off, with 2 above the feed stage and 7 below. The component material balance in lb/h is as follows: Component A B C Total Feed 350.0 650.0 140.0 1140.0 Solvent in 0.0 0.0 13834.3 13834.3 Extract 284.3 58.2 0.0 342.5 Raffinate 65.7 591.8 67.4 724.9 Solvent out 0.0 0.0 13906.9 13906.9 Reflux, LD 1364.7 279.5 0.0 1644.2 Exercise 8.27 Subject: Extraction of methylcyclohexane (MCH) from n-heptane (C) with aniline (S) at 25oC in a countercurrent-stage extractor. Given: Feed of 50 wt% MCH in C. On a solvent-free basis, extract contains 95 wt% MCH and raffinate contains 5 wt% MCH. Reflux at both ends as in Fig. 8.26a. Minimum extract reflux ratio = 3.49. Equilibrium data in Exercise 8.22. Find: With the Maloney-Schubert method on a Y - X Janecke diagram: (a) Raffinate reflux ratio. (b) Amount of aniline that must be removed by solvent removal at the top of the extractor, as shown in Fig. 8.26a. (c) Amount of solvent that must be added to the mixer at the bottom of the extractor, as shown in Fig. 8.26a. Analysis: The liquid-liquid phase equilibrium data are given, for each phase, as mass % MCH on an aniline-free basis, and mass of aniline per mass of aniline-free mixture. Thus, the data can be plotted directly to obtain the Janecke diagram, which is shown on the next page. Referring to Fig. 8.26a, the given compositions of the feed, F, the solvent-free extract, D, and the solvent-free raffinate, B, are plotted on the diagram. By construction on the Janecke diagram, on the previous page, the compositions of the solvent-containing extract, VN , and the solvent-containing raffinate, L1 , are determined as intersections of vertical lines (drawn up from the solvent-free points to the extract and raffinate equilibrium curves. The wt% compositions of these points are read from the diagram with the aid of the tabulated equilibrium data given in the exercise statementto be: Wt%: Component Extract, VN Raffinate, L1 Aniline, S 84.0 7.5 n-Heptane, C 1.0 88.0 Methylcyclohexane, MCH 15.0 4.5 Total: 100.0 100.0 Take a basis of 1,000 kg/h of feed and make material balances for the minimum extract reflux condition. Because the feed is 50 wt% MCH and 50 wt% C, and the two solvent-free product compositions are symmetrical with respect to MCH and C, the solvent-free material balance is quickly determined to be: kg/h: Component Feed, F Extract, D Raffinate, B n-Heptane, C 500 25 475 Methylcyclohexane, MCH 500 475 25 Total: 1,000 500 500 Exercise 8.27 (continued) Analysis: (continued) Exercise 8.27 (continued) Analysis: (continued) (b) Assume the given minimum extract reflux ratio = 3.49 = LR/D in Fig. 8.36a. Then, LR = 3.49(500) = 1,745 kg/h. Therefore, the flow rate of solvent-free extract in stream VN = 1,745 + 500 = 2,245 kg/h. But, from above, this stream contains 84 wt% aniline solvent. Therefore, the flow rate of aniline in stream VN = (84/16)(2,245) = 11,790 kg/h. This is the flow rate of aniline that must be removed by the separator at the top of the cascade. (c) From above, the raffinate L1 leaving at the bottom of the cascade is 7.5 wt% aniline. Therefore, aniline flow rate in B is (7.5/92.5)(500) = 40 kg/h and B = 540 kg/h The entering solvent rate SB is, therefore, 40 + 11,790 = 11,830 kg/h. This is the amount of fresh solvent that must be added to the mixer at the bottom of the column. (a) Referring to Fig. 8.36a, define the raffinate reflux ratio as (L1 - B)/B. Because the cascade is operating at minimum reflux, assume infinite stages so that VB is in equilibrium with L1. From the Janecke diagram on the previous page, the composition of VB is obtained by the dashed tie line from B, which gives 14.5 kg aniline/ kg solvent-free extract and 5.7 wt% MCH in the extract. This gives 14.5/15.5 x 100% = 93.5 wt% aniline. A total material balance around the bottom Mixer gives:SB - B = 11,830 - 540 = 11, 290 = VB - L1 . An aniline balance around the bottom Mixer gives:11,830 - 40 = 0.935VB - 0.075L1. Solving these two equations gives L1 = 1,435 kg/h. Therefore, the raffinate reflux ratio = (1,435 - 540)/540 = 1.66. Exercise 8.28 Subject: Design of a mixer-settler unit for extraction of acetic acid (A) from water (C) by isopropyl ether (S). Given: 12,400 lb/h of 3 wt% A in C. 24,000 lb/h of S. 1.5 minutes residence time in mixer. 4 gal/min-ft2 in settling vessel. Find: (a) (b) (c) (d) Diameter, DT , and height, H, of mixing vessel if H/DT = 1. Agitator Hp for mixing vessel. Diameter, DT , and length, L, of settling vessel, if L/DT =4. Residence time in settling vessel in minutes. Analysis: First compute total volumetric flows into mixer vessel. Because feed is dilute in A, assume density of feed is that of water = 62.4 lb/ft3. From Perry's Handbook, specific gravity of S = 0.725 or a density of 0.725(62.4) = 45.2 lb/ft3. Therefore, total flow rate of feed plus solvent = Q = 12,400/62.4 + 24,000/45.2 = 730 ft3/h or 12.2 ft3/min. (a) Vessel volume = V = Qtres = 12.2(1.5) = 18.3 ft3 For a cylindrical vessel of H/DT = 1, V = DT2H/4 = DT3/4 Solving, DT3 = 4V/ = 4(18.3)/3.14 = 23.3 ft3. Therefore, DT = H = 2.85 ft (b) Assume 4 Hp/1,000 gallons for mixer agitator power. Gallons in vessel = 7.48V = 7.48(18.3) = 137 gal Therefore, need 4(137)/1,000 = 0.55 Hp. (c) Assume volumetric flow into mixer = volumetric flow leaving settler Therefore, Q = 12.2 ft3/min or 7.48(12.2) = 91 gpm The ft2 disengaging area needed = 91/4 = 22.8 ft2 = DTL = DT(4DT) = 4DT2 Solving, DT = 2.4 ft and L = 4DT = 4(2.4) = 9.6 ft. (d) Settler volume = V = DT2L/4 = 3.14(2.4)2(9.6)/4 = 43.4 ft3 Residence time of phases in settler = tres = V/Q = 43.4/12.2 = 3.6 minutes Exercise 8.29 Subject: Extraction of acetic acid (A) from water (C) by ethyl acetate (S) in available extractor. Given: Cascade of 6 mixer-settler units. Each mixer is 10-ft diameter by 10-ft high with 20-Hp agitator. Each settler is 10-ft diameter by 40-ft long. Feed and solvent compositions , and solvent-to-feed mass ratio, are those in Fig. 8.1. Find: Allowable feed flow rate. Analysis: From Fig. 8.1, the total entering component flow rates are as follows, with liquid densities from Perry's Handbook. Assume ideal liquid solutions. Component Acetic acid Water Ethyl acetate lb/h 6,660 26,100 68,600 density, lb/gal 8.74 8.33 7.51 gpm 12.7 52.2 152.2 Therefore, total flow rate = Q = 12.7 + 52.2 + 152.2 = 217.1 gpm The volume of one mixer = V = DT2H/4 = 3.14(10)2(10)/4 = 785 ft3 or 5,872 gal. First consider the mixer residence time factor. With the flow rates in Fig. 8.1, residence time in mixer = tres = V/Q = 5,872/217 = 27 minutes. Since the liquid phase viscosities are each less than 5 cP (actually less than 1 cP) and the density difference computes to be 0.106, which is greater than 0.10, the required residence time in the mixer should not exceed 5 minutes, Therefore, should be able to handle a feed rate of the acetic acid-water mixture of (27/5)(6,660 + 23,600) = 163,400 lb/h or 27/5 = 5.4 times that in Fig. 8.1. Next, consider the mixer agitator Hp factor. The vessel volume in gallons = 5,872. Therefore, using the criterion of 4 Hp/1,000 gal, need 4(5,872)/1,000) = 23.5 Hp. This is greater than the given 20 Hp. So reduce allowable feed rate of acetic acid in water to 20/23.5(163,400) = 139,000 lb/h. This increases residence time in mixer to 5(23.5)/20 = 5.9 minutes. Now, check settler disengaging factor. Assume that the total volumetric flow leaving the settler is equal to that entering the mixer. Therefore, Q = (27/5)(5/5.9)(217.1) = 994 gpm. Using the settling criterion of 5 gpm/ft2 of DTL, Need DTL = 994/5 = 200 ft2. The available settler has DTL = 10(40) = 400 ft2, which is twice that needed. Therefore, a feed rate of acetic acid + water of 139,000 lb/h is safe. This is 4.6 times that in Fig. 8.1. Exercise 8.30 Subject: Sizing an agitator for extraction of acetic acid (A) from water (C) by isopropyl ether (S) at 25oC. Given: For one of the mixers, flow rates and properties of raffinate and extract are given. Raffinate is dispersed phase in mixer, with residence time of 2.5 minutes. Find: (a) (b) (c) (d) Dimensions of closed, baffled mixing vessel. Diameter of flat-bladed impeller. Minimum rate of rotation of impeller in rpm for complete and uniform dispersion. Power of agitator at minimum rpm. Analysis: First, compute the volumetric flow rates of the extract and raffinate. Extract flow rate = 52,000/45.3 = 1,148 ft3/h or 1,148(7.48)/60 = 143 gpm Raffinate flow rate = 21,000/63.5 = 331 ft3/h or 331(7.48)/60 = 41 gpm Total flow rate = Q = 1,148 + 331 = 1,479 ft3/h or 143 + 41 = 184 gpm Assume this total flow also enters and leave the mixer. (a) For a residence time of 2.5 minutes, volume of mixing vessel = V = Qtres = (1,479/60)(2.5) = 62 ft3 Assume H = DT , then V = DT2H/4 = DT3/4 Solving, DT3 = 4(62)/3.14 = 79 ft3 and DT = H = 4.3 ft. (b) Impeller diameter = Di = (1/3)DT = 4.3/3 = 1.43 ft. (c) To determine the minimum impeller rpm, assume that the volume fractions of the two phases in the mixer are in the same ratio as for the flow rates of the phases leaving the settler. Therefore, the volume fractions are: = 1,148/1,479 = 0.78 and R = 1 - 0.78 = 0.22. Because the raffinate volume fraction is less than 0.26, the raffinate is probably the dispersed phase. Therefore, D = 0.22 and C = 0.78. We are given the following properties, Phase Dispersed (raffinate) Continuous (extract) Density, lb/ft3 63.5 45.3 Viscosity, cP 3.0 1.0 Also, interfacial tension = I = 13.5 dyne/cm = 388,000 lb/h2 Density difference = = 63.5 - 45.3 = 18.2 lb/ft3 Phase mixture density from Eq. (8-23) = M = 45.3(0.78) + 63.5(0.22) = 49.3 lb/ft3 From Eq. (8-24), mixture viscosity is, 15 D D . 1.0 15(3.0)(0.22) . M = C 1+ = 1+ = 16 cP or 3.87 lb/ft-h . C C + D 0.78 10 + 3.0 . Exercise 8.30 (continued) Analysis: (c) (continued) Compute minimum rpm from Eq. (8-25), where, far-right dimensionless group is, 2 387 2 (388,000) . M = = 3.42 10 -16 5 2 2 5 Di M g ( ) (143) (49.3)(4.17 108 )(18.2) 2 . 2 N min M Di D = 1.03 T g Di 2 .76 0.106 D 2 M 5 Di M g 2 ( ) 2 0.084 = 1.03 4.3 1.43 2 .76 (0.22) 0.106 (3.42 10-16 ) 0.084 = 20.88 2 Therefore, N min = 20.88 g (4.17 108 )(18.2) = 20.88 = 9.9 107 (rph2) M Di (49.3)(143) . and Nmin = 9,950 rph or 166 rpm (d) Determine the mixing vessel agitator Hp from Fig. 8.36 (b), which requires NRe. Assume N = Nmin = 9,950 rph From Eq. (8-22), N Re = Di2 N M (143) 2 (9,950)(49.3) . = = 2.6 105 M 387 . From Fig. 8.36 (b), for a mixing vessel with baffles (thus, no vortex), Npo = 5.7 From Eq. (8-21), the power is given by, P= N Po N 3 Di5 M (5.7)(9,950) 3 (143) 5 (49.3) . = = 4.0 106 ft-lbf/h or 2 Hp 8 gc 4.17 10 Compare this to the rule of thumb of 4 Hp/1,000 gallons. Vessel volume = 62 ft3 or 464 gal. Therefore, Hp = 4(464)/1,000 = 1.9 Hp, which is in very good agreement. Exercise 8.31 Subject: Droplet characteristics for extraction of acetic acid (A) from water (C) by isopropyl ether (S) in a mixer unit. Given: Conditions in Exercise 8.27. Find: (a) Sauter mean drop size. (b) Range of drop sizes. (c) Interfacial area of emulsion. Analysis: (a) The Sauter (surface-mean) diameter, dvs , is given by Eq. (8-38) or (8-39), depending on the Weber number. From Eq. (8-37), N We 143 (9,950) 2 (45.3) . Di3 N 2C = = = 33,800 i 388,000 3 Since NWe > 10,000, Eq. (8-39) applies. d vs = 0.39 Di N We (b) -0 . 6 = 0.39(143)(33,900) -0.6 = 0.00107 ft or 0.32 mm . dmin = dvs /3 = 0.32/3 = 0.107 mm dmax = 3dvs = 3(0.32) = 0.96 mm (c) From Eq. (8-36), the interfacial surface area per unit volume of emulsion is, a= 6 D 6(0.22) = = 1, 234 ft2/ft3 d vs 0.00107 Exercise 8.32 Subject: Mass transfer characteristics for extraction of acetic acid (A) from water (C) by isopropyl ether (S) in a mixer unit. Given: Conditions in Exercises 8.27 and 8.28. Diffusivity of acetic acid in each phase. Distribution coefficient of acetic acid. Find: (a) (b) (c) (d) Dispersed-phase mass-transfer coefficient. Continuous-phase mass-transfer coefficient. Murphree dispersed-phase efficiency. Fraction of A extracted. Analysis: (a) For the raffinate, which is the dispersed phase, Eq. (8-40 applies. Diffusivity of A in C = DD = 1.3 x 10-9 m2/s and from Exercise 8.28, dvs = 0.32 mm = 0.00032 m DD 13 10-9 . = 6.6 = 26.8 10 -6 m/s or 0.32 ft/h From a rearrangement of Eq. (8-40), k D = 6.6 d vs 0.00032 (b) For the extract, which is the continuous phase, Eq. (8-50) applies. Diffusivity of A in S = DC = 2.0 x 10-9 m2/s or 7.75 x 10-5 ft2/h (10)(2.42) . C In Eq. (8-44), N Sc = = = 690 C DC (45.3)(7.75 10 -5 ) In Eq. (8-50), N Re = In Eq. (8-50), N Fr = Di2 NC (143) 2 (9,950)(45.3) . = = 381,000 C (2.42) Di N 2 (143)(9,950) 2 . = = 0.34 4.17 108 g In Eq. (8-50), Di /dvs = 1.43/(0.00032/0.3048) = 1,360 In Eq. (8-50, dvs/DT = (0.00032/0.3048)/4.3 = 2.44 x 10-4 In Eq. (8-50, N Eo = From Eq. (8-50), 2 D d vs g (63.5)(0.00032 / 0.3048) 2 (4.17 108 ) = = 0.0752 388,000 2 1/ 2 N Sh C k d = C vs = 1237 10-5 N Sc . DC 1/ 3 N Re 2/3 -1/ 2 D N Fr 5/12 Di d vs d vs DT N Eo 5/ 4 = 1.237 10 -5 (690)1/ 3 (381,000) 2 / 3 (0.22) -1/ 2 (0.34) 5/12 (1,360) 2 (2.44 10 -4 )1/ 2 0.0752 5/ 4 = 889 Exercise 8.32 (continued) Analysis: (b) (continued) kC = ( NSh )C DC d vs = 889(7.75 10 -5 ) = 65.6 ft/h (0.00032 / 0.3048) (c) To compute EMD from Eq. (8-33), KOD must be computed from Eq. (8-28). That equation requires the slope of the equilibrium curve for acetic acid, m = dcC/dcD . From Perry's Handbook, Section 15, the equilibrium constant under dilute conditions in massfraction composition units is KD = 0.429. Converting this to concentration units, mass/unit volume, m =KD (S/C) = 0.429(45.3/63.5) = 0.306. 1 1 = = 0.315 ft/h From Eq. (8-28), KOD = 1 1 1 1 + + k D mk C 0.32 0.306(65.6) Thus, mass transfer is controlled by the dispersed phase. In Eq. (8-33), need a, V, and QD . From Exercise 8.28, a = 1,234 ft2/ft3 , V = 62 ft3, and QD = 331 ft3/h. From Eq. (8-33), KOD aV / QD 0.315(1,234)(62) / 331 E MD = = = 0.986 or 98.6% 1 + KOD aV / QD 1 + 0.315(1,234)(62) / 331 (d) To compute the fraction of solute extracted, the relation in Example 8.7 can not be used because in that example, the extract is the dispersed phase. Here, the raffinate is the dispersed phase. Therefore, derive a new equation. By material balance on the solute, QD cD , in - cD , out = QC cC , out or QD 1 - cD , out c = QC C , out cD , in cD , in cD , out QC cC , out = cD , in QD cD , in 1- cD , out cD , in (1) Therefore, f extracted = 1 - cD , in - cD , out * cD , in - cD , out E MD = = cD , in - cD , out = cC , out 1 cC , out cD , in - 1- m m cD , in (2) Combining Eqs. (1) and (2) and solving for fextracted , E MD 0.986 f extracted = = = 0.511 or 51.1% extracted 0.986 331 E MD QD 1+ 1+ 0.306 1,148 m QC Exercise 8.33 Subject: Design of mixer unit for extraction of benzoic acid (A) from water (C) by toluene (S). Given: Conditions in Example 8.4. Properties of raffinate and extract phases. 6-flat-blade impeller in a closed mixer unit with baffles. Extract phase is dispersed. Find: (a) Minimum rpm of impeller for complete and uniform dispersion. (b) Power requirement of agitator. (c) Sauter mean diameter. (d) Interfacial area. (e) KOD. (f) NOD. (g) EMD. (h) Fractional extraction of A. Analysis: From Example 8.4, DT = 7.9 ft. Use Di = DT/3 = 7.9/3 = 2.63 ft. Assume QR = QF = 500 gpm, and QE = QS = 750 gpm. Therefore, total Q = 500 + 750 = 1,250 gpm. Assume phase volume holdups in the mixing vessel are in proportion to the volumetric flow rates. Therefore, with the extract phase dispersed, = 750/1,250 = 0.60 and C = 0.40. (a) We are given the following properties, Phase Dispersed (extract) Continuous (raffinate) Density, lb/ft3 53.7 62.1 Viscosity, cP 0.59 0.95 Diffusivity of A, cm/s 1.5 x 10-5 2.2 x 10-5 Also, interfacial tension = I = 22.0 dyne/cm = 633,000 lb/h2 Density difference = =62.1 - 53.7 = 8.4 lb/ft3 Phase mixture density from Eq. (8-23) = M = 53.7(0.60) + 62.1(0.40) = 57.0 lb/ft3 From Eq. (8-24), mixture viscosity is, 15 D D . 0.95 15(0.59)(0.60) . M = C 1+ = 1+ = 2.65 cP or 6.41 lb/ft-h C C + D 0.40 0.95 + 0.59 Compute minimum rpm from Eq. (8-25), where, far-right dimensionless group is, 2 6.412 (633,000) M = = 2.96 10 -19 Di5 M g 2 ( ) 2 (2.63) 5 (57.0)(4.17 108 )(8.4) 2 2 N min M Di D = 103 T . g Di 2 .76 0.106 D 2 M Di5 M g 2 ( ) 2 0.084 = 103 . 7.9 2.63 2 .76 (0.60) 0.106 (2.96 10-19 ) 0.084 = 1004 . Exercise 8.33 (continued) Analysis: (a) (continued) g (4.17 108 )(8.4) 2 Therefore, N min = 1004 . = 1004 . = 235 106 (rph2) . M Di (57.0)(2.63) and Nmin =4,840 rph or 81 rpm (b) Determine the mixing vessel agitator Hp from Fig. 8.36 (b), which requires NRe. Assume N = Nmin =4,840 rph From Eq. (8-22), N Re = Di2 N M (2.63) 2 (4,840)(57.0) = = 2.98 105 M 6.41 From Fig. 8.36 (b), for a mixing vessel with baffles (thus, no vortex), Npo = 5.7 From Eq. (8-21), the power is given by, N Po N 3 Di5 M (5.7)(4,840) 3 (2.63) 5 (57.0) . P= = = 1110 106 ft-lbf/h or 5.6 Hp 8 4.17 10 gc Compare this to the rule of thumb of 4 Hp/1,000 gallons. Vessel volume = 334 ft3 or 2,500 gal. Therefore, Hp = 4(2,500)/1,000 = 10 Hp, which is 80% higher. (c) The Sauter (surface-mean) diameter, dvs , is given by Eq. (8-38) or (8-39), depending on the Weber number. From Eq. (8-37), N We = 2.63 (4,840) 2 (57.0) Di3 N 2C = = 41,800 633,000 i 3 Since NWe > 10,000, Eq. (8-39) applies. d vs = 0.39 Di N We -0 . 6 = 0.39(2.63)(41,800) -0.6 = 0.00173 ft or 0.53 mm (d) From Eq. (8-36), the interfacial surface area per unit volume of emulsion is, a= 6 D 6(0.60) = = 2, 080 ft2/ft3 d vs 0.00173 Exercise 8.33 (continued) Analysis: (continued) (e) Eq. (8-28) applies, where, KOD = 1 1 1 + k D mk C For the extract, which is the dispersed phase, Eq. (8-40 applies. D 15 10 -5 . = 188 10-3 cm/s or 0.22 ft/h . From a rearrangement of Eq. (8-40), k D = 6.6 D = 6.6 d vs 0.0527 For the raffinate, which is the continuous phase, Eq. (8-50) applies, with DC = 2.2 x 10-5 cm2/s or 8.53 x 10-5 ft2/h. C (0.95)(2.42) = = 434 In Eq. (8-44), N Sc = C DC (62.1)(8.53 10-5 ) In Eq. (8-50), N Re = Di2 NC (2.63) 2 (4,840)(62.1) = = 904,000 C (0.95)(2.42) In Eq. (8-50), N Fr = Di N 2 (2.63)(4,840) 2 = = 0.148 g 4.17 108 In Eq. (8-50), Di /dvs = 2.63/(0.00173 = 1,520 In Eq. (8-50, dvs/DT = 0.00173/7.9 = 2.19 x 10-4 In Eq. (8-50, N Eo = From Eq. (8-50), 2 D d vs g (53.7)(0.00173) 2 (4.17 108 ) = = 0.1058 633,000 2 1/ 2 N Sh C k d = C vs = 1237 10-5 N Sc . DC 1/ 3 N Re 2/3 -1/ 2 D N Fr 5/12 Di d vs d vs DT N Eo 5/ 4 . . = 1.237 10 -5 (434)1/ 3 (904,000) 2 / 3 (0.60) -1/ 2 (0148) 5/12 (1,520) 2 (2.19 10-4 )1/ 2 01058 kC = N Sh C 5/ 4 = 1,052 1,052(8.53 10-5 ) = 518 ft/h . d vs 0.00173 The slope of the equilibrium curve for dilute conditions is given by Eq. (8-29) as m = dcC/dcD = cC/cD = 1/21 = 0.0476 1 1 1 From Eq. (8-28), K OD = = = = 0.202 ft/h 1 1 1 1 4.55 + 0.405 + + k D mkC 0.22 0.0476(51.8) where, it is seen that the dispersed phase controls the rate of mass transfer. = DC Exercise 8.33 (continued) Analysis: (continued) (f) From Eq. (8-32, N OD = K OD aV (0.202)(2, 080)(334) = = 26.3 QD (750)(60) / 7.48 N OD 26.3 = = 0.963 or 96.3% 1 + N OD 1 + 26.3 (g) From Eq. (8-33), E MD = (h) Because the extract is the dispersed phase, Eq. (3) in Example 8.7 applies, f extracted 750(0.963) QD E MD 1 500 QC m 21 = = = 0.968 or 96.8% extracted QD E MD 750(0.963) 1+ 1+ 1 QC m 500 21 Exercise 8.34 Subject: Diameter of an RDC column for extraction of acetic acid (A) from water (C) by isopropyl ether (S). Given: Conditions in Exercises 8.28 and 8.30. Find: Column diameter. Analysis: Conditions of Exercise 8.28: Flow rate of feed/raffinate = 12,400 lb/h. Flow rate of solvent/extract = 24,000 lb/h Assume raffinate is dispersed. The following properties are given in Exercise 8.30. R = D = 63.5 lb/ft3 and E = C = 45.3 lb/ft3 C = 1.0 cP or 0.000021 lbf-s/ft2 Interfacial tension = I = 13.5 dyne/cm or 0.000924 lbf/ft u Assume o C C = 0.01 i (0.000924)(63.5 - 45.3) Therefore, uo = 0.01 i = 0.01 = 0.177 ft/s C C (0.000021)(45.3) U D mD / D 12,400 / 635 . = = = 0.369 The flooding correlation of Fig. 8.39 is too difficult U C mC / C 24,000 / 45.3 to read for low values of UD/UC . Instead, use Eqs. (8-62) and (8-59). From Eq. (8-62), with UC/UD = 1/0.369 = 2.71, (1) Df = 1+ 8 UC / U D 1/ 2 -3 4 UC / U D - 1 = 1 + 8 2.71 1/ 2 -3 4 2.71 - 1 = 0.257 From Eq. (8-59), using D = (D)f = 0.257, UD UC UD UC + = + = uo 1 - D = 0.177(1 - 0.257) D 1 - D 0.257 1 - 0.257 which simplifies to, 3891U D + 1346U C = 01315 . . . (2) Solving Eqs. (1) and (2) simultaneously, UC = 0.0473 ft/s and UD = 0.0175 ft/s Therefore, at flooding, (UD + UC) = 0.0175 + 0.0473 = 0.0648 ft/s Size the column for 50% of flooding. Therefore, (UD + UC) = 0.0648/2 = 0.0324 ft/s Total volumetric flow rate = Q = 12,400/63.5 + 24,000/45.3 = 725 ft3/h = 0.201 ft3/s Column cross-sectional area = AC = Q/(UD + UC) = 0.201/0.0324 = 6.2 ft2 Column diameter = DC = (4AC/)1/2 = [4(6.2)/3.14]1/2 = 2.8 ft Exercise 8.34 (continued) Analysis: (continued) Conditions of Exercise 8.30: Flow rate of feed/raffinate = 21,000 lb/h. Flow rate of solvent/extract = 52,000 lb/h Raffinate is dispersed. The following properties are given in Exercise 8.30. R = D = 63.5 lb/ft3 and E = C = 45.3 lb/ft3 C = 1.0 cP or 0.000021 lbf-s/ft2 Interfacial tension = I = 13.5 dyne/cm or 0.000924 lbf/ft u Assume o C C = 0.01 i (0.000924)(63.5 - 45.3) Therefore, uo = 0.01 i = 0.01 = 0.177 ft/s (0.000021)(45.3) C C U D mD / D 21,000 / 63.5 = = = 0.288 The flooding correlation of Fig. 8.39 is too difficult U C mC / C 52,000 / 45.3 to read for low values of UD/UC . Instead, use Eqs. (8-62) and (8-59). (1) From Eq. (8-62), with UC/UD = 1/0.289 = 3.46, Df = 1+ 8 UC / U D 1/ 2 -3 4 UC / U D - 1 = 1 + 8 3.46 1/ 2 -3 4 3.46 - 1 = 0.239 From Eq. (8-59), using D = (D)f = 0.239, UD UC UD UC + = + = uo 1 - D = 0.177(1 - 0.239) D 1 - D 0.239 1 - 0.239 which simplifies to, 4.184U D + 1314U C = 0.1347 . (2) Solving Eqs. (1) and (2) simultaneously, UC = 0.0534 ft/s and UD = 0.0154 ft/s Therefore, at flooding, (UD + UC) = 0.0154 + 0.0534 = 0.0688 ft/s Size the column for 50% of flooding. Therefore, (UD + UC) = 0.0688/2 = 0.0344 ft/s Total volumetric flow rate = Q = 21,000/63.5 + 52,000/45.3 = 1,480 ft3/h = 0.411 ft3/s Column cross-sectional area = AC = Q/(UD + UC) = 0.411/0.0344 = 11.9 ft2 Column diameter = DC = (4AC/)1/2 = [4(11.9)/3.14]1/2 = 3.9 ft Exercise 8.35 Subject: Diameter of a Karr column for extraction of benzoic acid (A) from water (C) by toluene (S). Given: Conditions in Exercise 8.33. Find: Column diameter. Analysis: Volumetric flow rate of feed/raffinate = 500 gal/min Volumetric flow rate of solvent/extract = 750 gal/min Extract is dispersed. The following properties are given in Exercise 8.30. R = D = 0.860(62.4) = 53.7 lb/ft3 and E = C = 0.995(62.4) = 62.1 lb/ft3 C = 0.95 cP or 0.000020 lbf-s/ft2 Interfacial tension = I = 22 dyne/cm or 0.0015 lbf/ft u Assume o C C = 0.01 i (0.0015)(62.1 - 53.7) = 0.01 = 0101 ft/s . Therefore, uo = 0.01 i C C (0.000020)(62.1) U D QD 750 = = = 150 . U C QC 500 From Fig. 8.39, U D + UC uo f = 0.27 Therefore, (UD +UC)f = 0.27 uo = 0.27(0.101) = 0.0273 ft/s Size the column for 50% of flooding. Therefore, (UD + UC) = 0.0273/2 = 0.0137 ft/s Total volumetric flow rate = Q =(500 + 750)(60)/7.48 = 10,030 ft3/h = 2.79 ft3/s Column cross-sectional area = AC = Q/(UD + UC) = 2.79/0.0137 = 204 ft2 Column diameter = DC = (4AC/)1/2 = [4(204)/3.14]1/2 = 16.1 ft From Table 8.2, maximum Karr diameter is 1.5 m or 4.9 ft. Therefore, need multiple Karr D 2 314(4.9) 2 . columns in parallel of maximum cross-sectional area = Ac max = = = 18.8 ft2 4 4 Therefore, need 204/18.8 = 11 units in parallel of 1.5 m diameter each. Would probably be better to specify one RDC column of 16.1 ft diameter or 5 m , which is less than the maximum diameter of 8 m. Exercise 8.36 Subject: HETS of an RDC column for extraction of acetic acid (A) from water (C) by isopropyl ether (S). Given: Conditions in Exercises 8.28, 8.30, and 8.34. Find: HETS. Analysis: From Fig. 8.40, for an interfacial tension of 13.5 dyne/c from Exercise 8.30, HETS = 4 with HETS and DT in inches. 1 DT/ 3 From Exercise 8.34, for the conditions of Exercise 8.28, DT = 2.8 ft or 33.6 inches. Therefore, HETS = 4(33.6)1/3 = 12.9 inches. From Exercise 8.34, for the conditions of Exercise 8.30, DT = 3.9 ft or 46.8 inches. Therefore, HETS = 4(46.8)1/3 = 14.4 inches. Exercise 8.37 Subject: HETS of a Karr column for extraction of benzoic acid (A) from water (C) by toluene (S). Given: Conditions in Exercises 8.32 and 8.35. Find: HETS. Analysis: From Fig. 8.40, for an interfacial tension of 22 dyne/c from Exercise 8.33, HETS = 5 with HETS and DT in inches. 1 DT/ 3 From Exercise 8.35, DT = 4.9 ft or 58.8 inches. Therefore, HETS = 5(58.8)1/3 = 19.4 inches. Exercise 9.1 Subject: Selection of type condenser and operating pressure for distillation of propionic and n-butyric acids. Given: Distillate of 95 mol% propionic acid. Bottoms of 98 mol% n-butyric acid. Assumptions: Ideal solutions so that Raoult's K-values apply. Find: type condenser and operating pressure Analysis: From Perry's Handbook or other source, the normal boiling points are: Propionic acid n-Butyric acid 141.1oC 163.5oC Using Fig. 7.16, an operating pressure of 30 psia or less, with a total condenser, might be used. Could also consider using subcooled reflux with a top pressure of just above atmospheric pressure. Exercise 9.2 Subject: Selection of type condenser and operating pressure for distillation of a mixture of light paraffin hydrocarbons in two columns Given: Feed and product compositions as follows: kmol/h: Component Feed Product 1 C1 160 160 C2 370 365 C3 240 5 nC4 25 0 nC5 5 0 Total: 800 530 Find: Type condensers and operating pressures for: (a) Direct sequence. (b) Indirect sequence. Analysis: Use Chemcad or other simulator to make the bubble-point and dew-point calculations, using the SRK equation of state. The results should be close to hand calculations using K-values from Figs. 2.8 and 2.9. Use Fig. 7.16 to determine pressure and type condenser. Column 1 of direct sequence: Distillate is Product 1. Can not condense at 120oF because this temperature is above the critical temperatures of methane and ethane. At 415 psia, the dew point temperature is 16oF. Therefore, use a partial condenser with a refrigerant boiling at a temperature of about 6oF. One choice of refrigerant is propane. Bottoms temperature would not be in the decomposition range. Column 2 of direct sequence: Distillate is Product 2, which is mainly propane. Bubble-point pressure at 120oF is 256 psia, which is greater than 215 psia. Dew-point pressure at 120oF is 248 psia. Use partial condenser with about 250 psia pressure at the top. Use cooling water in the condenser. Bottoms conditions present no problems. Column 1 of indirect sequence: Distillate is a mixture of Products 1 and 2. At 120oF, pressure is very high, even for a partial condenser. Dew-point temperature at 415 psia is 83oF. Raise pressure a little and use cooling water in the condenser. Bottoms conditions present no problems. Column 2 of indirect sequence: Conditions at the top are the same as Column 1 of the direct sequence. Therefore, use a partial condenser with a refrigerant for removing heat at 16oF. Product 2 0 5 230 1 0 236 Product 3 0 0 5 24 5 34 Exercise 9.3 Subject: Selection of type condenser and operating pressure for the distillation of a mixture of methane, ethane, benzene, and toluene by the direct sequence in two columns. Given: Feed and product conditions as follows from Fig. 9.20: kmol/h: Component Feed Product 1 Product 2 Methane 20 20.000 0.000 Ethane 5 4.995 0.005 Benzene 500 5.000 485.000 Toluene 100 0.000 0.500 Total: 625 29.995 485.505 Find: Type condenser and operating pressure in each column. Analysis: Use Chemcad or other simulator to make the bubble-point and dew-point calculations, with the SRK equation of state. Use Fig. 7.16 to determine pressure and type condenser. Column 1: Distillate is Product 1. Bubble point on Product 1 at 120oF gives an impossible very high pressure because the critical temperatures of methane and ethane are less than 120oF. A dew point on Product 1 at 120oF gives 31.4 psia because of the presence of benzene. Therefore, can use a partial condenser with cooling water. No problems occur at the bottom. Column 2: Distillate is Product 2. Bubble point on Product at 120oF gives 5.1 psia. No need to operate at a vacuum. Raise condenser pressure to 30 psia or just above atmospheric pressure and use a total condenser, possibly with subcooled reflux. No problems at the column bottom. Product 3 0.0 0.0 10.0 99.5 109.5 Subject: Exercise 9.4 Number of equilibrium stages for a deethanizer. Given: Deethanizer with feed and average relative volatilities as follows from Fig. 9.21: Component C1 C2 C3 nC4 nC5 Feed, kmol/h Avg. Relative volatility 160 8.22 370 2.42 240 1.00 25 0.378 5 0.150 Distillate is to contain 2 kmol/h of C3. Bottoms is to contain 2 kmol/h of C2. Therefore, LK is C2 and HK is C3. Assumptions: Relative volatilities are relative to C3. Find: Number of equilibrium stages for 2.5 times minimum number of stages. Analysis: To compute the minimum number of stages, use component flow rate form of the Fenske equation given by (9-12): log N min = d C2 dC3 bC3 bC 2 = log log C 2 ,C3 370 - 2 240 - 2 2 2 log 2.42 = 11.3 N = 2.5Nmin = 2.5(11.3) = 28.3 Exercise 9.5 Subject: Determination of the minimum number of stages by the Fenske equation for each of three sections in a distillation column with a vapor sidestream Given: The following feed and product compositions for a distillation column: kmol/h: Component Feed Distillate Vapor sidestream Bottoms Benzene, B 260 257.0 3.0 0.0 Toluene, T 80 0.1 79.4 0.5 Biphenyl, BP 5 0.0 0.2 4.8 Pressure, kPa 165 130 180 200 Assumptions: Raoult's law for K-values. Find: Using the Fenske equation, the minimum number of stages for sections between: (a) Distillate and feed. (b) Feed and vapor sidestream. (c) Sidestream and bottoms. Analysis: (a) For this section, take the LK as benzene and the HK as toluene. At the top, have almost pure benzene at 130 kPa = 19 psia. Therefore, from Fig. 2.4, the distillate temperature = boiling point of benzene at 19 psia = 195oF. At the feed point, the pressure = 165 psia = 23.9 psia. Assume a saturated liquid feed. The bubble-point temperature is estimated to be 215o F. Take an average temperature = 205oF, at which the vapor pressures for benzene and toluene, respectively, are 23 psia and 9 psia respectively. From Eq. (2-44), the corresponding K-values at the average section pressure of 21.5 psia are 1.07 and 0.42. The relative volatility, from Eq. (221) = 1.07/0.42 = 2.55. From the Fenske equation (9-12), applied between the distillate and feed compositions: dB fT 257 80 log log dT fB 0.1 260 N min = = = 7.1 stages log B,T log 2.55 (b) For the section between the feed and the vapor sidestream, again take the LK as benzene and the HK as toluene. The pressure at the vapor sidestream = 180 kPa = 26.1 psia. A dew-point on the vapor sidestream gives a temperature of 265oF. Assume an average temperature for the section of 240oF, at which the vapor pressures for benzene and toluene, respectively, are 37 psia and 16 psia respectively. From Eq. (2-44), the corresponding K-values at the average section pressure of 25 psia are 1.48 and 0.64. The relative volatility, from Eq. (2-21) = 1.48/0.64 = 2.31. Analysis: (b) (continued) Exercise 9.5 (continued) From the Fenske equation (9-12), applied between the feed and vapor sidestream compositions: log N min = fB fT vsT vsB log = 260 79.4 80 3 log 2.31 log B,T = 5.3 stages (c) In the section between the vapor sidestream and bottoms, take toluene as the LK and biphenyl as the HK. At the bottom of the column, the pressure = 200 kPa = 29 psia. Biphenyl boils at 255oC = 491oF at 14.7 psia, which is much higher than toluene. From Perry's Handbook, the following vapor pressure data are obtained for biphenyl: T, oF Vapor pressure, psia 243 0.19 274 0.39 307 0.77 165.2 1.16 329 1.93 400 3.87 445 7.74 A bubble point on the bottoms pressure gives 465oF. At this temperature, the relative volatility, T,BP = vapor pressure of T/vapor pressure of BP = 220/10.6 = 20.8. At the temperature of the vapor sidestream, 265oF, T,BP = vapor pressure of T/vapor pressure of BP = 27.9/0.37 = 75.4. Use a geometric mean for this wide range. Thus, T,BP = [20.8(75.4)]1/2 = 40. From the Fenske equation (9-12), applied between the vapor sidestream and bottoms compositions: vsT bBP 79.4 4.8 log log vsBP bT 0.2 0.5 N min = = = 2.3 stages log B,T log 40 Exercise 9.6 Subject: Comparison of minimum number of stages by Fenske equation and McCabeThiele method for a non-ideal system of acetone (A) and water (W). Given: A feed of 25 mol% A and 75 mol% W. Distillation at 130 kPa (975 torr) to obtain a distillate of 95 mol% A and a bottoms of 2 mol% A. Infinite-dilution liquid-phase activity coefficients of 8.12 for acetone and 4.13 for water. Find: Minimum stages by the Fenske equation and by the McCabe-Thiele method. Analysis: Because the pressure is close to 1 atm, estimate vapor-liquid equilibria from the modified Raoult's law, Eq. (4) in Table 2.3. Obtain vapor pressures from Antoine equations and liquid-phase activity coefficients over the entire composition range from fitting the infinitedilution coefficients to the van Laar equation, Eq. (3) in Table 2.9, using Eqs. (2-75). Thus, KA = yA A PAs = xA P AAW x A 1+ A WA xW AAW (1) and 2 KW = s yW W PW = xW P (2) AWA 2 A = exp (3) and W = exp x A 1 + W AW xA AWA (4) AAW = ln A = ln(8.12) = 2.094 AWA = ln = ln(4.13) = 1418 . W Substituting these van Laar coefficients into Eqs. (3) and (4), 2.094 1418 . A = exp (5) and W = exp 2 1 - xA xA 1 + 0.6772 1+1.477 xA 1 - xA From the 7th edition of Perry's Chemical Engineers' Handbook, page 13-21, 1210.595 1730.630 s log PAs = 7.11714 - (7) and log PW = 8.07131 - (8) T + 229.664 T + 233.426 where Pi s = vapor pressure of component i in torr and T = o C 2 (6) Minimum stages by the Fenske equation: First determine the relative volatilities at the top and bottom and take the geometric average for use in the Fenske equation. Exercise 9.6 (continued) Analysis: Fenske equation (continued) Bubble point for the distillate composition. Ps From Eqs. (4-12), (1), and (2): x Di Ki = x Di i i =1 (9) P i i In the distillate, xA = 0.95 and xW = 0.05. From Eqs. (3) and (4), 2.094 1418 . A = exp = 1.003 and W = exp 2 0.95 2.094 0.05 1418 . 1+ 1+ 0.05 1418 . 0.95 2.094 Eq. (9) becomes: 0.95 2 = 3.75 1003 PAs . 3.75 PAs + 0.05 = 1 where vapor pressure is in torr. 975 975 Using Eqs. (7) and (8) with a spreadsheet, a trial and error calculation gives a distillate temperature of 64oC. Then from Eqs. (1) and (2), A,W = KA/KW = 1.48 Bubble point for the bottoms composition. In the bottoms, xA = 0.02 and xW = 0.98. From Eqs. (3) and (4), 2.094 1418 . A = exp = 7.19 and W = exp = 1.001 2 2 0.02 2.094 0.98 1418 . 1+ 1+ 0.98 1.418 0.02 2.094 7.19 PAs 1001PAs . + 0.98 = 1 where vapor pressure is in torr. 975 975 Using Eqs. (7) and (8) with a spreadsheet, a trial and error calculation gives a distillate temperature of 95oC. Then from Eqs. (1) and (2), A,W = KA/KW = 27.7 The geometric mean relative volatility = [(1.48)(27.7)]1/2 = 6.40 Eq. (9) becomes: 0.02 log From the Fenske equation (9-11), N min = xDA xBA xBW xDW = log log A,W 0.95 0.98 0.02 0.05 log 6.40 = 3.7 Minimum stages by the McCabe-Thiele method: To obtain a y-x equilibrium curve at 975 torr in terms of acetone mole fractions, we can run bubble-point temperature calculations, as above, for a set of points in the range of liquid compositions between the distillate and bottoms compositions. For each point, the K-value and vapor mole fraction for acetone are computed from Eq. (1). The results from a spreadsheet are as follows: Analysis: McCabe-Thiele method (continued) T, oC yA xA ------------------------------------63.6 1.000 1.000 64.0 0.965 0.950 64.5 0.936 0.900 65.0 0.911 0.850 65.5 0.889 0.800 66.0 0.870 0.750 66.5 0.855 0.700 67.0 0.842 0.650 67.5 0.833 0.600 67.9 0.824 0.550 68.3 0.817 0.500 68.6 0.811 0.450 68.9 0.806 0.400 69.2 0.801 0.350 69.6 0.796 0.300 70.2 0.789 0.250 71.2 0.777 0.200 73.1 0.754 0.150 76.8 0.705 0.100 84.9 0.578 0.050 95.1 0.361 0.020 107.1 0.000 0.000 Exercise 9.6 (continued) The McCabe-Thiele plot for these data are shown on the next page. The equilibrium stages for total reflux are stepped off between the equilibrium curve and the 45oline, which represents the operating line, from the acetone xD = 0.95 down to the acetone xB = 0.02. From the plot, there are 5 minimum stages, which is 35% higher than the 3.7 minimum stages computed from the Fenske equation. The Fenske equation is not reliable for highly nonideal binary systems. Exercise 9.6 (continued) Analysis: McCabe-Thiele method (continued) Exercise 9.7 Subject: Minimum equilibrium stages and distribution of nonkey components using the Fenske equation for distillation of a paraffin hydrocarbon mixture. Given: Column pressure of 700 kPa. Feed composition and split for two key components in Fig. 9.23. K-values from Figs. 2.8 and 2.9. Find: Minimum number of equilibrium stages. Distribution of nonkey components at total reflux. Analysis: To apply the Fenske equation, the geometric mean relative volatility between the distillate and bottoms is needed. From the following crude estimate of the split of all nonkey components, compute the bubble-point temperatures of the distillate and bottoms and use the Kvalues at those temperatures to get the relative volatilities. Use the bubble-point equation, Eq. (4-12). The following results are obtained for a distillate bubble point of 68oF (20oC) and a bottoms bubble point of 257oF (125oC): Component Propane Isobutane n-Butane Isopentane n-Pentane n-Hexane n-Heptane n-Octane Total Feed, kmol/h 2500 400 600 100 200 40 50 40 3930 Distillate, kmol/h 2500 399 594 15 5 0 0 0 3513 K-value, distillate 1.250 0.490 0.340 0.138 0.108 0.038 0.013 0.005 KxD 0.8895 0.0557 0.0575 0.0006 0.0002 1.0035 Bottoms, kmol/h 0 1 6 85 195 40 50 40 417 K-value, bottoms 5.70 3.25 2.65 1.45 1.23 0.65 0.33 0.19 KxB 0.0078 0.0382 0.2955 0.5751 0.0623 0.0222 0.0177 1.0188 For the two key components, nC4 and iC5, the relative volatilities at the top and bottom are, respectively (0.34/0.138) = 2.46 and (2.65/1.45) = 1.83. The geometric mean nC4,iC5 = [(2.46)(1.83)]1/2 = 2.12. From the Fenske equation, (9-12), log N min = d nC4 d iC5 biC5 bnC4 = log log nC4 ,iC5 594 85 6 15 log 2.12 = 8.4 To compute the distribution of nonkey components at total reflux, use for the lighter than light key components, LLK, Eq. (9-15), and use Eq. (9-16) for the heavier than heavy key, HHK, with iC5 as the reference component, r, and the above value of Nmin. Analysis: (continued) Exercise 9.7 (continued) fi Thus, for the LLK, i, bi = 1+ d iC5 biC5 d iC5 biC5 d iC5 biC5 = N min i,iC5 fi 15 8.4 1+ i,iC5 85 (1) fi For the HHK, i, iN min ,iC 5 iN min ,iC5 di = 1+ 15 8.4 i ,iC5 85 = 15 8.5 1+ i ,iC5 85 fi (2) Using the above K-values to compute geometric mean values of i, iC5 , followed by use of Eqs. (1) and (2) with the material balance, fi = di + bi , the following results are obtained: Component Propane Isobutane n-Butane Isopentane n-Pentane n-Hexane n-Heptane n-Octane i, iC5 at top 9.06 3.55 1.00 0.783 0.272 0.094 0.038 i, iC5 at bottom i, iC5 average 3.93 5.97 2.24 2.82 1.00 0.848 0.448 0.228 0.128 0.815 0.349 0.146 0.070 di , kmol/h 2499.996 399.63 594 15 6.12 0.001 8.2x10-7 1.92x10-10 bi , kmol/h 4.2x10-3 0.37 6 85 193.88 39.999 50 40 Exercise 9.8 Subject: Type condenser and operating pressure. Minimum equilibrium stages and distribution of nonkey components using the Fenske equation for distillation of a light hydrocarbon mixture. Given: Feed composition and split of two key components in Fig. 9.24. K-values from Figs. 2.8 and 2.9. Find: Type condenser. Operating pressure. Minimum equilibrium stages. Distribution of nonkey components at total reflux. Analysis: To determine the type condenser and operating pressure, follow the algorithm in Fig. 7.16. The estimated distillate and bottoms compositions from the given data are: Component Methane Ethylene Ethane Propylene Propane n-Butane Feed, lbmol/h 1000 2500 2000 200 100 50 Distillate, lbmol/h 1000 2500 1999 5 0 0 Bottoms, lbmol/h 0 0 1 195 100 50 Using Fig. 2.8, it is quickly shown that at 120oF, both the bubble-point and the dew-point pressures for the estimated distillate composition are greater than 1000 psia because the K-value for ethane is greater than 1.0 at 120oF and 1000 psia. Therefore, select the operating pressure as 415 psia and use a partial condenser. Run a dew point on the distillate at 415 psia. Use Eq. (4-13) in the following form: di 1000 2500 1999 5 =D= + + + = 5504 Ki K C1 KC = KC 2 KC = 2 3 (1) i By trial and error, using Fig. 2.9 for K-values, the temperature that satisfies Eq. (1) is 0oF, which gives: di 1000 2500 1999 5 =D= + + + = 5505 5.0 105 0.69 018 . . i Ki The relative volatility of the two key components is C2, C3= = 0.69/0.18 = 3.83 Analysis: (continued) Now run a bubble-point on the estimated bottoms composition at 415 psia. Use Eq. (4-12) in the following form: bi Ki = B = KC 2 + 195KC= + 100 KC3 + 50 KnC4 = 346 (2) By trial and error, using Fig. 2.8 for K-values, the temperature that satisfies Eq. (1) is 155oF, . . which gives: bi Ki = B = 2.7 + 195(115) + 100(101) + 50(0.39) = 347 i i 3 Exercise 9.8 (continued) The relative volatility of the two key components is C2, C3= =2.7/1.15 = 2.35 The geometric mean relative volatility = [(3.83)(2.35)]1/2 = 3.0 From the Fenske equation, (9-12), log N min = d C2 d C= 3 bC= 3 bC2 = 2 , C3 log = log C 1999 195 5 1 log 3.0 = 10. 3 To compute the distribution of nonkey components at total reflux, use for the lighter than light key components, LLK, Eq. (9-15), and use Eq. (9-16) for the heavier than heavy key, HHK, with C= as the reference component, r, and the above value of Nmin. 3 fi fi = Thus, for the LLK, i, bi = (3) 5 d C= N 10.3 3 min 1+ i,C= 1 + 195 i,C= 3 3 b C= 3 fi For the HHK, i, d C= 3 di = 1+ bC = 3 iN min ,C = 3 d C= 3 bC= 3 iN min ,C = 3 5 10.3 = 195 i ,C3 = 5 1+ 10.3 = 195 i ,C3 fi (4) Using the above K-values to compute geometric mean values of i, C3= , followed by use of Eqs. (3) and (4) with the material balance, fi = di + bi , the following results are obtained: Component Methane Ethylene Ethane Propylene Propane n-Butane i, C3= at top 27.8 5.83 3.83 1.00 0.78 0.18 i, C3= at bottom 8.70 3.22 2.35 1.00 0.88 0.34 i, C3= average 15.55 4.33 3.00 1.00 0.83 0.25 di , lbmol/h 1000 2499.973 1999 5 0.375 8.1x10-7 bi , lbmol/h 2.1x10-8 0.027 1 195 99.625 50 Exercise 9.9 Subject: Recovery of a key component as a function of distillate flow rate for the distillation of a paraffin hydrocarbon mixture when the minimum number of equilibrium stages is fixed. Given: Feed composition and flow rate of 1000 lbmol/h. Column pressure of 250 psia. Minimum equilibrium stages = 15. K-values from Figs. 2.8 and 2.9. Find: Percent recovery of propane in the distillate as a function of distillate flow rate. Analysis: Assume that the average relative volatility between propane and n-butane is not sensitive to the separation. If a perfect separation were made between propane and n-butane at 250 psia, the distillate temperature would be approximately 110oF and the bottoms temperature would be approximately 270oF. Assume an average temperature of 190oF. From Fig. 2.8, the Kvalues and relative volatilities at 250 psia and 190oF, with the light key as propane and the heavy key as n-butane are: Component Ethane Propane n-Butane n-Pentane n-Hexane Total: Feed rate, lbmol/h 30 200 370 350 50 1000 K-value 4.3 1.74 0.72 0.30 0.133 referred to nC4 5.97 2.42 1.00 0.417 0.185 Based on these relative volatility values, assume that little of the nonkey components will distribute. Therefore, D = d C2 + d C3 + d nC 4 + d nC5 + d nC6 = 30 + d C 3 + d nC 4 (1) The Fenske equation, (9-12) becomes: log N min = 15 = d C3 bnC4 d nC4 bC 3 avg log = dC3 d nC 4 370 - d nC4 200 - d C3 log C3 , nC4 200 - d C 3 d C3 log 2.42 200 - d C 3 d C3 Rearranging, 370 - d nC4 d nC4 = 2.4215 = 571800 (2) Assume a value of d for C3. Calculate d for nC4 from Eq. (2). Then compute D from Eq. (1). Then compute recovery of C3 = d for C3/200 and make plot. Exercise 9.9 (continued) Analysis: (continued) d of C3, lbmol/h % recovery of C3 d of nC4, lbmol/h 198 99 0.064 190 95 0.012 180 90 0.006 150 75 0.002 100 50 0.00065 D, lbmol/h 228.1 220.0 210.0 180.0 130.0 We see that with 15 minimum stages, the distribution of nC4 to the distillate is almost negligible. Exercise 9.10 Subject: Minimum reflux ratio by the Underwood equation for the separation of a binary mixture as a function of feed vaporization. Given: Binary feed of 30 mol% propane and 70 mol% propylene. Distillate to contain 99 mol% propylene, and bottoms to contain 98 mol% propane. Column pressure of 300 psia. Kvalues in Figs. 2.8 and 2.9. Find: Minimum reflux ratio by the Underwood equation for: (a) Bubble-point liquid feed. (b) Feed of 50 mol% vapor. (c) Dew-point vapor feed. Analysis: Because we have a binary mixture, this is a Class I separation, where all components distribute. Therefore, the Underwood equation that applies is (9-20): xC = , D 3 Lmin = D xC = , 3 - C = ,C 3 3 3 xC 3 3 ,D 3 , xC C = ,C - 1 = 0.99 0.01 - C = ,C 3 3 xC = , xC , 3 3 C = ,C - 1 3 3 (1) (a) For a bubble-point liquid feed, the liquid-phase mole fractions in the pinch zone are those of the total feed. Thus, xC= , = 0.70 and xC , = 0.30. The relative volatility is that at the pinch. A 3 bubble-point temperature on the feed composition at 300 psia gives 126oF, with = 1.035/0.92 = 1.125. Substitution in Eq. (1) gives Lmin/D = 11.0 (b) For 50 mol% vaporization, a flash gives about the same temperature of 126oF, and therefore the same , with liquid-phase mole fractions of 0.688 for propylene and 0.312 for propane. Substitution into Eq. (1) gives Lmin/D =11.2. (c) For a dew-point feed, the temperature still is 126oF, with liquid-phase mole fractions of 0.68 for propylene and 0.32 for propane. Substitution into Eq. (1) gives Lmin/D =11.4 Because the relative volatility is close to 1, the % vaporization of the feed has only a small effect on the minimum reflux ratio. 3 Exercise 9.11 Subject: Minimum reflux rate and distribution of nonkey components for the distillation of a paraffin hydrocarbon mixture. Given: Column pressure of 700 kPa. Feed composition and split for two key components in Fig. 9.23. Feed is a bubble-point liquid. K-values from Figs. 2.8 and 2.9. Find: Minimum external reflux rate. Distribution of nonkey components at minimum reflux ratio. Analysis: Because of the wide range of volatility of the components in the feed, and the relative sharpness of the separation between the LK, n-butane, and the HK, isopentane, assume a Class 2 separation for estimating the minimum reflux ratio by the method of Underwood. Compute the minimum reflux using relative volatilities at the feed temperature. Because the feed is at the bubble point, use Eq. (4-12). For the pressure of 700 kPa (102 psia), a trial-and-error calculation for the bubble-point temperature gives the following results for 80oF, where nC8 is by extrapolation: Component Propane iso-Butane n-Butane iso-Pentane n-Pentane n-Hexane n-Heptane n-Octane Total: z, feed mole fraction 0.6361 0.1018 0.1527 0.0254 0.0509 0.0102 0.0127 0.0102 1.0000 K at 102 psia and 800F 1.37 0.56 0.39 0.16 0.12 0.038 0.013 0.004 Kz 0.8715 0.0570 0.0596 0.0041 0.0061 0.0004 0.0002 0.0000 0.9989 referred to iso-pentane 8.56 3.50 2.44 1.00 0.75 0.24 0.081 0.025 The Underwood equations that apply are Eqs. (9-28) and (9-29). For a bubble-point feed, 1 - q = 0 and Eq. (9-28) becomes: i ,r zi , F i i ,r - = 1- q = 0 = 8.56(0.6361) 35(01018) 2.44(01527) 10(0.0254) . . . . + + + + 8.56 - 3.5 - 2.44 - 10 - . 0.75(0.0509) 0.24(0.0102) 0.081(0.0127) 0.025(0.0102) + + + 0.75 - 0.24 - 0.081 - 0.025 - (1) Eq. (1) has 8 roots for . However, only 3 roots are probably necessary. These are those between values of for: (1) iso-butane and n-butane, (2) n-butane and iso-pentane, and (3) iso- Analysis: (continued) Exercise 9.11 (continued) pentane and n-pentane, because iso-butane and n-pentane boil closely to n-butane and isopentane, respectively. Using a spreadsheet, the 3 roots of are: (1) 2.7159 between 3.50 and 2.44, (2) 1.02575 between 2.44 and 1.00, and (3) 0.78274 between 1.00 and 0.75. Eq. (9-29) is applied in the following form for each of the three values of , in terms of the three unknowns, diC4 , dnC5 , and Lmin. The form used is obtained by multiplying Eq. (9-29) by the total distillate rate, D, which converts distillate mole fractions to component distillate rates, and the reflux ratio, R to L. 3.50 (d iC4 ) 0.75 (d nC5 ) 8.56 (2500) 2.44 (594) 1.00 (15) D + Lmin = + + + + 8.56 - 2.7159 350 - 2.7159 2.44 - 2.7159 1.00 - 2.7159 0.75 - 2.7159 . = 36618 + 4.4637(d iC 4 ) - 5253.2 - 8.7418 - 0.3815(d nC5 ) = 4.4637(d iC4 ) - 1600.14 - 0.3815(d nC5 ) . D + Lmin = 0.75 (d nC5 ) 3.50 (d iC4 ) 1.00 (15) 8.56 (2500) 2.44 (594) + + + + 8.56 - 102575 350 - 102575 2.44 - 1.02575 100 - 102575 0.75 - 102575 . . . . . . = 2840.36 + 141457(d iC 4 ) + 1024.83 - 582.52 - 2.71986(d nC5 ) = 141457(d iC4 ) + 3282.67 - 2.71986(d nC5 ) . . D + Lmin = 3.50 (d iC4 ) 0.75 (d nC5 ) 8.56 (2500) 2.44 (594) 1.00 (15) + + + + 8.56 - 0.78274 3.50 - 0.78274 2.44 - 0.78274 100 - 0.78274 0.75 - 0.78274 . = 2751.61 + 1.28806(d iC 4 ) + 874.55 + 69.04 - 22.9078(d nC5 ) = 128806(d iC4 ) + 3695.20 - 22.9078(d nC5 ) . Because the total distillate rate, D, is also an unknown, the following equation is also needed: D = dC3 + diC4 + dnC4 + diC5 + dnC5 = 2500 + diC4 +594 + 15 + dnC5 = 3109 + diC4 + dnC5 Solving these 4 linear equations, the following results are obtained: D = 4712.8 kmol/h, Lmin = 795.4 kmol/h, diC4 = 1593.4 kmol/h, dnC5 = 10.45 kmol/h Because the flow rate of iC4 in the distillate is greater than its feed rate, iC4 does not distribute. Recalculate omitting the first of the four equations that uses = 2.7159. Replace diC4 = fiC4 = 400 kmol/h. The equations to be solved now are: D + Lmin + 2.71986 d nC5 = 3848.5 D + Lmin + 22.9078 d nC5 = 4210.4 D - d nC5 = 3509.0 Analysis: (continued) Exercise 9.11 (continued) Solving, D = 3526.9 kmol/h, Lmin = 272.8 kmol/h, and dnC5 = 17.9 kmol/h Because the distillate rate of nC5 is positive and less than its feed rate, it does distribute. If we had assumed that only the two key components distributed, then it is only necessary to solve the following two equations using only = 1.02575, the one between the two key components, with dnC5 = 0.0 : D + Lmin = 3848.5 D = 3509.0 Then, Lmin = 3848.5 - 3509.0 = 339.5 kmol/h = internal reflux rate, which is high and incorrect. Assume that the minimum external reflux rate = minimum internal reflux rate = 272.8 kmol/h The product compositions at minimum reflux rate are: Component Propane iso-Butane n-Butane iso-Pentane n-Pentane n-Hexane n-Heptane n-Octane Total: : Feed, kmol/h 2500 400 600 100 200 40 50 40 3930 Distillate, kmol/h 2500 400 594 15 17.9 0 0 0 3526.9 Bottoms, kmol/h 0 0 6 85 182.1 40 50 40 403.1 Exercise 9.12 Subject: Minimum reflux ratio and minimum number of equilibrium stages as a function of product purity for the distillation of a binary mixture. Given: Equimolar bubble-point feed of isobutane and n-butane at column pressure of 100 psia. Purity of iC4 in the distillate equal to purity of nC4 in the bottoms. Purity range from 90 mol% to 99.99 mol% Find: Minimum reflux ratio. Minimum number of equilibrium stages. Discuss significance of results. Analysis: From the vapor pressure plot of Fig. 2.4, for 100 psia, the boiling points are 120oF for iC4 and 146oF for nC4. Take an average temperature of 133oF and assume a relative volatility equal to the vapor pressure ratio. Thus, again using Fig. 2.4, iC4, nC4 = 115/85 = 1.35. For Nmin , use the Fenske equation, (9-11), which simplifies to: log N min = xiC 4 xiC 4 2 D 2 B log = xiC 4 01303 . 2 D 1 - xiC 4 2 D log 135 . (1) For Rmin , use the Class 1 form of the Underwood equation, (9-20), which simplifies to: Rmin = xiC 4 , D / 0.5 - 135 1 - xiC4 , D / 0.5 . 0.35 (2) Using a spreadsheet, the following values are computed: % Purity 90 92 94 95 96 98 99 99.5 99.9 99.95 99.99 Nmin 14.6 16.3 18.3 19.6 21.2 25.9 30.6 35.3 46.0 50.7 61.4 Rmin 4.37 4.64 4.91 5.04 5.18 5.45 5.58 5.65 5.70 5.71 5.71 Exercise 9.12 (continued) Analysis: (continued) A plot of the results is shown below. It is clearly evident from the table above or the plot below that the minimum reflux increases little and less and less as the purity approaches 100%, while the minimum number of equilibrium stages increases greatly and more and more as the purity approaches 100%, Exercise 9.13 Subject: mixture. Reflux ratio by the FUG method for the distillation of a propylene-propane Given: Binary feed of 360 kmol/h of propylene and 240 kmol/h of propane at the bubble point at column pressure. Distillate to contain 347.5 kmol/h of propylene and 3.5 kmol/h of propane. Average relative volatility = 1.11. N/Nmin = 2. Total condenser and partial reboiler. Assumptions: External reflux ratio = internal reflux ratio at the upper pinch. Find: Operating reflux ratio by the FUG method. Analysis: By material balance, the bottoms product contains 12.5 kmol/h of propylene and 236.5 kmol/h of propane. Using the Fenske equation (9-12): log N min = d C = bC 3 3 3 d C bC = 3 3 3 log = 347.5 236.5 35 . 12.5 log 111 . log C= ,C = 72.2 avg N = 2 Nmin = 2(72.2) = 144.4 For minimum reflux, use the Class 1 Underwood equation (9-20), which applies for a binary mixture, where the pinch composition is that of the liquid feed. The distillate mole fractions are 347.5/351 = 0.99 for propylene and 0.01 for propane. In the feed, the mole fractions are 360/600 = 0.60 for propylene and 0.40 for propane. Thus, xC = , D 3 Rmin = Lmin = D xC = , F 3 - C = ,C 3 3 3 xC 3 3 ,D 3 ,F xC C = ,C - 1 = 0.99 0.01 - 111 . 0.60 0.40 111 - 1 . = 14.75 Use either the Gilliland graphs of Figs. 9.10 or 9.11, or the Gilliland correlation equation, (9-34). The latter is more accurate, but requires iteration when solving for X. Apply the equation. N - N min 144.4 - 72.2 Y= = = 0.497 N +1 144.4 + 1 Using a spreadsheet to solve Eq. (9-34), which is nonlinear in X, we obtain X = 0.159 R - Rmin X= = 0.159 Solving, R = 17.73 and R/Rmin = 17.73/14.75 = 1.202 R +1 Exercise 9.14 Subject: Number of equilibrium stages by the FUG method for the distillation of a binary mixture of dichlorobenzene (DCB) isomers. Given: Feed of 62 mol% p-DCB and 38 mol% o-DCB at approximately 1 atm. Distillate to contain 98 mol% p-DCB, and bottoms to contain 96 mol% o-DCB. Total condenser and partial reboiler. Average relative volatility = 1.154 with the para isomer as the LK. Feed condition is q = 0.9 (10 mol% vaporized). Want R/Rmin = 1.15. Assumptions: External reflux ratio = internal reflux ratio at the upper pinch. Find: Number of equilibrium stages by the FUG method. Analysis: Using the Fenske equation (9-11): log N min = xp-DCB xp-DCB D B xo-DCB xo-DCB B D log = (0.98) 0.96 (0.04) 0.02 log(1154) . log p-DCB,o-DCB = 49.3 For minimum reflux, use the Class 1 Underwood equation (9-20), which applies for a binary mixture, where the pinch composition is that of the liquid portion of the feed. From the feed composition, determine the equilibrium liquid composition by a flash calculation for 10 mol% vaporization. Do this by combining Eqs. (7-26) for the q-line with (4-8) for equilibrium at a constant relative volatility. Thus, applying these equations to the LK, p-DCB, x 1154 x . q zF 0.62 = = x- = -9 x + = 6.2 - 9 x 1 + x ( - 1) 1 + 0154 x . q -1 q -1 01 . Solving this nonlinear equation for a positive root of x between 0 and 1 gives x = xp-DCB = 0.617 Then, xp-DCB, D Rmin = Lmin = D xp-DCB, F - p - DCB,o - DCB p - DCB,o - DCB - 1 xo-DCB, D xo - DCB, F 0.98 0.02 . - 1154 0.617 0.387 1154 - 1 . = = 9.93 R = 1.15 Rmin =1.15(9.93) = 11.42 In the Gilliland equation, (9-34), X = (R - Rmin)/(R + 1) = (11.42-9.93)/(11.42+1) = 0.120. Using Eq. (9.34), Y = 0.534 = (N - Nmin)/(N + 1). Solving, N = 107 stages. Exercise 9.15 Subject: Shortcomings of the Gilliland correlation Given: Column with a small ratio of rectifying to stripping stages. Find: Explanation for possible inapplicability of the Gilliland correlation. Analysis: When a column has a small ratio of rectifying to stripping stages, stripping of a light key in the stripping section may dominate over absorption of the heavy key in the rectifying section, in determining stage requirements. The Gilliland correlation is based on reflux considerations in the rectifying section. Alternatively, a correlation could have been developed based on boilup considerations in the stripping section, using a modified Underwood equation for minimum boilup ratio. Such a correlation, since it would be empirical, would seem to be favored over the Gilliland correlation, which emphasizes rectification, rather than stripping. Subject: mixture. Exercise 9.16 Use of the FUG method for the distillation of a normal paraffin hydrocarbon Given: Bubble-point liquid feed at 300 psia, with composition in mole fractions: C3 nC4 nC5 nC6 nC7 Component C2 Mole fraction 0.08 0.15 0.20 0.27 0.20 0.10 Find: (a) For a sharp separation between nC4 and nC5, determine column pressure and type condenser for condenser outlet temperature of 120oF. (b) At total reflux, determine the separation for 8 equilibrium stages, with 0.01 mole fraction of nC4 in the bottoms. (c) Minimum reflux ratio for separation in Part (b). (d) Number of equilibrium stages for R/Rmin = 1.5 using Gilliland correlation. Analysis: (a) For a sharp split, the distillate composition is as follows: Component C2 C3 nC4 Total: xd 0.186 0.349 0.465 1.000 K at 120oF, 250 psia 2.8 1.13 0.34 Kxd 0.520 0.394 0.158 1.072 In this table, a distillate bubble-point pressure of 250 psia is assumed and the bubble-point equation (4-12) is applied, using K-values in Fig. 2.8. The summation is 1.072. In order for the sum to be 1.0, an even higher pressure is needed, since the K-value is almost inversely proportional to pressure. But, in Fig. 7.16, we are already > 215 psia, therefore use a partial condenser. To determine the pressure, run a dew point using Eq. (4-13). Component C2 C3 nC4 Total: yD 0.186 0.349 0.465 1.000 K at 120oF, 140 psia 4.7 1.58 0.53 yD/K 0.040 0.221 0.877 1.138 K at 120oF, 120 psia 5.4 1.80 0.60 yD/K 0.030 0.194 0.775 0.999 On the second trial in the above table, the dew-point pressure is found to be 120 psia. Set the column pressure at this value with a partial condenser. Exercise 9.16 (continued) Analysis: (continued) (b) Run a bubble point on an assumed bottoms composition, using Eq. (4-12): xB K at 320oF, xB K K at 290oF, xB K Component 120 psia 120 psia nC4 0.010 3.3 0.033 2.75 0.028 nC5 0.469 1.7 0.797 1.38 0.647 nC6 0.347 0.92 0.319 0.72 0.250 nC7 0.174 0.48 0.084 0.36 0.063 Total: 1.000 1.233 0.988 o A third trial gives a bottoms temperature of 295 F The K-values and relative volatilities referred to the nC5, the heavy key, at distillate and bottoms conditions are as follows for a pressure of 120 psia: Component K at 120oF, I, nC5 at 120oF, K at 295oF, I, nC5 at 295oF, Geometric 120 psia 120 psia 120 psia 120 psia average C2 5.40 25.7 12.0 8.28 14.6 C3 1.80 8.6 5.8 4.0 5.9 nC4 0.60 2.86 2.8 1.93 2.35 nC5 0.21 1.00 1.45 1.00 1.00 nC6 0.074 0.352 0.74 0.51 0.424 nC7 0.028 0.133 0.38 0.262 0.187 log From the Fenske Eq. (9-12), N min = 8 = d nC4 bnC5 d nC5 bnC4 avg log = d nC 4 bnC5 d nC5 bnC4 log 2.35 (1) log nC4 ,nC5 Take as a basis, 1 mol of feed. Assume that the distribution of the nonkeys is negligible. That is, only the LK and the HK are found in both the distillate and bottoms. Assume the total bottoms, B, is 0.57 mol. Then bnC4 = 0.01(0.57) = 0.0057. Therefore, dnC4 = 0.20 - 0.0057 = 0.1943 mol. Solving Eq. (1), bnC5/dnC5 = 27.3. Since bnC5 + dnC5 = fnC5 = 0.27, bnC5 = 0.2605 mol and dnC5 = 0.0095. Therefore, B = 0.0057 + 0.2605 + 0.20 + 0.10 = 0.5662. Repeating the calculations with this value of B, the result is bnC4 = 0.00566, dnC4 = 0.19434, bnC5 = 0.2604, and dnC5 = 0.0096. Using Eq. (1) for each of the nonkey components shows that they do not distribute to any extent. Therefore the distillate and bottoms compositions at total reflux are: xD xB Component f, mol d, mol b, mol C2 0.08 0.0800 0.184 0.0000 0.000 C3 0.15 0.1500 0.346 0.0000 0.000 nC4 0.20 0.1943 0.448 0.0057 0.010 nC5 0.27 0.0096 0.022 0.2604 0.460 nC6 0.20 0.0000 0.000 0.2000 0.353 nC7 0.10 0.0000 0.000 0.1000 0.177 Total 1.00 0.4339 1.000 0.5661 1.000 Exercise 9.16 (continued) Analysis: (c) Because the split of the two key components is quite sharp, assume a Class 2 separation for the purpose of minimum reflux. Thus, only components methane to n-pentane will distribute to the distillate. To determine the phase condition of the feed, perform an adiabatic flash with bubble-point conditions at 300 psia upstream of the feed valve to the column, and 120 psia downstream of the valve, using Chemcad. The result is a V/F ratio for the feed of 0.334. From Eq. (7-19), q = 1 - (V/F) = 1 - 0.334 = 0.666 or 1 - q = 0.334. From Underwood Eq. (9-28), i ,nC5 zi , F i i ,nC5 - = 1 - q = 0.334 = 14.6(0.08) 5.9(015) 2.35(0.20) 1.00(0.27) 0.424(0.20) 0.187(010) . . + + + + + 14.6 - 015 - . 2.35 - 100 - . 0.424 - 0187 - . Solving this equation for the root between 2.35 and 1.00 gives = 1.5962 From Underwood Eq. ((9-29), i ,nC5 xi , D 14.6(0184) . 5.9(0.346) 2.35(0.448) 1.00(0.022) = + + + = 2.04 = 1 + Rmin 14.6 - 15962 5.9 - 15962 2.35 - 15962 1.00 - 15962 . . . . i i , nC5 - Therefore, the internal minimum reflux ratio = 2.04 - 1.00 = 1.04 Assume the external minimum reflux ratio = 1.04 (d) For the number of theoretical stages, use the Gilliland equation, (9-34), where for this exercise, R = 1.5R = 1.5(1.04) = 1.56. X = (R - Rmin)/(R + 1) = (1.56 - 1.04)/(1.56 + 1) = 0.203. Using Eq. (9.34), Y = 0.458 = (N - Nmin)/(N + 1). Solving, N = 15.6 equilibrium stages. Exercise 9.17 Subject: mixture. Use of the FUG method for the distillation of a light paraffin hydrocarbon Given: Feed mixture of: Component C3 iC4 nC4 iC5 nC5 lbmol/h 5 15 25 20 35 Column pressure = 120 psia. Liquid distillate contains 92.5% of the nC4 in the feed. Bottoms contains 82 mol% of iC5 in the feed. Use Chemcad with SRK instead of Figs. 2.8 and 2.9 for Kvalues. Find: (a) Minimum number of equilibrium stages. (b) Distribution of nonkey components. (c) Minimum reflux ratio for a bubble-point liquid feed. (d) Number of equilibrium stages for R = 1.2 Rmin. (e) Feed-stage location. Analysis: (a) To obtain the average relative volatilities, run bubble points on the following assumed distillate and bottoms compositions, using the given key-component recoveries: lbmol/h: Component Feed Distillate C3 5 5.000 iC4 15 15.000 nC4 25 23.125 iC5 20 3.600 nC5 35 0.000 Total: 100 46.725 Bottoms 0.000 0.000 1.875 16.400 35.000 53.275 The results with K-values and relative volatilities are as follows, using the SRK equation of state for K-values and the bubble-point equation, (4-12): Component C3 iC4 nC4 iC5 nC5 K, distillate, 140oF 2.100 1.056 0.818 0.401 0.329 i,iC5 , 140oF 5.24 2.63 2.04 1.00 0.82 K, bottoms, 230oF 3.880 2.260 1.863 1.068 0.922 i,iC5 , 230oF 3.63 2.12 1.74 1.00 0.86 i,iC5 , geometric mean 4.36 2.36 1.88 1.00 0.84 Analysis: (continued) (a) Using the Fenske equation (9-12), with nC4 as the LK and iC5 as the HK, log N min = d nC4 biC5 d iC5 bnC4 log = Exercise 9.17 (continued) ( 23.125 ) (16.4 ) ( 3.6 ) (1.875 ) log1.88 log ( nC4,iC5 )avg = 6.38 (b) To compute the distribution of nonkey components at total reflux, use for the lighter than light key components, LLK, Eq. (9-15), and use Eq. (9-16) for the heavier than heavy key, HHK, with iC5 as the reference component, r, and the above value of Nmin. fi fi = (1) Thus, for the LLK, i, bi = d iC5 N min 1 + 3.6 6.38 i,iC5 1+ i,iC5 16.4 b iC5 fi For the HHK, i, d iC5 biC5 d iC5 biC5 iN min ,iC 5 iN min ,iC5 di = 1+ 3.6 6.38 i ,iC5 16.4 = 3.6 6.38 1+ i ,iC5 16.4 fi (2) Using the above geometric mean values of i, iC5 , followed by use of Eqs. (1) and (2) with the material balance, fi = di + bi , the following results are obtained: lbmol/h: Component Feed Distillate Bottoms C3 5 4.998 0.002 iC4 15 14.720 0.280 nC4 25 23.125 1.875 iC5 20 3.600 16.400 nC5 35 2.356 32.644 Total: 100 48.799 51.201 (c) Because the split of the LK and HK components is not sharp and because iC4 boils close to nC4, and iC5 boils close to nC5, it is possible that both iC4 and nC5 will distribute at minimum reflux. It is questionable that C3 will distribute. Nevertheless, first assume a Class 1 Underwood minimum reflux and use the following rearrangement of Eq. (9-21) for a feed at the bubble point: F Lmin = d nC 4 f nC 4 - nC 4 ,iC5 nC4 ,iC5 - 1 d iC5 f iC5 = 100 23125 . 3.6 - 188 . 25 20 188 - 1 . = 66.7 l Exercise 9.17 (continued) Analyze: (c) (continued) Now, use the Class 2 Underwood equation, Eq. (9-28),which for a bubble-point feed is: i ,iC5 zi , F . . 4.36(0.05) 2.36(015) 188(0.25) 1.00(0.20) 0.84(0.35) = 1- q = 0 = + + + + . 4.36 - 2.36 - 188 - 1.00 - 0.84 - i i ,iC5 - (3) There are four roots to Eq. (3), which from a spreadsheet calculation are: = 0.9248, 1.317, 2.173, 3.996 Assume that the C3 does not distribute. Then, the 3.996 root is not needed and dC3 = 5 lbmol/h. Eq. (9-29) is applied in the following form for each of the three remaining values of , in terms of the three unknowns, diC4 , dnC5 , and Lmin. The form used is obtained by multiplying Eq. (9-29) by the total distillate rate, D, which converts distillate mole fractions to component distillate rates, and the reflux ratio, R to L. 2.36 (d iC 4 ) 1.88 (23125) 0.84 (d nC5 ) 4.36 (5) . 1.00 (3.6) D + Lmin = + + + + 4.36 - 0.9248 2.36 - 0.9248 188 - 0.9248 1.00 - 0.9248 0.84 - 0.9248 . = 6.346 + 1644(d iC4 ) + 45514 + 47.872 - 9.9057(d nC5 ) = 1.644(d iC4 ) + 93.386 - 9.9057(d nC5 ) . . D + Lmin = 0.84 (d nC5 ) 2.36 (d iC 4 ) 1.88 (23125) 1.00 (3.6) 4.36 (5) . + + + + . . . . 4.36 - 1317 2.36 - 1.317 188 - 1317 100 - 1317 0.84 - 1317 . . . . = 7.164 + 2.263(d iC 4 ) + 77.22 - 1136 - 1.761(d nC5 ) = 2.263(d iC4 ) + 73.02 - 1761(d nC5 ) D + Lmin = 2.36 (d iC4 ) 1.88 (23125) 1.00 (3.6) 0.84 (d nC5 ) 4.36 (5) . + + + + 4.36 - 2.173 2.36 - 2.173 188 - 2.173 100 - 2.173 0.84 - 2.173 . . = 9.968 + 12.62(d iC4 ) - 148.38 - 3.07 - 0.63(d nC5 ) = 12.62(d iC4 ) - 14148 - 0.63(d nC5 ) . Because the total distillate rate, D, is also an unknown, the following equation is also needed: D = dC3 + diC4 + dnC4 + diC5 + dnC5 = 5 + diC4 +23.125 + 3.6 + dnC5 = 31.725 + diC4 + dnC5 Solving these 4 linear equations, the following results are obtained: D = 53.27 lbmol/h, Lmin = 64.74 lbmol/h, diC4 = 20.61 lbmol/h, dnC5 = 0.93 lbmol/h Note that the distillate rate for iC4 is impossible because it exceeds the feed rate of 15 lbmol/h. Therefore, iC4 does not distribute, and the calculations must be repeated, eliminating diC4 as an unknown by replacing it with a value of 15 lbmol/h. Also, the root of 2.173 is not used. The following three equations result: Exercise 9.17 (continued) Analyze: (c) (continued) D + Lmin = 106.96 - 1761d nC5 . D + Lmin = 118.05 - 9.9057d nC5 D = 46.725 + d nC5 Solving these three equations: D = 48.085 lbmol/h, Lmin = 56.49 lbmol/h, and dnC5 = 1.36 lbmol/h This is the correct value of internal minimum reflux rate. Assume the external minimum reflux ratio = internal minimum reflux ratio = 56.49/48.085 = 1.175. Note that this value is about 15% less than that obtained by assuming a Class 1 separation. (d) The operating reflux ratio = 1.2(1.175) = 1.41. In the Gilliland equation, (9-34), X = (R Rmin)/(R + 1) = (1.41 - 1.175)/1.41+1) = 0.0975. Using Eq. (9.34), Y = 0.556 = (N - Nmin)/(N + 1). Solving, N = 15.6 stages. (e) Apply the Kirkbride equation (9-36): From above, using the Fenske distribution, D = 48.799 lbmol/h and B = 51.201 lbmol/h NR = NS ziC5 , F znC4 , F xnC4 , B xiC5 , D 2 0.206 B D = 0.20 0.25 0.0366 0.0738 2 51.201 48.799 0.206 = 0.723 Therefore, of 15.6 equilibrium stages, (0.723/1.723)(15.6) = 6.5 stages are in the rectifying section. Thus, the feed stage is equilibrium stage 6 or 7 from the top. Exercise 9.18 Subject: Use of the FUG method for the distillation of a chlorination effluent. Given: Bubble-point liquid feed with the following composition and average K-values: Component C2H4 , A HCl , B C2H6 , C C2H5Cl , D Mole fraction 0.05 0.05 0.10 0.80 Aver. K-value 5.10 3.80 3.40 0.15 Partial condenser and partial reboiler. Column pressure = 240 psia. (xD/xB) = 0.01 for C2H5Cl and 75 for C2H6. Find: Minimum equilibrium stages Component distribution Minimum reflux ratio Number of equilibrium stages for R = 1.5 Rmin Feed stage location Analysis: For minimum equilibrium stages, use the Fenske equation (9-11): log N min = ( xD / xB ) C H ( xD / xB )C H Cl 2 6 2 5 log C,D 75 0.01 = = 2.86 log(3.4 / 0.15) log For component distribution, use the following rearrangements of the Fenske equation: xD xB xD xB = C2 H 4 xD xB N C2min4 ,C2 H5Cl = 0.01 H C 2 H 5Cl 51 . 015 . 2 .86 = 240 = 103 HCl x = D xB C 2 H 5Cl N min HCl,C 2 H 5Cl 3.8 = 0.01 015 . 2 .86 Because the separation between the two key components is quite sharp, assume that the two nonkeys do not distribute at minimum reflux. The relative volatilities with respect to the heavy key are: 34 for A, 25.3 for B, and 22.7 for C. Now, use the Class 2 Underwood equation, Eq. (928),which for a bubble-point feed is: Analysis: (continued) i , D zi , F i Exercise 9.18 (continued) 34(0.05) 25.3(0.05) 22.7(0.10) 1.00(0.80) + + + (1) . 34 - 25.3 - 22.7 - 100 - i ,D - = 1- q = 0 = There are three roots to Eq. (1). The one of interest is the one between 22.7 and 1.0, which from a spreadsheet calculation is = 4.318. The composition of the distillate by material balances is: Component distillate, mole fraction C2H4 , A 0.25405 HCl , B 0.25405 C2H6 , C 0.48195 C2H5Cl , D 0.00995 Apply Underwood Eq. (9-29): i , D xi , D 34(0.25405) 25.3(0.25405) 22.7(0.48195) 100(0.00995) . = + + + = 1 + Rmin 34 - 4.318 25.3 - 4.318 22.7 - 4.318 1.00 - 4.318 i i ,D - = 0.291 + 0.306 + 0.595 - 0.003 = 1189 = 1 + Rmin . Therefore, Rmin = 0.189. Assume this is equal to the external minimum reflux ratio. Operating reflux ratio = 1.5(0.189) = 0.284. In the Gilliland equation, (9-34), X = (R - Rmin)/(R + 1) = (0.284 - 0.189)/0.284 + 1) = 0.074. Using Eq. (9.34), Y = 0.581 = (N - Nmin)/(N + 1). Solving, N = 8.2 stages. For feed stage location, apply the Kirkbride equation (9-36): From above, using the Fenske distribution with a material balance, D/B = 0.245. Also, mole fractions in the bottoms product are: xC = 0.0064 and xD = 0.9936 NR = NS zD , F zC , F xC , B xD , D 2 0.206 B D = 0.80 0.10 0.0064 0.00995 2 1 0.245 0.206 = 171 . Therefore, of 8.2 equilibrium stages, (1.71/2.71)(8.2) = 5.2 stages are in the rectifying section. With a partial condenser acting an equilibrium stage, the feed is about stage 4 from the top stage in the column. Exercise 9.19 Subject: Use of the FUG method for the distillation of a ficticious ternary mixture. Given: 100 kmol/h of a bubble-point feed of: Component A B C Feed rate, kmol/h 40 20 40 i,C 5 3 1 Find: (a) Product distribution for a distillate rate of 60 kmol/h and 5 minimum theoretical stages. (b) Minimum reflux and boilup ratios for separation of part (a). (c) Number of equilibrium stages and feed-stage location for R = 1.2Rmin. Analysis: (a) With a distillate rate of 60 kmol/h, the LK is B and the HK is C. From the rearranged form of the Fenske equation, given by Eq. (9-14), using component C as the reference, dA d d = C 55 = 3125 C bA bC bC dB d d = C 35 = 243 C bB bC bC The following 4 material balances also apply: dA + bA = 40 dB + bB = 20 dC + bC = 40 dA + dB + dC = 60 Solving these 6 equations, two of which are nonlinear, in 6 unknowns gives the following results: Component A B C zF 0.4 0.2 0.4 xD 0.662 0.307 0.031 xB 0.007 0.039 0.954 (b) Assuming that component A does not distribute, the Underwood Eq. (9-28) for Class 2 applies: i , D zi , F i i ,D - = 1- q = 0 = 5(0.40) 3(0.20) 1(0.40) + + 5- 3- 1- (1) Analysis: (continued) Exercise 9.19 (continued) There are two roots to Eq. (1). The one of interest is the one between 3 and 1.0, which from a spreadsheet calculation is = 1.425. If component A does not distribute, the composition of the distillate by material balances is as follows, if the total distillate rate of 60 kmol/h is maintained: kmol/h: xD xB Component Feed Distillate Bottoms A 40 40 0.0 0.6667 0.0000 B 20 18.3 1.7 0.3050 0.0425 C 40 1.7 38.3 0.0283 0.9575 Total: 100 60 40 1.0000 1.0000 Apply Underwood Eq. (9-29): i , D xi , D i i ,D - = 5(0.6667) 3(0.3050) 1(0.0283) + + = 1 + Rmin 5 - 1425 . 3 - 1.425 1 - 1.425 = 0.9324 + 0.5810 - 0.0666 = 1447 = 1 + Rmin . Therefore, Rmin = 1.447 - 1.000 = 0.447. Assume this is also the external reflux ratio. Assume constant molar overflow. Then, the reflux rate in the rectifying section = 0.447(60) = 26.8 kmol/h. The reflux rate in the stripping section = 100 + 26.8 = 126.8 kmol/h. The boilup rate in the stripping section = 126.8 - 40 = 86.8 kmol/h. The boilup ratio = 86.8/40 = 2.17. (c) Operating reflux ratio = 1.2(0.0.447) = 0.536. In the Gilliland equation, (9-34), X = (R - Rmin)/(R + 1) = (0.536 - 0.447)/0.536 + 1) = 0.058. Using Eq. (9.34), Y = 0.60 = (N - Nmin)/(N + 1). Solving, N = 14 stages. For feed stage location, apply the Kirkbride equation (9-36): From above, using the Fenske distribution with a material balance, D/B = 1.5. NR = NS zC , F zB , F xB, B xC , D 2 0.206 B D = 0.40 0.20 0.039 0.031 2 1 15 . 0.206 = 1166 . Therefore, of 14 equilibrium stages, (1.166/2.166)(14) = 7.5 stages are in the rectifying section. Assuming a total condenser, the feed stage is stage 7 or 8 from the top. Exercise 9.20 Subject: Comparison of the Kirkbride, Fenske, and McCabe-Thiele methods for determining feed-stage location. Given: A feed mixture of 25 mol% acetone (A) and 75 mol% water (W). Distillate to contain 95 mol% A, and bottoms to contain 2 mol% A. Infinite dilution activity coefficients from Exercise 9.6. Find: Ratio of rectifying to stripping stages by: (a) Fenske equation. (b) Kirkbride equation. (c) McCabe-Thiele method. Analysis: (a) Fenske equation: From Exercise 9.6: zF xD xB i,W, D i,W, B Component Acetone 0.25 0.95 0.02 1.48 27.7 Water 0.75 0.05 0.98 1.00 1.0 The relative volatility at feed conditions is also needed. Assume a bubble-point liquid feed. Also from vapor-liquid, y-x table of Exercise 9.6, ,W = [(0.789/0.250)]/[(0.211/0.750)] = 11.2 From Eq. (9-35), 1/ 2 N R log ( 0.95 / 0.25 )( 0.75 / 0.05 ) log[(27.7)(11.2)] 1.756(1.246) = = = 2. 96 1/ 2 1.213(0.6097) N S log ( 0.25 / 0.02 )( 0.98 / 0.75 ) log[(1.48)(11.2)] (b) Kirkbride equation: By material balances on A and W, B/D = 0.753/0.247 = 3.05 NR = NS zW, F zA, F xA, B xW, D 2 0.206 B D = 0.75 0.25 0.02 0.05 2 0.206 3.05 = 1.08 (c) McCabe-Thiele method: From the McCabe-Thiele graph of Exercise 9.6, shown on the next page, where the total number of minimum stages = 5. N R 4.2 = = 5.25 N S 0.8 The McCabe-Thiele method is the most accurate. The Fenske equation is better than the Kirkbride equation because the Kirkbride equation does not take into account the drastic change in relative volatility from the distillate to the bottoms. Exercise 9.20 (continued) Analysis: McCabe-Thiele method (continued) Exercise 9.21 Subject: Derivation of Eq. (9-37) for the effective stripping factor Given: Equations (5-46) and (5-47) Analysis: By analogy to Eqs. (5-46) and (5-47), the fraction of species in the entering feed liquid that is not stripped is given for just 2 stages by: S = 1 1 = 2 S N S1 + S N + 1 Se + Se + 1 Therefore, Se2 + Se + 1 = S N S1 + S N + 1 Solving for Se , Se2 + Se - S N S1 + 1 = 0 Se = -1 + 1 + 4 S N S1 + 1 2 = 0.25 + S N S1 + 1 - 0.5 : Which is Eq. (9-37). Exercise 9.22 Subject: Effect of number of stages and pressure on multicomponent absorption Given: Feed gas of 2,000 lbmol/h of light paraffin hydrocarbons at 60oF, and 500 lbmol/h of nC10 absorbent at 90oF. K-values in Fig. 2.8. Find: Separation at following conditions: (a) 6 stages and 75 psia. (b) 3 stages and 150 psia. (c) 6 stages and 150 psia Analysis: Apply Kremser's method, with absorption factors, A = L/KV, computed for L = entering absorbent rate = 500 lbmol/h, V = entering gas rate = 2,000 lbmol/h, and K-values from Fig. 2.8 at the average of the two entering temperatures, (90 + 60)/2 = 75oF and each of the two pressures. The feed gas composition and the resulting K-values and values of A are as follows: Component C1 C2 C3 nC4 nC5 Total: Feed, lbmol/h 1660 168 96 52 24 2000 K at 75 psia 29 5.8 1.68 0.48 0.145 A at 75 psia 0.00862 0.0431 0.149 0.521 1.724 K at 150 psia 15 3.1 0.93 0.27 0.08 A at 150 psia 0.0167 0.0806 0.269 0.926 3.13 The fraction of a gas component not absorbed, A, is given by the Kremser Eq. (5-48) and by material balance the amount absorbed follows: A -1 A = N +1 and l N = 1 - A f A -1 The results for all three cases are as follows: Component C1 C2 C3 nC4 nC5 Total: Case (a): Case (b): A lN, lbmol/h A lN, lbmol/h 0.991 14.9 0.983 28.2 0.957 7.2 0.919 13.6 0.851 14.3 0.731 25.8 0.484 26.8 0.280 37.4 0.0164 23.6 0.0224 23.5 86.8 128.5 Case (c): A lN, lbmol/h 0.983 28.2 0.919 13.6 0.731 25.8 0.178 42.7 0.00072 23.98 134.28 The results show that doubling the pressure is much more effective than doubling the number of equilibrium stages. Even Case (c) results in little additional absorption over Case (b) when the stages are doubled for the same pressure. Exercise 9.23 Subject: Effect of the number of stages on the absorption of a light normal paraffin hydrocarbon gas by the Kremser method. Given: 1000 kmol/h of gas at 70oF of composition below, and 500 kmol/h of nC10 absorbent at 90oF. Absorber operating at 4 atm. K-values in Fig.2.8. Find: Percent absorption for each component for (a) 4 stages, (b) 10 stages, and (c) 30 stages. Analysis: : Apply Kremser's method, with absorption factors, A = L/KV, computed for L = entering absorbent rate = 500 kmol/h, V = entering gas rate = 1,000 kmol/h, and K-values from Fig. 2.8 at the average of the two entering temperatures, (90 + 70)/2 = 80oF and a pressure of 4 atm = 59 psia. The feed gas composition and the resulting K-values and values of A are as follows: Component C1 C2 C3 nC4 nC5 Total: f, kmol/h 250 150 250 200 150 1000 K-value 35 7.7 2.2 0.65 0.20 A = L/KV 0.0142 0.065 0.227 0.769 2.50 The fraction of a gas component not absorbed, A, is given by the Kremser Eq. (5-48), and by material balance the amount absorbed follows: A -1 A = N +1 and Percent absorbed = 100 1- A A -1 The results for all three cases are as follows: Component C1 C2 C3 nC4 nC5 N = 4 stages: A 0.986 0.935 0.773 0.316 0.0155 % absorbed 1.4 6.5 22.7 68.4 98.4 N = 10 stages: N = 30 stages: % % A A absorbed absorbed 0.986 1.4 0.986 1.4 0.935 6.5 0.935 6.5 0.773 22.7 0.773 22.7 0.245 75.5 0.231 76.9 0.0001 99.99 0.0000 100.0 The above results show that as N increases, the percent absorption of the heavier components increases, but the absorption of the lighter components does not increase. However, beyond 10 stages, little change occurs. Exercise 9.24 Subject: Flow rate of stripping steam by the Kremser method Given: Feed of composition below enters a flash drum operating at 150oF and 2 atm. Equilibrium liquid is sent to a 5-equilibrium-stage stripper operating at 2 atm to give 0.5 kmol/h of nC5 in stripper bottoms. Find: Flow of steam needed in the stripper. Analysis: Use Chemcad with the SRK equation of state to flash the feed to determine the liquid feed to the stripper. The result is: Component C1 C2 C3 nC4 nC5 nC12 Total: Feed, kmol/h Distillate, kmol/h 13.7 13.6 101.3 98.2 146.9 134.9 23.9 19.1 5.7 3.4 196.7 0.4 488.2 269.6 Bottoms, kmol/h 0.1 3.1 12.0 4.8 2.3 196.3 218.6 The stripping steam enters the stripper at 2 atm and 300oF. For the stripper, the average temperature of the two feeds = (150 + 300) = 225oF. From Fig. 2.8, the K-value of nC5 at 2 atm and 225oF = 3.0. The average stripping factor for n-pentane can be taken as: S = KV/L = 3V/218.6 = 0.0137 V The fraction of nC5 not stripped = S = 0.5/2.3 = 0.217 Use Kremser Eq. (5-50) to compute the value of S needed for N = 5 theoretical stages. By trial and error, the result from: S = is S = 0.89. Therefore, S -1 = 0.217 S N +1 - 1 Flow rate of stripping steam = V = S/0.0137 = 0.89/0.0137 = 65 kmol/h Exercise 9.25 Subject: Stripping of a hydrocarbon liquid with superheated steam. Given: 1000 kmol/h of feed liquid at 250oF with composition below. 100 kmol/h of superheated steam at 300oF and 50 psia. Stripper operates at 50 psia and has 3 equilibrium stages. Assumptions: Negligible stripping of nC10 and negligible condensation of steam Find: Flow rates and compositions of the exiting streams by the group (Kremser) method. Analysis: Apply Kremser's method, with stripping factors, S = KV/L , computed for L = entering feed = 1000 kmol/h, V = entering steam rate = 100 kmol/h, and K-values from Fig. 2.8 at the average of the two entering temperatures, (250 + 300)/2 = 275oF and a pressure of 50 psia. The feed gas composition and the resulting K-values and values of S are as follows: Component C1 C2 C3 nC4 nC5 nC10 Total: f, kmol/h 0.3 2.2 18.2 44.7 85.9 848.7 1000.0 K-value 67 26 12.2 6.0 2.85 S = KV/L 6.7 2.6 1.22 0.60 0.285 The fraction of a gas component not stripped, S, is given by the Kremser Eq. (5-50), and by material balance the amounts stripped and not stripped follow: S -1 S = N +1 and l1 = S f and v3 = f - l1 S -1 The results are as follows: Component C1 C2 C3 nC4 nC5 Total: S 0.0028 0.036 0.181 0.46 0.72 l1, kmol/h 0.0 0.1 3.3 20.6 61.8 85.8 v3, kmol/h 0.3 2.1 14.9 24.1 24.1 65.5 The exiting gas is 65.5 + 100 = 165.5 kmol/h. The exiting liquid is 85.8 + 848.7 = 934.5 kmol/h. Exercise 9.26 Subject: Liquid-liquid extraction of a hydrocarbon mixture by diethylene glycol (DEG) by the group (Kremser) method. Given: 100 kmol/h of an equimolar mixture of benzene (B), toluene (T), n-hexane (C6), and n-heptane (C7). Extraction at 150oC with 300 kmol/h of DEG. Extractor has 5 equilibrium stages. Distribution coefficients given below. Find: Flow rates and compositions of extract and raffinate. Analysis: Treat the problem like a stripper. Apply Kremser's method, with extraction factors, Eq. (9-39), E = KDV/L , computed for L = entering feed = 100 kmol/h, V = entering DEG solvent rate = 300 kmol/h, and given values of KD = mole fraction in extract/mole fraction in raffinate. The extract is like the vapor in stripping and the raffinate is like the liquid in stripping. The feed composition and the K-values and values of S are as follows: Component Benzene Toluene n-Hexane n-Heptane Total: f, kmol/h 25 25 25 25 100 KD 0.33 0.29 0.050 0.043 E = KDV/L 0.99 0.87 0.15 0.129 The fraction of a feed component not extracted, E , is given by the Kremser Eq. (9-44), and by material balance the amounts stripped and not stripped follow: E -1 E = N +1 and l1 = E f and v3 = f - l1 E -1 The results are as follows: Component Benzene Toluene n-Hexane n-Heptane Total: E 0.177 0.230 0.850 0.871 l1, kmol/h 4.3 5.8 21.2 21.8 53.1 v3, kmol/h 20.7 19.2 3.8 3.2 46.9 Now compute the transfer of DEG to the raffinate using Eqs. (9-40) and (9-43), using KD = 30 U = L/KDV = 100/[30(300)] = 0.0111. U = 0.9889. Therefore 0.9889(300) = 296.7 kmol/h of DEG leaves in the extract and 300 - 296.7 = 3.3 kmol/h in the raffinate. The exiting extract flow rate = 46.9 + 296.7 = 343.6 kmol/h The exiting raffinate flow rate = 53.1 + 3.3 = 56.4 kmol/h. Exercise 9.27 Subject: Reboiled stripping of a normal paraffin hydrocarbon feed liquid by the group method. Given: Feed liquid at 39oF and 300 psia and of composition given below is adiabatically flashed to 150 psia before entering a stripper containing 7 equilibrium stages and a partial reboiler. The bottoms flow rate from the stripper = B = 99.3 lbmol/h. Find: Compositions of the vapor and liquid products from stripper. Analysis: First, adiabatically flash the feed given below, using Chemcad with the SRK equation of state. The results of the flash are as follows, where the calculated flash temperature = 26.3oF: Lbmol/h: Component Feed Vapor C1 59.5 37.67 C2 73.6 14.83 C3 153.2 8.90 nC4 173.5 2.58 nC5 58.2 0.23 nC6 33.6 0.04 Total: 551.6 64.25 Liquid 21.83 58.77 144.30 170.92 57.97 33.56 487.35 The liquid from the adiabatic flash is sent to the reboiled stripper where it is separated into an overhead vapor and a liquid bottoms. The overhead vapor is mixed with the vapor from the adiabatic flash to give the final vapor product. To solve the reboiled stripper by the group method, apply Edmister's method, rather than Kremser's method, because we do not know the composition and flow rate of the vapor leaving the reboiler and entering the column. For a reboiled stripper, Eq. (5-64) applies: lF S B AX + 1 KV = (1) where S B = B B (2) b SX B For AX and , use average absorption and stripping factors computed for L = entering flash liquid = 487.35 kmol/h, V = vapor leaving the reboiler, which can be taken as L - B = 487.35 99.30 = 388.05 kmol/h, and K-values from Fig. 2.8 at an average temperature of say 200oF and a pressure of 150 psia. The feed gas composition and the resulting K-values, values of S = KV/L and A = 1/S, and corresponding values of AX and EX , computed from Eqs. (5-48) and (5-50), AX = are as follows: A -1 S -1 and SX = N +1 N +1 A -1 S -1 Exercise 9.27 (continued) Analysis: (continued) Component C1 C2 C3 nC4 nC5 nC5 Total: lF , lbmol/h 21.83 58.77 144.30 170.92 57.97 33.56 487.35 K-value 22 7.1 2.9 1.23 0.52 0.23 S = KV/L 17.5 5.65 2.31 0.979 0.414 0.183 A = 1/S 0.0571 0.177 0.433 1.021 2.415 5.46 SX 0.000 0.000 0.004 0.152 0.587 0.817 AX 0.943 0.823 0.569 0.134 0.003 0.000 To compute values of SB from Eq. (2), the assumed value of VB is critical. We seek the value of VB that will give values of SB , which when substituted into Eq. (1) will give values of bi that will sum to the specified value of B = 99.3 lbmol/h. Assume the temperature at the bottom is 250oF and obtain from Fig. 2.8 values of K at 250oF and 150 psia. To begin, assume a value of VB = 500 lbmol/h, calculate B = bi , and then adjust the values of bi so they sum to the specified B. The results are as follows: Component C1 C2 C3 nC4 nC5 nC5 Total: KB 23 8.8 4.0 1.85 0.88 0.45 SB 116 44.3 20.1 9.31 4.43 2.27 lF/b ---3100 14.8 1.73 1.22 lbmol/h: b 0.00 0.00 0.05 11.55 33.51 27.44 72.55 Adj. b 0.00 0.00 0.07 25.92 45.87 27.44 99.30 The final product distribution is as follows, where it is compared to the result obtained by using Chemcad with the Tower model to also calculate the reboiled stripper: Edmister method: Rigorous Chemcad: Component Feed, Vapor Stripped Vapor Stripped lbmol/h product, liquid, product, liquid, lbmol/h lbmol/h lbmol/h lbmol/h C1 59.5 59.5 0.0 59.5 0.0 C2 73.6 73.6 0.0 73.6 0.0 C3 153.2 153.1 0.1 153.1 0.1 nC4 173.5 147.6 25.9 150.8 22.7 nC5 58.2 12.3 45.9 12.7 45.5 nC5 33.6 6.2 27.4 2.6 31.0 Total: 551.6 452.3 99.3 452.3 99.3 The Edmister method is uncertain because average temperature and vapor rate are difficult to estimate. Exercise 9.28 Subject: Comparison of FUG method with Edmister group method for distillation. Given: Bubble-point liquid feed in kmol/h of 100 ethylbenzene, 100 paraxylene, 200 metaxylene, and 100 orthoxylene. Column pressures of 25 psia at the top and 35 psia at the bottom. Distillate to contain 1 kmol/h of orthoxylene and bottoms to contain 2 kmol/h metaxylene. Column to be equipped with a total condenser and partial reboiler. Assumption: Applicability of Raoult's law for estimating K-values. Find: (a) Number of stages and reflux ratio by the FUG method for R/Rmin = 1.1 Feed-stage location by the Kirkbride equation. (b) Distribution of components between distillate and bottoms by Edmister method. Analysis: Need K-values at distillate, feed, and bottoms conditions. Because metaxylene and paraxylene have very close boiling points, assume the d/b ratio of paraxylene is the same as that for metaxylene. Therefore, have 1 kmol/h of paraxylene in the bottoms. Use Chemcad to obtain K-values by running bubble points on the feed, and estimated distillate and bottoms compositions. The results are: K-values: b, 25 psia, 30 psia, 35 psia, Component d, kmol/h kmol/h 159.3oC 168.4oC 180.5oC Ethylbenzene, EB 100 0 1.051 1.080 1.206 Paraxylene, PX 99 1 0.986 1.012 1.131 Metaxylene, MX 198 2 0.982 1.012 1.136 Orthoxylene, OX 1 99 0.855 0.884 0.996 Total: 398 101 distillat feed bottoms e (a) Using the Fenske equation (9-12), with MX as the LK and OX as the HK, with a geometric mean relative volatility = [(0.982/0.855)(1.136/0.996)]1/2 = 1.144 log N min = d MX bOX d OX bMX log = log ( MX,OX )avg (198 ) ( 99 ) (1) ( 2 ) log1.144 = 68. 3 To compute the distribution of nonkey components at total reflux, use for the lighter than light key components, LLK, Eq. (9-15), with OX as the reference component, r, and the above value of Nmin. Exercise 9.28 (continued) Analysis: (a) (continued) Thus, for the LLK, i, bi = fi d N min 1 + OX i,OX bOX = fi 1 68.3 1+ i,OX 99 (1) Using geometric mean values of i, OX, with Eq. (1) and the material balance, fi = di + bi , the following results are obtained: kmol/h: Component Feed Distillate Bottoms EB 100 99.99 0.01 PX 100 99.00 1.00 MX 200 198.00 2.00 OX 100 1.00 99.00 Total: 500 397.99 102.01 Because it is highly likely that PX, and possibly EB, will distribute, use the Class 2 Underwood equation, Eq. (9-28),which for a bubble-point feed, using relative volatilities at the feed conditions is: i ,OX zi , F 12217(0.20) 11448(0.20) 11448(0.40) 100(0.20) . . . . (2) = 1- q = 0 = + + + 12217 - 11448 - 11448 - 100 - . . . . i i ,OX - There are two roots to Eq. (2), which from a spreadsheet calculation are: = 1.0280 and 1.2027 Eq. (9-29) is applied in the following form for each of the two values of , in terms of the two unknowns, dEB and Lmin. The form used is obtained by multiplying Eq. (9-29) by the total distillate rate, D, which converts distillate mole fractions to component distillate rates, and the reflux ratio, R to L. The distillate rate of PX is not an unknown because it has the same relative volatility as MX. Thus, it must have the same d/b ratio as MX, making its distillate rate as 99 kmol/h. D + Lmin = 12217 (d EB ) . 1.1448 (297) 1.00 (1) + + 12217 - 10280 11448 - 1.0280 1.00 - 10280 . . . . = 6.307(d EB ) + 2875.3 D + Lmin = 12217(d EB ) . 1.1448(297) 1.00 (1) + + 12217 - 12027 11448 - 1.2027 1.00 - 1.2027 . . . = 64.30(d EB ) - 5877.2 Exercise 9.28 (continued) Analysis: (a) (continued) Because the total distillate rate, D, is also an unknown, the following equation is also needed: D = dEB + 99 + 198 + 1 = dEB + 298 Solving these 3 linear equations, the following results are obtained: D =448.9 kmol/h, Lmin = 3378 kmol/h, and dEB = 150.9 kmol/h. Thus, the assumption that EB distributes is incorrect because only 100 kmol/h of EB is in the feed. Therefore discard the root of = 1.2027. Also, D now equals 100 + 99 + 198 + 1 = 398 kmol/h. From above for the root of 1.0280, D + Lmin = 6.307 dEB + 2875.3. Therefore, Lmin = 6.307(100) + 2875.3 - 398 = 3108 kmol/h or Rmin = Lmin/D = 3108/398 = 7.81 The operating reflux ratio = 1.1(7.81) = 8.59. In the Gilliland equation, (9-34), X = (R - Rmin)/(R + 1) = (8.59 - 7.81)/(8.59 + 1) = 0.0813. Using Eq. (9.34), Y = 0.573 = (N - Nmin)/(N + 1). Solving, for Nmin = 68.3, N = 161.3 equilibrium stages. The Kirkbride Eq. (9-36) is used to determine the feed-stage location. NR = NS zOX , F zMX , F x MX , B xOX , D 2 0.206 B D = 0.20 0.40 0.0196 0.00251 2 102.01 397.99 0.206 = 153 . Therefore, the feed stage is located at (1.53/2.53) 161.3 = stage 97.5 from the top. (b) The Edmister group method is applied by using Eq. (5-66) to estimate the b/d ratio for each component: b = AF d AC SE + 1 AE S B AX + 1 SX (3) For a total condenser, AC = the reflux ratio = 8.59 for all components. To compute values of for each component, assume the following values of flow rates and stages for the enriching (rectifying) section and exhausting (stripping) section, based on constant molar overflow and noting that the feed and reboiler stages are not included in these two sections. Variable Number of stages Vapor rate, kmol/h Liquid rate, kmol/h Enriching Section 97 3817 3419 Exhaustin g Section 63 3817 3919 Exercise 9.28 (continued) Analysis: (b) (continued) Because in the above table the K-values do not vary much, use arithmetic mean values for each section. These values and those for the feed stage and partial reboiler are as follows along with the corresponding values of the absorption and stripping factors, as defined by equations (5-38), (5-51), below (5-64), and (5-65), where VB/B = 3817/102 = 37.4 Variable KE AE SE = 1/AE KF AF KX AX SX = 1/AX KB SB EB 1.066 0.840 1.190 1.080 0.951 1.143 0.898 1.114 1.206 45.1 PX 0.999 0.897 1.115 1.012 1.015 1.072 0.958 1.044 1.131 42.3 MX 0.997 0.898 1.114 1.012 1.015 1.074 0.956 1.046 1.136 42.5 OX 0.870 1.030 0.971 0.884 1.161 0.940 1.092 0.918 0.996 37.3 Values of A and S are obtained from Eqs. (5-48) and (5-50), respectively: Variable AE SE X SX EB PX MX 0.160 0.103 0.102 0.0000 0.0000 0.0000 0.102 0.0449 0.0466 0.0001 0.0030 0.00274 OX 0.00175 0.0307 0.0003 0.0823 Now apply Eq. (3) above to obtain the component distribution: Variable f, kmol/h b/d d, kmol/h b, kmol/h EB 100 0.0001 99.99 0.01 PX 100 0.0102 98.99 1.01 MX 200 0.0091 198.20 1.80 OX 100 68.2 0.68 99.32 Comparing these Edmister method results with the FUG method, we see that the Edmister method predicts a somewhat better split of the two key components than does the FUG method for this problem. Exercise 10.1 Subject: Independency of MESH equations Given: Equations (10-1), (10-3), (10-4), and (10-6). Prove: Equation (10-6) is not independent of the other 3 equations Analysis: Eq. (10-6) can be derived from the other 3 equations as follows, as outlined in the text between Eqs. (10-5) and (10-6).. Summing Eq. (10-1) over all C components: L j -1 C i =1 xi , j -1 + V j +1 C i =1 yi , j +1 + Fj C i =1 zi , Fj - L j + U j C i =1 xi , j - V j + Wj C i =1 yi , j = 0 (1) From Eqs. (10-3) and (10-4), all 5 sums in Eq. (1) are equal to 1. Therefore, Eq. (1) becomes: L j -1 + V j +1 + F j - L j + U j - V j + W j = 0 Writing Eq. (2) for each stage from Stage 1 to Stage j: (2) L0 + V2 + F1 - L1 - U 1 - V1 - W1 = 0 L1 + V3 + F2 - L2 - U 2 - V2 - W2 = 0 L2 + V4 + F3 - L3 - U 3 - V3 - W3 = 0 ....... L j - 2 + V j + Fj -1 - L j -1 - U j -1 - V j -1 - Wj -1 = 0 L j -1 + V j +1 + F j - L j - U j - V j - W j = 0 (3) Summing Eqs. (2), noting that L0 = 0 and that many variables cancel, we obtain: V j +1 + j m =1 Fm - U m - Wm - L j - V1 = 0 j m =1 or L j = V j +1 + Fm - U m - Wm - V1 But this is Eq. (10-6). Therefore, it is not independent of Eqs. (10-1), (10-3), and (10-4). Exercise 10.2 Subject: Revision of MESH equations to account for entrainment, occlusion, and chemical reaction (in the liquid phase). Given: MESH Eqs. (10-1) to (10-5). Find: Revised set of MESH equations. Analysis: Entrainment: Let j = ratio of entrained liquid (in the exiting vapor) that leaves Stage j to the liquid (Lj + Uj) leaving Stage j. Then, the entrained component liquid flow rate leaving Stage j = jxi,j (Lj + Uj). Correspondingly, the entrained component liquid flow rate entering Stage j = j+1xi,j+1 (Lj+1+ Uj+1). Occlusion: Let j = ratio of occluded vapor (in the exiting liquid) that leaves Stage j to the vapor (Vj + Wj) leaving Stage j. Then the occluded component liquid flow rate leaving Stage j = j yi,j (Vj + Wj). Correspondingly, the occluded component liquid flow rate entering Stage j = j-1 yi,j-1(Vj-1 + Wj-1 ). Chemical Reaction: Let: Mj = molar liquid volume holdup on Stage j M = number of independent chemical reactions i,m = stoichiometric coefficient of component i in chemical reaction m rk,m,j = chemical reaction rate, dck,m,j/dt, of the mth chemical reaction for the reference reactant component, k, on Stage j Then, the formation or disappearance of component i by chemical reaction on Stage j is: Mj M m =1 i ,mrk ,m, j The component material balance equations, (10-1), now become: L j -1 C i =1 xi , j -1 + V j +1 C i =1 yi , j +1 + Fj C i =1 zi , F j - L j + U j C i =1 xi , j - V j + Wj C i =1 yi , j + M m =1 j -1 V j -1 + Wj -1 yi , j -1 - j V j + Wj yi , j + j +1 L j +1 + U j +1 xi , j +1 - j L j + U j xi , j - M j i ,mrk ,m, j = 0 The equilibrium equations, (1-2) and the summation equations (10-3) and (10-4) do not change. The energy balance is as follows, where the heat of reaction does not appear because it is assumed that enthalpies are referred to the elements: L j -1hL j -1 + V j +1hV j +1 + Fj hFj + j +1 V j -1 + Wj -1 hV j -1 + j +1 L j +1 + U j +1 hL j +1 - 1 + j V j + Wj hV j - 1 + j L j + U j hL j = 0 Exercise 10.3 Subject: Revised set of MESH equations. Given: MESH Eqs. (10-1) to (10-5). Find: Revised set of MESH equations to account for liquid pumparounds as shown in Fig. 10.2. Revised M equations like Eq. (10-7). Possible partitioning to tridiagonal sets. Analysis: As an example, consider the section of equilibrium stages below, which includes a bypass pumparound and a recycle pumparound. Let Pu,v designate a total molar liquid pumparound flow rate from Stage u to Stage v. Exercise 10.3 (continued) Analysis: (continued) The revised MESH equations for Stage j are: M i , j = L j -1 xi , j -1 + V j +1 yi , j +1 + Pj + 2, j xi , j + 2 + Pj - 2, j xi , j - 2 + F j zi , j - ( L j + U j ) xi , j - (V j + W j ) yi , j = 0 (1) (2) Ei , j = yi , j - K i , j xi , j = 0 H j = L j -1hL j-1 + V j +1hV j+1 + Fj hFj + Pj - 2, j hL j-2 + Pj + 2, j hL j+2 - (L j + U j ) hL j - (V j + W j ) hV j - Q j = 0 j m =1 (5) L j = V j +1 + ( Fm - U m - Wm ) - V1 + Pj + 2, j = 0 (6) Equations (1), (2), and (6) are combined to give the following modified Mi,j equations in terms of component mole fractions, assuming that all other variables are specified or guessed: M i , j = xi , j -1 V j + j -1 m =1 Fm - U m - Wm - V1 - Pj - 2 , j + xi , j +1 Ki , j +1V j +1 + Pj + 2 , j xi , j + 2 + Pj - 2 , j xi , j - 2 + j m =1 Fj zi , j - xi , j V j +1 + Fm - U m - Wm - V1 + Pj + 2 , j + U j + V j + Wj Ki , j = 0 Compared to Eqs. (10-7) to (10-11), we see that the modified Mi,j equations contain 5, rather than 3, unknown liquid-phase mole fractions. The additional mole fractions are xi,j+2 and xi,j-2. Therefore, the set is no longer tridiagonal. We have a banded matrix with 5 bands, one additional band due to the recycle and one due to the bypass. Exercise 10.4 Subject: Application of tridiagonal matrix algorithm. Given: Thomas algorithm equations (10-13) to (10-18) and a 3x3 matrix equation below. Find: Solution to the matrix equation. Analysis: The matrix equation is: B1 A2 0 C1 B2 A3 0 B3 x1 x3 D1 D3 -160 0 200 0 x1 x3 0 0 C2 x2 = D2 = 50 - 350 180 x2 = -50 150 - 230 Applying the algorithm Eqs. (10-13) to (10-15): p1 = p2 = q2 = C1 200 = = -1.25 B1 -160 q1 = D1 0 = =0 B1 -160 C2 180 = = -0.626 . B2 - A2 p1 -350 - 50( -125) D2 - A2 q1 -50 - 50(0) . = = 01739 B2 - A2 p1 -350 - 50( -1.25) . D3 - A3q2 0 - 150(01739) = = 01917 . B3 - A3 p2 -230 - 150( -0.626) p3 = 0 q3 = Using Eqs. (10-17) and (10-18): x3 = q3 = 0.1917 x2 = q2 - p2x3 = 0.1739 - (-0.626)(0.1917) = 0.2939 x1 = q1 - p1x2 = 0 - (-1.25)(0.2939) = 0.3674 Exercise 10.5 Subject: Application of tridiagonal matrix algorithm. Given: Thomas algorithm equations (10-13) to (10-18) and a 3x3 matrix equation below. Find: Solution to the matrix equation. Analysis: The matrix equation is: B1 A2 0 0 0 C1 B2 A3 0 0 0 C2 B3 A4 0 0 0 C3 B4 A5 0 0 C4 B5 x1 x2 x4 x5 D1 D2 D4 D5 -6 0 0 0 3 0 3 0 0 3 4.5 0 0 0 3 - 4.5 x1 x2 x4 x5 0 0 0 0 3 - 4.5 0 0 0 x3 = D3 = 1.5 - 7.5 0 x3 = 100 4.5 - 7.5 Applying the algorithm Eqs. (10-13) to (10-15): p1 = p2 = p3 = q3 = p4 = q4 = q5 = C1 3 = = -0.5 B1 -6 q1 = D1 0 = =0 B1 -6 q2 = 0 C2 3 = = -1.0 B2 - A2 p1 -4.5 - 3( -0.5) C3 3 = = -0.5 B3 - A3 p2 -7.5 - 15( -1.0) . 100 - 15(0) D3 - A3q2 . = = -16.67 B3 - A3 p2 -7.5 - 15( -10) . 3 C4 = = -0.57143 B4 - A4 p3 -7.5 - 4.5( -0.5) D4 - A4 q3 0 - 4.5( -16.67) = = -14.28571 B4 - A4 p3 -7.5 - 4.5( -0.5) D5 - A5q4 0 - 4.5( -14.28571) = = -33.333 B5 - A5 p4 -4.5 - 4.5( -0.57143) Using Eqs. (10-17) and (10-18): x5 = q5 = -33.333 x4 = q4 - p4x5 = -14.28571 - (-0.57143)(-33.333) = -33.333 x3 = q3 - p3x4 = -16.6667 - (-0.5)(-33.333) = -33.333 x2 = q2 - p2x3 = 0.0 - (-1.0)(-33.333) = -33.333 x1 = q1 p1x2 = 0.0 - (-0.5)(-33.333) = -16.667 Exercise 10.6 Subject: Solution of tridiagonal matrix equation for liquid-phase mole fractions to avoid subtraction of nearly equal quantities as claimed by Wang and Henke. Given: Equations (10-1) to (10-18) Prove: Claim of Wang and Henke. Analysis: First investigate the values of pj. From Eq. (10-14), a subtraction occurs in the denominator: Cj pj = B j - A j p j -1 Starting at the top of the column, p1 = C1/B1, therefore, the denominator of Eq. (10-14) becomes: C B2 - A2 p1 = B2 - A2 1 Clearing the fraction, we obtain B1 B2 - A2 C1 B1 Examine this difference by substituting for each of the four quantities. From Eq. (10-8), A2 = V2 + F1 - W1 - U1 - V1 However, by material balance around Stage 1 in Fig. 10.3, V2 + F1 - W1 - U 1 - V1 = L1 Therefore, A2 = L1 From Eq. (10-9), B1 = -[V2 + F1 - W1 - U 1 - V1 + U 1 + V1 + W1 Ki ,1 ] However, by material balance around Stage 1 in Fig. 10.3, V2 + F1 - W1 - U 1 - V1 = L1 Therefore, B1 = - V1 + W1 Ki ,1 + L1 + U 1 Similarly, B2 = - V2 + W2 Ki ,2 + L2 + U 2 Thus, it is clear that B1 B2 = a summation of all positive quantities, that is no subtractions. From Eq. (10-10), C1 = V2 Ki ,2 and A2 C1 = L1V2 Ki ,2 When B1B2 is expanded, it contains a term -L1V2Ki,2 . Thus, this potential subtraction in B1 B2 - A2 C1 is cancelled. There are no subtractions in p2 . The same is found for all pj . Now investigate values of qj. From Eq. (10-15), subtractions occur in the numerator and denominator: Exercise 10.6 (continued) Analysis: (continued) qj = D j - A j q j -1 B j - A j p j -1 The denominator is identical to that already examined in pj. So we only have to investigate the numerator. Using Eqs. (10-9), as above, and (10-11), q1 = - F1zi ,1 D1 = B1 - V1 + W1 Ki ,1 + L1 + U 1 which contains no subtraction problems Therefore, the numerator in Eq. (10-15) becomes: D2 - A2 q1 = - F2 zi ,2 - L1 - F1zi ,2 - V1 + W1 Ki ,1 + L1 + U 1 = F2 zi ,2 V1 + W1 Ki ,1 + L1 + U 1 + L1 F1zi ,1 - V1 + W1 Ki ,1 + L1 + U 1 Thus, no subtractions appear in this numerator. The result is obtained for all qj. Therefore, the claim of Wang and Henke is verified. Exercise 10.7 Subject: Derivation of stagewise component material balance equations in terms of component vapor flow rates rather than liquid-phase mole fractions. Given: Eq. (10-7). Find: Component flow rates form of Eq. (10-7). Whether equations can still be partitioned into C tridiagonal matrix equations. Analysis: Start with Eqs. (10-1), (10-2), and (10-6) for Lj and Lj-1: M i , j = L j -1 xi , j -1 + V j +1 yi , j +1 + Fj zi , j - L j + U j xi , j - V j + Wj yi , j = 0 Ei , j = yi , j - Ki , j xi , j = 0 L j = V j +1 + L j -1 = V j + j m=1 j -1 m=1 (1) (2) (3) (4) Fm - U m - Wm - V1 Fm - U m - Wm - V1 Substituting Eqs. (2) to (4) into (1), to eliminate variables in L and x, M = V + i, j j - V j -1 m + j i, j - 1 F - U - W -V + +F z m m m i, j + 1 j i, j 1 V K i, j - 1 i, j - 1 i, j i, j F - U - W -V + U - -W =0 j +1 m m m 1 j V K i, j j V m j i, j j Collecting terms in the component vapor flow rates, M = V + i, j j - V j -1 ( Fm - U m - Wm ) - V1 V m j i, j - 1 + K i, j - 1 i, j - 1 W 1 j -1 - + +F z =0 i, j i, j + 1 j i, j V K V j i, j j j +1 + ( Fm - U m - Wm ) - V1 + U j m Thus, we will have a set of equations that can be partitioned into C tridiagonal sets. Exercise 10.8 Subject: Computer storage locations for tridiagonal matrix equations Given: Eqs. (10-7) to (10-11) Find: Minimum storage locations needed. Analysis: For a 100-stage column, only need: 99 storage locations for pj 100 storage locations for qj 4 storage locations for current values of Aj, Bj, Cj, and Dj This is total of 203 storage locations. If Aj, Bj, Cj, and Dj expressions are substituted into the pj and qj equations, then only need 199 storage locations. Exercise 10.9 Subject: Solving a set of two nonlinear equations by the Newton-Raphson method. Given: Two nonlinear equations in x1 and x2: 2 x12 + x2 = 17 8 x1 1/ 3 1 + x2 / 2 = 4 Find: Solutions for four different starting guesses Analysis: (a) Starting guess is x1 = 2, x2 = 5 The Newton-Raphson method converges in 5 iterations: Iteration Start 1 2 3 4 5 x1 2.000 0.5237 0.8866 0.9943 0.9999873 1.0000 x2 5.000 4.3911 4.0569 4.0032 4.0000 4.0000 (b) Starting guess is x1 = 4, x2 = 5 The Newton-Raphson method fails. After only one iteration, x1 = -6.2135, x2 = 10.7719 (c) Starting guess is x1 = 1, x2 = 1 The Newton-Raphson method fails. After only one iteration, x1 = -15.4910, x2 = 24.9880 (d) Starting guess is x1 = 8, x2 = 1 The Newton-Raphson method fails. After only one iteration, x1 = 5.1320, x2 = -0.4449 Note that by a homotopy continuation method, two roots in the positive domain are obtained: x1 = 1, x2 = 4 x1 = 4.0715, x2 = 0.65027 If the Newton-Raphson method is damped by using a line search, the method fails less often. Exercise 10.10 Subject: Solving a set of two nonlinear equations by the Newton-Raphson method. Given: Two nonlinear equations in x1 and x2: sin x1 x2 - exp 2 x1 x2 - x1 = 0 2 1 1 1- + exp 1 - 1 - 2 x1 + x2 = 0 4 4 Find: Solutions for three different starting guesses Analysis: (a) Starting guess is x1 = 0.4, x2 = 0.9 The Newton-Raphson method converges in 4 iterations: Iteration x1 x2 Start 0.4000 0.9000 1 0.1415 0.8394 2 0.2806 0.9073 3 0.2984 0.9026 4 0.2994 0.9030 (b) Starting guess is x1 = 0.6, x2 = 0.9 The Newton-Raphson method converges to a different solution in just 3 iterations. Iteration x1 x2 Start 0.6000 0.9000 1 0.4901 1.0235 2 0.4999 1.0002 3 0.5000 1.0000 (c) Starting guess is x1 = 1, x2 = 1 The Newton-Raphson method converges with some oscillations in 14 iterations. Iteration x1 x2 Iteration x1 x2 Start 1.0000 1.0000 8 1.0042 0.4451 1 0.9091 0.6914 9 0.7160 1.2837 2 1.0720 0.2987 10 0.6514 0.9886 3 0.9727 0.5500 11 0.5135 1.0452 4 1.1633 -0.0252 12 0.5045 1.0008 5 1.0511 0.3321 13 0.5001 1.0001 6 0.9512 0.6053 14 0.5000 1.0000 7 1.0876 0.2072 Exercise 10.11 Subject: Calculations for the first iteration of the Bubble-Point method Given: Feed of 1000 kmol/h of a bubble-point mixture of 60 mol% methanol (M), 20 mol% ethanol (E), and 20 mol% n-propanol (P) is sent to the middle stage of a column with a total condenser, 3 equilibrium stages in the column, and a partial reboiler, and operating at 1 atm. The distillate rate = 600 kmol/h and the reflux rate = 2,000 kmol/h. Assumptions: Ideal solutions with Raoult's law K-values. Constant molar overflow. Find: Liquid-phase component mole fractions and stage temperatures from the first iteration of the bubble-point method. Analysis: By material balance, the bottoms flow rate = 1000 - 600 = 400 kmol/h. The liquid rate above the feed = 2000 kmol/h. The liquid rate below the feed = 3000 kmol/h. The boilup rate over the entire column = 3000 - 400 = 2600 kmol/h. Based on the distillate rate, methanol will be mainly in the distillate, with ethanol and n-propanol mainly in the bottoms. The boiling points of these three components are, respectively, 64.7oC, 78.4oC, and 97.8oC. As a first approximation, assume a distillate temperature of 64.7oC and a bottoms temperature of (78.4 + 97.8)/2 = 88.1oC. Therefore, the initial guesses, assuming a linear temperature profile, for the tear variables and the total liquid flow rates are: Stage j 1 (Total condenser) 2 3 4 5 (Partial reboiler) Vj , kmol/h 0 2600 2600 2600 2600 Tj , oC 64.7 70.6 76.4 82.3 88.1 Lj , kmol/h 2000 2000 3000 3000 400 Compute K-values at 1 atm and the above temperatures for each component by Eq. (2-44), K = Pis/P. From the 7th edition of Perry's Handbook, on page 2-50 to 2-54, in Table 2-6, vapor pressure data are fitted to an extended Clausius-Clapeyron equation: P s = exp C1 + C2 + C 3 ln T + C 4T C 5 , P s is in Pa and T is in K , where the constants are: T C1 81.768 74.475 88.134 C2 -6876 -7164.3 -8498.6 Component M E P C3 -8.7078 -7.327 -9.0766 C4 7.1926E-6 3.1340E-6 8.3303E-18 C5 2 2 6 Exercise 10.11 (continued) Analysis: (continued) Computed K-values at the five assumed stage temperatures for P = 101300 Pa are: Stage Temperature, K K-values: M E P 1 337.9 1.0086 0.5705 0.2505 2 343.7 1.2639 0.7309 0.3294 3 349.6 1.5702 0.9272 0.4281 4 355.4 1.9350 1.1654 0.5504 5 361.3 2.3665 1.4521 0.7002 A plot of the K-values as a function of temperature is shown on the next page. The component material balance equations, combined with the equilibrium equations to eliminate vapor-phase mole fractions, for each of the 5 stages are as follows: Stage 1: Stage 2: - D + L1 xi ,1 + V2 Ki ,2 xi ,2 = 0 - 2600 xi ,1 + 2600 Ki ,2 xi ,2 = 0 L1 xi ,1 - V2 Ki ,2 + L2 xi ,2 + V3 Ki ,3 xi ,3 = 0 2000 xi ,1 - 2600 Ki ,2 + 2000 xi ,2 + 2600 Ki ,3 xi ,3 = 0 (2) L2 xi ,2 - V3 Ki ,3 + L3 xi ,3 + V4 Ki ,4 xi ,4 = - F3 zi ,3 2000 xi ,2 - 2600 Ki ,3 + 2000 xi ,3 + 2600 Ki ,4 xi ,4 = -1000zi ,3 (3) L3 xi ,3 - V4 Ki ,4 + L4 xi ,4 + V5 Ki ,5 xi ,5 = 0 3000 xi ,3 - 2600 Ki ,4 + 3000 xi ,4 + 2600 Ki ,5 xi ,5 = 0 (4) L4 xi ,4 - V5 Ki ,5 + B xi ,5 = 0 3000 xi ,4 - 2600 Ki ,5 + 400 xi ,5 = 0 (5) (1) Stage 3: Stage 4: Stage 5: For each component, the K-values for each stage are substituted from the table above. The resulting equations form 3 tridiagonal matrix equations, one each for each of the 3 components. These equations are as follows, as produced from a spreadsheet: Exercise 10.11 (continued) Analysis: (continued) Analysis: (continued) Exercise 10.11 (continued) ___________________________________________________________________________________________ Exercise 10.11 (continued) Analysis: (continued) A new set of stage temperatures is computed by running bubble-point temperature calculations to find a temperature that satisfies Eq. (10-21), using the normalized mole fractions with the vapor pressure equations and a pressure of 1 atm. The results are as follows compared to the initially assumed values: Stage j 1 (Total condenser) 2 3 4 5 (Partial reboiler) Tj , o C Tj , o C initially from 1st iteration assumed 64.7 67.2 70.6 69.6 76.4 73.1 82.3 77.8 88.1 84.1 Exercise 10.12 Subject: Solving a block tridiagonal matrix equation. Given: 9 linear equations in 9 unknowns. Find: Values of the 9 unknowns using the block form of the Thomas algorithm. Analysis: To form a block tridiagonal matrix structure, interchange the last 3 equations with the middle group of 3 equations. The result is: 0 1 2 1 0 1 1 1 1 1 2 1 0 1 2 0 0 1 0 0 0 0 0 0 0 0 0 2 0 1 1 3 0 0 1 1 1 1 2 2 1 1 1 2 1 1 2 1 1 1 2 0 1 1 0 0 0 0 0 0 0 0 0 3 1 0 1 1 3 2 1 1 2 2 1 3 0 1 1 2 1 x1 x2 x3 x4 x5 x6 x7 x8 x9 = - -7 -6 -6 -13 -14 -10 -11 -8 -8 The solution proceeds as in Example 10.7, with a matrix of the form: B1 C1 0 0 A 3 B3 -0.5 0.5 0.5 = 0 -1 1 B1 0.5 0.5 - 0.5 X1 X 3 F1 = - F2 F3 -0.5 C1 = 1.5 2 1 0 1 which replaces C1 0 A 2 B2 C2 X 2 B1 -1 -1 -1 - 2 B1 -1 -2.5 F1 = 0 -3.5 which replaces F1 1 0 and I = 0 0 1 0 0 0 which replaces B1 1 Exercise 10.12 (continued) Analysis: (continued) 2 B2 - A 2 C1 = 0 -0.5 B2 - A 2 C1 -1 2 1 1 0.5 1 0 0 1 - 0.5 1 -3 3 which replaces C2 Upon inversion, B2 - A 2 C1 -1 0.5 = 0 -1 1 0 0 1 0.25 -1.5 C2 = -7 1.25 - 0.25 - 3.5 F2 - A 2 F1 = -7 -6.5 B2 - A 2 C1 -1 which replaces F2 3.5 F2 = 0 -7 2.25 0 0 0 0.75 0 1.25 which replaces F2 0 0 0 1.5 8 1.5 Upon inversion, B3 - A 3 C2 1 B2 is replaced by 0 0 0 1 0 0 0 1 -1 A 2 is replaced by 0 0 -1.75 B3 - A 3 C2 = -1 -1.25 -2.75 F3 - A 3 F2 = -9 -2.25 B3 - A 3 C2 -1 1 = 9.25 -10 -1.25 0.75 6 12.5 8.5 - 0.75 1.25 0.75 which replaces F3 0.13514 x7 which replaces F3 = - x8 x9 x4 which replaces F2 = - x5 x6 x1 which replaces F1 = - x2 x3 F3 = -1.13514 -1.10811 - 0.45947 F2 - C2 F3 = - 2.67567 - 2.08110 2.62161 F1 - C1 F2 = - 3.72971 - 0.13513 Exercise 10.12 (continued) Analysis: (continued) Therefore, the solution to the 9 equations is: x1 = -2.62161 x2 = 3.72971 x3 = 0.13513 x4 = 0.45947 x5 = 2.67567 x6 = 2.08110 x7 = -0.13514 x8 = 1.13514 x9 = 1.10811 Subject: Given: Exercise 10.13 Matrix structure when equations are ordered by type. Mesh equations (10-58) to (10-60). Find: Whether matrix structure is block tridiagonal. Analysis: Many matrix structures are possible depending on the order of the equations and the order of the variables. Referring to Eqs. (10-61) to (10-64), try the following orders for a three component system, where j refers to the stage number: F = [ Hj , M1,j , M2,j , M3,j , E1,j , E2,j , E3,j ] X = [ 1,j , 2,j , 3,j , Tj , l1,j , l2,j , l3,j ] For a three-stage system, the incidence matrix is on the next page, where x denotes that the variable for that column appears in the equation for that row. It is seen in the incidence diagram that the matrix is not block tridiagonal or banded. If the order of the equations is maintained, it is not possible to permute the variable columns to get a blockbanded structure. Analysis: (continued) Exercise 10.13 (continued) Incidence Matrix 1,1 1,2 1,3 2 ,1 2 ,2 2 ,3 3,1 3,2 3,3 T1 T2 T3 l1,1 l1,2 l1,3 l2 ,1 l2 ,2 l2 ,3 l3,1 l3,2 l3,3 H1 H2 H3 M1,1 M1,2 M1,3 M 2,1 M 2,2 M 2,3 M 3,1 M 3,2 M 3,3 E1,1 E1,2 E1,3 E 2,1 E 2,2 E 2,3 E 3,1 E 3,2 E 3,3 x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x x Exercise 10.14 Subject: Obtaining derivatives of K-values for use in the Jacobian of the simultaneouscorrection method. Given: Chao-Seader correlation for the K-value based on Eq. (2) in Table 2.3. Ki , j Tj , Find: Analytical derivatives for: Ki , j i ,k , and Ki , j li , k . where, I,k = yi,kVk and l = xi,kLk and component molar flow rates. Analysis: From Eq. (2) in Table 2.3, Ki = iL iL iV From Eqs. (2-64) and (2-62), (1) v iL = exp iL i - RT xi viL x j viL C j =1 2 j j (2) From Eq. (2-61), where for a given component, i , viL and i are constants, i = C (3) j =1 It is important to distinguish between the symbols i,k (the molar vapor component flow rate) and viL (the component molar volume in the liquid phase = a constant). For the liquid-phase, pure-component, fugacity coefficient, the Chao-Seader equations are: (0) (1) log iL = log iL + i log iL (0) logiL = A0 + (4) Pri2 A8 + A9Tri - log Pri (1) logiL = A10 + A11Tri + ( A1 + A2Tri + A3Tri2 + A4Tri3 + Pri A5 + A6Tri + A7Tri2 + Tri ( ) ) (5) A12 + A13Tri3 + A14 Pri - 0.6 Tri ( ) (6) where, Tr = T/Tc Exercise 10.14 (continued) Analysis (continued) From Eqs. (2-56), (2-46), (2-47), (2-48), (2-49), and (2-50), iV = exp ZV - 1 Bi A A B B - ln ZV - B - 2 i - i ln 1 + B B A B ZV bP RT (7) (8) 3 2 ZV - ZV + ZV A - B - B 2 - AB = 0 A= aP R 2T 2 (9) B= (10) with a and b given Table 2.5 and Eqs. (2 - 49) and (2 - 50) Derivative Ki , j Tj : Ki , j Tj = iL iL + iL iL - iL 2 iL iV iV Tj iV Tj iV Tj C j =1 2 Using Eq. (1) with the chain rule, (11) iL From Eq. (2), =- Tj iL viL i - j j RTj2 (0 (1 iL log iL ) log iL) = 2.3026 iL + i Tj Tj Tj where from Eqs. (5) and (6), (0 log iL ) 1 A1 = - 2 + A2 + 2 A3Tri + 3 A4 Tri2 + Pri A6 + 2 A7 Tri + Pri2 A9 Tj Tci Tri (1 log iL) 1 A = A11 - 12 + 3 A13Tri2 Tj Tci Tri2 To obtain iV , we note that in Eq. (7), ZV, A, Ai, B, and Bi are functions of T. Tj Therefore, from Eq. (7), Exercise 10.14 (continued) Analysis: (continued) Bi ZV 1 Bi Bi B 1 + ZV - 1 - 2 - B T B T B T ZV - B iV = iV Tj 1 B ZV B - - T T A Ai A B B - 2 i ln 1 + - Ai T B T ZV Ai Bi A A A 2 B B 1 - - 2 2 i - i + A B T B A B T B B B ZV - T ZV T 1+ A BZV B ZV 2 Ai Bi - A B where, A A Ai A B B = -2.5 , = -2.5 i , =- , T T T T T T Bi B =- i T T ZV is obtained as follows by implicit integration from Eq. (8): T ZV Z ZV A B B B A - 2 ZV V + A - B - B 2 + ZV - - 2B -A -B =0 T T T T T T T T Solving, 2 3ZV ZV = T ZV 2 A B B B A Z -2.5 A + B + 2 B + 35 AB . - - 2B -A -B V T T T T T T T T T = 2 2 2 ZV - 3ZV - A - B - B 2 2 ZV - 3ZV - A - B - B 2 Substitution of the equations below Eq. (11) into Eq. (11) gives an expression for Ki for any stage j . T Derivative Ki , j i , k : Using Eq. (1) with the chain rule, noting that only iV is a function of i and letting the subscript k = j, but dropping that subscript: Ki = - iL 2 iL iV i iV i (12) Exercise 10.14 (continued) Analysis: (continued) The mixture vapor fugacity coefficient, iV , is given by Eq. (7). In this equation only ZV, A, and B are functions of vapor composition. Also, ZV, is a function of A and B, which are given by Eqs. (9) and (10), where, from Eqs. (2-49) and (2-50), after substituting yi = C C n =1 C i V = i C m =1 : m a= r =1 r n ar an C 2 and b= r =1 C r bn m m =1 m m =1 Therefore, A P ya 2 = 2 2 i i+ i R T V V B P bi 1 = - i RT V V C r =1 C n =1 n i yn ai an - 2 V C r =1 C n =1 yr yn a r a n (13) yr br (14) ZV is obtained as follows by implicit integration from Eq. (8): i 2 3ZV ZV Z ZV A B B B A - 2 ZV V + A - B - B 2 + ZV - - 2B -A -B =0 i i i i i i i i A B B B A - - 2B -A -B i i i i i 2 2 ZV - 3ZV - A - B - B 2 Solving, ZV = i ZV (15) From Eq. (7), Exercise 10.14 (continued) Analysis: (continued) Bi ZV 1 Bi Bi B 1 + ZV - 1 - 2 - B i B i B i ZV - B iV = iV i 1 B ZV B - - i i A Ai A Bi B - 2 ln 1 + - Ai i B i ZV Ai Bi A A A 2 B B 1 - - 2 2 i - i + A B i B A B i B B B ZV - i ZV i 1+ A BZV B ZV 2 Ai Bi - A B The final result is obtained by substituting equations below (13), (14), and (15) into this equation, followed by substitution of the result into Equation (12). Derivative Ki , j li ,k : Using Eq. (1) with the chain rule, noting that only iL is a function of li and letting the subscript k = j, but dropping that subscript: Ki iL iL = li iV li In Eq. (2), using: m = (16) xmvmL lmvmL l v = = m CmL vL Lv L v L ln n =1 C 2 Eq. (1) becomes: v iL = exp iL i - RT m =1 lmvmL m C n =1 vL ln Taking the derivative of this equation with respect to li gives: Exercise 10.14 (continued) Analysis: (continued) C iL 2viL = iL i - li RT m =1 lmvmL m C n =1 C viL i vL C n =1 - m =1 lmvmL m C n =1 2 vL ln ln C vL ln = iL C 2viL i - m m RT m =1 i i - xi L m =1 m m L Substitution of this equation into Eq. (16) gives the final result Exercise 10.15 Subject: Given: Development of a partial simultaneous correction method. MESH equations (10-1) to (10-6). Find: (a) The two indexed equations for solving for temperatures and vapor flow rates. (b) Truncated Taylor series expansions for the two indexed equations. (c) Order of equations and variables to permit an efficient solution procedure. Analysis: Instead of computing a new set of total vapor flow rates from the energy balance equations as in the bubble-point method and a new set of temperatures from the energy balance equations as in the sum-rates method, a good procedure to simultaneously compute a new set of total vapor flow rates and a new set of temperatures from a combination of the energy balance, sum rates, and equilibrium equations. This procedure is as follows: (a) Combine the energy balance equation (10-5) with the total mole balance (10-6) to eliminate variables in Lj so as to obtain the following equation in terms of Vj and Tj , assuming that mole fraction compositions have just been calculated: H j = H j V j , Tj = V j + V j +1 + j m =1 j -1 m =1 Fm - U m - Wm - V1 hL j -1 + V j +1hV j+1 + Fj hFj - (1) Fm - U m - Wm - V1 + U j hL j - V j + Wj hV j - Q j = 0 Instead of using the sum-rates equation (10-33) in terms of total liquid flow rates, use the following similar equation in terms of the total vapor flow rates: V j( k +1) = V j( k ) C i =1 yi , j Combining this equation with equation (10 - 2) V j( k +1) = V j( k ) C i =1 Ki , j xi , j Writing this equation as a sums function, S j , gives: S j = S j V j , Tj = V j( k +1) - V j( k ) C i =1 Ki , j xi , j = 0 (2) Equations (1) and (2) are the indexed equations for computing new sets of Vj and Tj , where in Eq. (1), Vj = Vj(k+1) and all Tj implicit in the K-values and enthalpies are Tj(k+1). Exercise 10.15 (continued) Analysis: (continued) (b) The truncated Taylor's series expansions for Eqs. (1) and (2) are: 0 = H (j k ) + H j Tj -1 S j V j (k ) Tj -1 + H j Tj (k ) (k ) Tj + H j Tj +1 (k ) Tj +1 + H j V j (k ) V j + H j V j +1 (k ) V j +1 (3) (k ) 0= S (k ) j + V j + S j Tj Tj (4) The derivatives for Eq. (3) are: H j Tj -1 H j Tj H j Tj +1 H j V j H j V j +1 = Vj + j -1 m =1 j m =1 Fm - U m - Wm - V1 Fm - U m - Wm - V1 hL j -1 Tj -1 hL j Tj - V j + Wj hV j Tj = V j +1 + = V j +1 hV j+1 Tj +1 = hL j -1 - hV j = hV j+1 - hL j The derivatives for Eq. (4) are: S j V j S j Tj =1 = -V j( k ) C i =1 Ki , j Tj xi , j Exercise 10.15 (continued) Analysis: (continued) (c) Consider a simple case of 3 equilibrium stages. The set of 3 equations each for Eq. (3) and Eq. (4), will form a block tridiagonal matrix structure if the equations and correction variables are ordered as shown below. More details on this method are given by J. F. Tomich, AIChE Journal, 16, 229 (1970). The following is the incidence matrix for the Jacobian matrix: V1 S1 H1 S2 H2 S3 H3 x x T1 x x x V2 x x x T2 x x x x V3 T3 x x x x x x Exercise 10.16 Subject: Effect of interlinking streams of a thermally coupled distillation system on the Jacobian matrix structure. Given: Thermally coupled system of Fig. 10.31. Solution of the equations by the NewtonRaphson method. Find: Whether the matrix equation retains a block tridiagonal structure. Analysis: Consider the system shown below, where the stages in the first column are numbered from 12 to 15 and the stages in the second column are numbered from 1 to 11 and from 16 to 19. The component material balance equations, similar to Eq. (10-58, around stages 4, 12, 11, and 16, respectively are: M i ,4 = li ,4 (1 + s4 ) + i ,4 - li ,12 - li ,3 - i ,5 = 0 M i ,12 = li ,12 + i ,12 - li ,4 s4 - i ,13 = 0 M i ,11 = li ,11 + i ,11 - li ,10 - i ,16 = 0 M i ,16 = li ,16 + i ,16 (1 + S16 ) - li ,15 - li ,11 - i ,17 = 0 Note that: Stage 4 includes flow rates for stages 3, 4, 5, and 12. Stage 12 includes flow rates for stages 4, 12, and 13. Stage 11 includes flow rates for stages 10, 11, and 16. Stage 16 includes flow rates for stages 11, 15, 16, and 17. The same stages appear in the energy balance equations. Therefore, the matrix structure has the diverse A and C blocks as shown in the diagram below. The matrix is not block tridiagonal. Exercise 10.16 (continued) Analysis: (continued) Exercise 10.16 (continued) Analysis: (continued) Matrix Structure Exercise 10.17 Subject: Given: Find: Effect of changing variable and equation ordering. Ordering equations (10-63) and (10-64). Consequences of changing ordering in the above two equations. With the variable ordering v, T, l, as seen in Eq. (10-68), the non-zero variables in Analysis: the A j submatrix are pushed to the right, close to the B j submatrix, while the non-zero variables in the C j submatrix are pushed to the left, close to the B j submatrix. Thus, the nonzero variables are arranged across the block row of these three submatrices in the most compact manner. If the ordering is changed to l, v, T , the nonzero variables are spread out over the entire width of the row of the three submatrices. That is, the non-zero variables in the A j submatrix are now pushed to the left, away from the B j submatrix, while the non-zero variables in the C j submatrix are pushed to the right, away from the B j submatrix. This is the least compact arrangement. If the ordering of the equations is changed from H, M, E to E, M, H, there is little change. Thus, the orderings given by Eqs. (10-63) and (10-64) appear to as efficient or more efficient than the other orderings given above. Exercise 10.18 Subject: Detailed method for determining the scalar multiplier, Sb , in (10-104). Given: Inside-out method procedure in Section 10.5. Find: Details of the subject method. Analysis: From Eq. (10-104), when uncorrected, Si,j = I,j Sb,j . This product appears in the component material balance equations (10-83). Using initial estimates from steps 1-7 of the Initialization Procedure, values of li,j are computed from step 9. If the sum of li,j for the Nth stage, i.e. B= is not the specified B, then use, C i =1 li , N Si , j = Sb i , j Sb , j where Sb is determined by a search, e.g. a Fibonacci search to make B = specified B. Typically, the search might be over the range of 0.75 to 1.25. Exercise 10.19 Subject: Error function for the Inside-out method. Given: Error function for the simultaneous correction method. Find: Error function Analysis: Rewrite Eqs. (10-82) and (10-84) in residual form: Ei , j = i , j - i , j Sb , j li , j = 0 M i , j = Li , j -1 - RLj + i , j ..... li , j + i , j .... li , j +1 + f i , j = 0 H j = hL j RLj L j +.........- hF j Fj - Q j = 0 Then, following Eq. (10-75), = N j =1 Hj scale factor 2 + C i -1 (M ) + (E ) 2 i, j i, j 2 where the scale factor is a nominal heat of vaporization, some hV j - hL j Subject: Distillation of a light normal paraffin hydrocarbon mixture by a rigorous equilibrium-stage method. Given: Feed of composition below as a bubble-point liquid at column pressure of 250 psia. Column has a partial condenser, partial reboiler, and 15 equilibrium stages in the column with feed entering the middle stage. Distillate rate = 23.0 lbmol/h and reflux rate = 150.0 lbmol/h. Soave-Redlich-Kwong equation of state for K-values and enthalpies. Find: Product compositions. Stage temperatures, vapor and liquid flow rates, and compositions. Reboiler and condenser duties. Analysis: The calculations were made with the TOWER model (Inside-out method) of the Chemcad program, with the following specifications: Partial condenser Top pressure = 250 psia No condenser or stage pressure drop. Therefore, bottom pressure also = 250 psia Number of stages = 17 (includes partial condenser and partial reboiler) Feed stage = 9 from the top At the top, reflux rate = 150 lbmol/h At the bottom, bottoms rate = B = F - D = 100 - 23 = 77 lbmol/h Estimated distillate rate = 23 lbmol/h (actual rate) Estimated reflux rate = 150 lbmol/h (actual rate) Estimated temperature of stage 1 = 120oF Estimated temperature of bottom stage = 300oF Estimated temperature of stage 2 = 140oF Although the feed was given as a bubble-point liquid at 250 psia and 213.9oF, the SRK equation of state gave a bubble-point temperature of 210.7oF, which was the value used. Based on the feed composition below, the distillate rate of 23 lbmol/h corresponded to a desired split between propane and n-butane. The calculations converged rapidly in 3 outer-loop iterations with results given below and on the following pages. Condenser duty = 862,300 Btu/h Reboiler duty = 1,119,800 Btu/h Top temperature = 115.8oF compared to assumed value of 120oF Bottom temperature = 264.2oF compared to assumed value of 300oF Stage 2 temperature = 125.4oF compared to assumed value of 140oF Exercise 10.20 Exercise 10.20 (continued) Analysis: (continued) Analysis: (continued) Exercise 10.20 (continued) Exercise 10.21 Subject: Optimal feed-stage location for separation of Exercise 10.20 Given: Conditions given in Exercise 10.20 except that the optimal feed location is to be determined rather than using the middle stage (stage 9). Find: Optimal feed stage, assuming that this corresponds to the stage that gives the best separation between propane and n-butane. Analysis: The calculations were made with the TOWER model (Inside-out method) of the Chemcad program, in exactly the same way as with Exercise 10.20 except that several runs were made with different feed stage locations. The results are as follows for the flow rate of propane (the light key) in the bottoms: Feed stage location (from top) 6 7 8 9 (Exercise 10.20) 10 11 lbmol/h C3 in bottoms 0.5258 0.4743 0.4730 0.5165 0.6138 0.7769 The optimal feed stage is seen to be stage 8, with stage 9 giving almost the same result. Exercise 10.22 Subject: Distillation with a vapor sidestream Given: Conditions given in Exercise 10.20 except that a vapor side stream is withdrawn from the fourth stage (stage 13 from the top) from the bottom at a rate of 37.0 lbmol/h in an attempt to obtain a nC4-rich product. Therefore, the remaining specifications are still a feed of composition below as a bubble-point liquid at column pressure of 250 psia. Column has a partial condenser, partial reboiler, and 15 equilibrium stages in the column with feed entering the middle stage. Distillate rate = 23.0 lbmol/h and reflux rate = 150.0 lbmol/h. Soave-Redlich-Kwong equation of state for K-values and enthalpies. Find: Product compositions. Stage temperatures, vapor and liquid flow rates, and compositions. Reboiler and condenser duties. Analysis: The calculations were made with the TOWER model (Inside-out method) of the Chemcad program, with the following specifications: Partial condenser Top pressure = 250 psia No condenser or stage pressure drop. Therefore, bottom pressure also = 250 psia Number of stages = 17 (includes partial condenser and partial reboiler) Feed stage = 9 from the top Sidestream from stage 13 from the top, as a vapor at a flow rate of 37 lbmol/h At the top, reflux rate = 150 lbmol/h At the bottom, bottoms rate = B = F - D - S = 100 - 23 - 37 = 40 lbmol/h Estimated distillate rate = 23 lbmol/h (actual rate) Estimated reflux rate = 150 lbmol/h (actual rate) Estimated temperature of stage 1 = 120oF Estimated temperature of bottom stage = 300oF Estimated temperature of stage 2 = 140oF Although the feed was given as a bubble-point liquid at 250 psia and 213.9oF, the SRK equation of state gave a bubble-point temperature of 210.7oF, which was the value used. Based on the feed composition below, the distillate rate of 23 lbmol/h corresponded to a desired split between propane and n-butane. The calculations converged rapidly in 3 outer-loop iterations with results given below and on the following pages. The vapor side stream contains 69.5 mol% nC4. Condenser duty = 914,000 Btu/h Reboiler duty = 1,440,000 Btu/h Top temperature = 124.4oF compared to assumed value of 120oF Bottom temperature = 290.5oF compared to assumed value of 300oF Stage 2 temperature = 140.9oF compared to assumed value of 140oF Analysis: (continued) Exercise 10.22 (continued) Analysis: (continued) Exercise 10.22 (continued) Exercise 10.23 Subject: Distillation with an intercooler and an interheater Given: Conditions given in Exercise 10.20 except that an intercooler with a duty of 200,000 Btu/h is placed on the fourth stage from the top (Stage 5 counting the partial condenser), and an interheater with a duty of 300,000 Btu/h is placed on the fourth stage from the bottom (stage 13 from the top counting the partial condenser). Therefore, the remaining specifications are still a feed of composition below as a bubble-point liquid at column pressure of 250 psia. Column has a partial condenser, partial reboiler, and 15 equilibrium stages in the column with feed entering the middle stage. Distillate rate = 23.0 lbmol/h and reflux rate = 150.0 lbmol/h. Soave-RedlichKwong equation of state for K-values and enthalpies. Find: Product compositions. Stage temperatures, vapor and liquid flow rates, and compositions. Reboiler and condenser duties. Analysis: The calculations were made with the TOWER model (Inside-out method) of the Chemcad program, with the following specifications, where TOWER treats intercooler and interheater stages as feed stages with heat transferred out as negative and heat transferred in as positive: Partial condenser and Top pressure = 250 psia No condenser or stage pressure drop. Therefore, bottom pressure also = 250 psia Number of stages = 17 (includes partial condenser and partial reboiler) Feed stage one = 5 from the top for the intercooler Feed stage two = 9 from the top for the feed to the column Feed stage three = 13 from the top for the interheater At the top, reflux rate = 150 lbmol/h At the bottom, bottoms rate = B = F - D = 100 - 23 = 77 lbmol/h Estimated distillate rate = 23 lbmol/h (actual rate) Estimated reflux rate = 150 lbmol/h (actual rate) Estimated temperature of stage 1 = 120oF Estimated temperature of bottom stage = 300oF Estimated temperature of stage 2 = 140oF Although the feed is given as a bubble-point liquid at 250 psia and 213.9oF, the SRK equation of state gave a bubble-point temperature of 210.7oF, which was the value used. Based on the feed composition below, the distillate rate of 23 lbmol/h corresponded to a desired split between propane and n-butane. The calculations converged rapidly in 3 outer-loop iterations with results given below and on the following pages: Condenser duty = 856,200 Btu/h Reboiler duty = 1,014,000 Btu/h o Top temperature = 115 F compared to assumed value of 120oF Bottom temperature = 264.4oF compared to assumed value of 300oF Stage 2 temperature = 123.7oF compared to assumed value of 140oF Analysis: (continued) Exercise 10.23 (continued) Analysis: (continued) Exercise 10.23 (continued) Exercise 10.24 Subject: Determination of optimal feed locations for two feeds of light normal paraffin hydrocarbon mixtures by a rigorous equilibrium-stage method. Given: Feeds of compositions below as a bubble-point liquids at a constant column pressure of 250 psia. Column has a partial condenser, partial reboiler, and 30 equilibrium stages in the column. Distillate rate = 36.0 lbmol/h and reflux rate = 150.0 lbmol/h. Peng-Robinson equation of state for K-values and enthalpies. Suggested feed locations are stage 15 (stage 17 from the top counting the partial condenser), and stage 6 (stage 26 from the top counting the partial condenser) from the bottom for the heaviest feed, F2. Peng-Robinson equation of state for Kvalues and enthalpies. Find: Optimal feed-stage locations. Analysis: The calculations were made with the TOWER model (Inside-out method) of the Chemcad program, with the following specifications: Partial condenser Top pressure = 250 psia No condenser or stage pressure drop. Therefore, bottom pressure also = 250 psia Number of stages = 32 (includes partial condenser and partial reboiler) Feed one initially to stage = 17 from the top Feed two initially to stage = 26 from the top At the top, reflux rate = 150 lbmol/h At the bottom, bottoms rate = B = F - D = 100 - 36 = 64 lbmol/h Estimated distillate rate = 36 lbmol/h (actual rate) Estimated reflux rate = 150 lbmol/h (actual rate) Estimated temperature of stage 1 = 120oF Estimated temperature of bottom stage = 250oF Estimated temperature of stage 2 = 140oF Based on the feed compositions below, the distillate rate of 36 lbmol/h corresponds to a desired split between propane and n-butane. The calculations converged rapidly in 3 outer-loop iterations with results given below and on the following pages. The separation with feed stages at 12 and 15 from the top is much better than the separation with feed stages at 17 and 26. For the better stage locations: Top temperature = 118.3oF compared to assumed value of 120oF Bottom temperature = 252.8oF compared to assumed value of 300oF Analysis: (continued) Exercise 10.24 (continued) Exercise 10.25 Subject: Distillation of a hydrocarbon mixture and comparison with solution in Perry's Handbook. Given: Bubble-point liquid feed of composition given below. Column has partial condenser, partial reboiler, and 9 equilibrium stages in the column, with operation at 120 psia. Reflux rate = 126.1 lbmol/h and vapor distillate rate = 48.9 lbmol/h. Grayson-Streed correlation for K-values and Redlich-Kwong equation of state for enthalpies. Find: Product compositions, stage temperatures, interstage flow rates and compositions, reboiler duty, and condenser duty. Analysis: The calculations were made with the TOWER model (Inside-out method) of the Chemcad program, with the following specifications: Partial condenser Top pressure = 120 psia No condenser or stage pressure drop. Therefore, bottom pressure also = 120 psia Number of stages = 11 (includes partial condenser and partial reboiler) Feed stage = 6 from the top At the top, reflux rate = 126.1 lbmol/h At the bottom, bottoms rate = B = F - D = 100 - 48.9 = 51.1 lbmol/h Estimated distillate rate = 48.9 lbmol/h (actual rate) Estimated reflux rate = 126.1 lbmol/h (actual rate) Estimated temperature of stage 1 = 120oF Estimated temperature of bottom stage = 250oF Estimated temperature of stage 2 = 140oF The feed is given as a bubble-point liquid at 120 psia. Using the Grayson-Streed correlation, the temperature is computed to be 180.8oF. Based on the feed composition below, the distillate rate of 48.9 lbmol/h corresponds to a desired split between n-butane and isopentane. The calculations converged rapidly in 3 outer-loop iterations with results given below and on the following pages. Condenser duty = 1,002,600 Btu/h Reboiler duty = 1,427,000 Btu/h Top temperature = 161.4oF compared to assumed value of 120oF Bottom temperature = 230.0oF compared to assumed value of 250oF Stage 2 temperature = 173.9oF compared to assumed value of 140oF Analysis: (continued) Exercise 10.25 (continued) Analysis: (continued) Exercise 10.25 (continued) Exercise 10.25 (continued) Analysis: (continued) The above results for the product compositions are compared to the results in Perry's Handbook as follows: Component Propane Isobutane n-Butane Isopentane n-Pentane Total: Feed, lbmol/h 5.00 15.00 25.00 20.00 35.00 100.00 Perry's Handbook: Distillate, Bottoms, lbmol/h lbmol/h 5.00 0.00 14.82 0.18 23.60 1.40 3.21 16.79 2.30 32.70 48.93 51.07 This exercise: Distillate, Bottoms, lbmol/h lbmol/h 5.00 0.00 14.71 0.29 23.42 1.58 3.71 16.29 2.06 32.94 48.90 51.10 The example in Perry's Handbook is for a column with a total condenser, while this exercise calls for a partial condenser. Thus, this exercise has 11 instead of 10 equilibrium stages. Nevertheless, the solution from Perry's Handbook shows a slightly sharper separation. This is probably due to the differences in K-values. For example at a mid-column temperature of 210oF, the Grayson-Streed correlation predicts a relative volatility between n-butane and isopentane of 1.82, while the graphical correlation used in Perry's Handbook predicts a higher value of 1.94. Subject: Exercise 10.26 Distillation of an light alcohol mixture using UNIFAC equation. Given: Feed of 1000 kmol/h of a bubble-point mixture of 60 mol% methanol (M), 20 mol% ethanol (E), and 20 mol% n-propanol (P) is sent to the middle stage of a column with a total condenser, 3 equilibrium stages in the column, and a partial reboiler, and operating at 1 atm. The distillate rate = 600 kmol/h and the reflux rate = 2,000 kmol/h. Find: Product compositions, stage temperatures and vapor and liquid traffic, reboiler and condenser duties, and stage compositions for converged solution. Analysis: The calculations were made with the TOWER model (Inside-out method) of the Chemcad program, with the following specifications: Total condenser Top pressure = 1 atm No condenser or stage pressure drop. Therefore, bottom pressure also = 1 atm Number of stages = 5 (includes total condenser and partial reboiler) Feed stage = 3 from the top At the top, reflux rate = 2,000 kmol/h At the bottom, bottoms rate = B = F - D = 1000 - 600 = 400 kmol/h Estimated distillate rate = 600 kmol/h (actual rate) Estimated reflux rate = 2000 kmol/h (actual rate) Estimated temperature of stage 1 = 30oC Estimated temperature of bottom stage = 100oC Estimated temperature of stage 2 = 40oC The feed is given as a bubble-point liquid at 1 atm. The UNIFAC equation gives a bubble-point temperature of 70.1oC. Based on the feed composition, the distillate rate of 600 kmol/h corresponds to a desired split between methanol and ethanol. The enthalpies were computed by the Latent Heat model with a correction for the heat of solution based on liquid-phase activity coefficients from the UNIFAC equation. The calculations converged in 8 outer-loop iterations with results given below and on the following pages. Condenser duty = 93,592 MJ/h Reboiler duty = 93,841 MJ/h Top temperature = 70.1oC compared to assumed value of 30oC Bottom temperature = 79.2oC compared to assumed value of 100oC Stage 2 temperature = 67.3oC compared to assumed value of 40oC Exercise 10.26 (continued) Analysis: (continued) Exercise 10.26 (continued) Analysis: (continued) Subject: Distillation of a 4-component mixture with two sidestreams in an attempt to obtain four nearly pure products. Given: Feed of composition given below at 150oF and 25 psia. Column pressure of 20 psia at top and 25 psia at the bottom Column has a total condenser giving saturated liquid, a partial reboiler, and 28 equilibrium stages in the column with feed entering stage 14. Distillate rate = 14.08 lbmol/h and reflux ratio = 20. A liquid sidestream of 19.53 lbmol/h is withdrawn at stage 10 from the top, and a vapor sidestream of 24.78 lbmol/h is withdrawn at stage 24 from the top. Peng-Robinson equation of state for K-values and enthalpies. Find: Product compositions. Stage temperatures, vapor and liquid flow rates, and compositions. Reboiler and condenser duties. Analysis: The calculations were made with the TOWER model (Inside-out method) of the Chemcad program, with the following specifications: Total condenser Top pressure = 20 psia Total stage pressure drop = 5 psia. Therefore, bottom pressure = 25 psia Number of stages = 30 (includes total condenser and partial reboiler) Feed stage = 15 from the top Liquid sidestream from stage 11 from the top, at a flow rate of 19.53 lbmol/h Vapor sidestream from stage 25 from the top, at a flow rate of 24.78 lbmol/h At the top, reflux ratio = 20 At the bottom, bottoms rate = B = F - D - S1 - S2 = 98.33 - 14.08 - 19.53 - 24.78 = 39.94 lbmol/h Estimated distillate rate = 14.08 lbmol/h (actual rate) Estimated reflux rate = 280 lbmol/h (close to actual rate of 281.6 lbmol/h) Estimated temperature of stage 1 = 120oF Estimated temperature of bottom stage = 200oF Estimated temperature of stage 2 = 140oF Estimated 1st draw rate = 19.53 lbmol/h (actual rate) Estimated 2nd draw rate = 24.78 lbmol/h (actual rate) The feed condition at the specified temperature and pressure is computed to be just slightly greater than the bubble-point temperature. Based on the feed composition below, the distillate, sidestreams, and bottoms correspond to the flow rates of the four components in the feed. The calculations converged rapidly in 4 outer-loop iterations with results given below and on the following pages. Condenser duty = 2,834,600 Btu/h Reboiler duty = 3,493,700 Btu/h o Top temperature = 47.7 F compared to assumed value of 120oF Bottom temperature = 289.0oF compared to assumed value of 200oF Stage 2 temperature = 49.9oF compared to assumed value of 140oF Exercise 10.27 Analysis: (continued) A high reflux ratio is needed to minimize the effect of feed composition when attempting to obtain nearly pure sidestreams. Exercise 10.27 (continued) Analysis: (continued) Exercise 10.27 (continued) Exercise 10.28 Subject: Distillation of a hydrocarbon mixture given in Exercise 10.25. Given: Bubble-point liquid feed of composition given below. Column has a partial condenser and a partial reboiler, with operation at 120 psia. Recovery of light key, n-butane, to distillate = 98%. Recovery of heavy key, isopentane, to bottoms = 98%. Grayson-Streed correlation for Kvalues and Redlich-Kwong equation of state for enthalpies. Find: Product compositions, reflux ratio, numbers of equilibrium stages above and below the feed, stage temperatures, interstage flow rates and compositions, reboiler duty, and condenser duty. Analysis: The calculations were made with the TOWER model (Inside-out method). Initially, it was assumed that the same number of stages and feed stage location could be used as in Exercise 10.25. Also, the same reflux rate and distillate rate as specified for Exercise 10.25 were given for initial estimates. It was understood that the Chemcad TOWER model would adjust the reflux rate and distillate rate in an attempt to achieve the desired LK and HK recoveries. Thus, the initial specifications for the TOWER model of the Chemcad program were: Partial condenser Top pressure = 120 psia No condenser or stage pressure drop. Therefore, bottom pressure also = 120 psia Number of stages = 11 (includes partial condenser and partial reboiler) Feed stage = 6 from the top At the top, component recovery of n-butane (component 3) = 0.98 At the bottom, component recovery of isopentane (component 4) = 0.98 Estimated distillate rate = 48.9 lbmol/h (actual rate) Estimated reflux rate = 126.1 lbmol/h (actual rate) Estimated temperature of stage 1 = 120oF Estimated temperature of bottom stage = 250oF Estimated temperature of stage 2 = 140oF The feed is given as a bubble-point liquid at 120 psia. Using the Grayson-Streed correlation, the temperature is computed to be 180.8oF. After 50 iterations, the calculations were not converged, even though the reflux rate was increased from 126.1 lbmol/h to 8000.4 lbmol/h. The 98% recoveries of the two key components could not be achieved. This indicated that the number of stages was probably below the minimum number. Analysis: (continued) Exercise 10.28 (continued) To obtain a better estimate of the stage requirements, the Fenske equation can be applied in the following manner, where the separation achieved in Exercise 10.25 for 11 stages is ratioed to the separation desired here. That separation in Exercise 10.25 is 93.66% recovery of nC4 and 81.4% recovery of iC5. Thus, applying Eq. (9-12) in this manner: log N= log 98 2 93.7 6.3 98 2 81.5 8.5 (11) = 17. 3 equilibrium stages With this result, it was decided to try 19 stages (including the partial condenser and partial reboiler) with the feed to stage 10. Otherwise, the specifications for the TOWER model were unchanged from above. This time convergence was achieved rapidly in just 2 iterations. The reflux rate was increased by TOWER from 126.1 to 161.0 lbmol/h and the distillate rate was reduced from 48.9 lbmol/h to 44.93 lbmol/h. Other results were as follows: Condenser duty = 1,195,400 Btu/h Reboiler duty = 1,586,600 Btu/h Top temperature = 147.9oF compared to assumed value of 120oF Bottom temperature = 232.3oF compared to assumed value of 250oF Stage 2 temperature = 152.9oF compared to assumed value of 140oF Exercise 10.28 (continued) Analysis: (continued) Analysis: (continued) Exercise 10.28 (continued) Exercise 10.29 Subject: Distillation of a light hydrocarbon mixture with a specified split. Given: Saturated liquid feed at 125 psia of 200 lbmol/h of 5 mol% iC4, 20 mol% nC4, 35 mol% iC5, and 40 mol% nC5. Column is equipped with a total condenser and a partial reboiler, operating at 125 psia. Recovery of nC4 to the distillate = 95%. Recovery of iC5 to the bottoms = 95%. Number of equilibrium stages = 2 times that of the minimum stages by the Fenske equation. SRK for K-values and enthalpies. Find: A suitable design. Assume this consists of products, profiles, and stage compositions. Analysis: The Shortcut model of Chemcad computes a minimum number of stages = 10.7 and a minimum reflux ratio of 3.36. Therefore, select 2(10.7) = 21.4 or 22 stages (including the partial reboiler). Estimate the reflux ratio at 1.3(3.36) = 4.4. For the rigorous TOWER model of chemcad, use the following specifications: Total condenser Top pressure = 125 psia No condenser or stage pressure drop. Therefore, bottom pressure also = 125 psia Number of stages = 23 (includes total condenser and partial reboiler) Feed stage = 12 from the top At the top, component recovery of n-butane (component 2) = 0.95 At the bottom, component recovery of isopentane (component 3) = 0.95 Estimated distillate rate = 50 lbmol/h (actual rate) Estimated reflux rate = 50(3.36) = 168 lbmol/h (actual rate) Estimated temperature of stage 1 = 120oF Estimated temperature of bottom stage = 250oF Estimated temperature of stage 2 = 140oF The feed is given as a bubble-point liquid at 125 psia. Using the SRK correlation, the temperature is computed to be 210oF. After 4 iterations, the calculations are converged, while increasing the reflux rate to 267.9 lbmol/h, and the distillate rate to 52.11 lbmol/h, at which the 95% recoveries of the two key components are achieved. The computed results are as follows on this page and the next pages. Condenser duty = 2,517,800 Btu/h Reboiler duty = 2,582,500 Btu/h Top temperature = 160.8oF compared to assumed value of 120oF Bottom temperature = 233.2oF compared to assumed value of 250oF Stage 2 temperature = 165.8oF compared to assumed value of 140oF Exercise 10.29 (continued) Analysis: (continued) Analysis: (continued) Exercise 10.29 (continued) Exercise 10.30 Subject: Design of a depropanizer distillation column. Given: Feed at column average pressure of 315 psia and 66 mol% vaporized. Composition of feed in lbmol/h is as follows: 26 C1, 9 C2, 25 C3, 17 nC4, 11 nC5, 12 nC6. Separation is between C3 (0.4 lbmol/h in bottoms) and nC4 (0.3 lbmol/h in distillate). Cooling water available at 65oF and steam available at 315 psia. Total column pressure drop = 2 psi. Assumptions: Partial reboiler. Grayson-Streed correlation for K-values and Redlich-Kwong equation for enthalpies. Find: (a) Type condenser. (b) Feed temperature. K-values and relative volatilities at feed temperature and pressure. (c) Reflux ratio for 1.3 times minimum reflux ratio. Number of equilibrium stages in rectifying and stripping sections. (d) Separation. Separation for reflux ratio = 1.5 times minimum, 15 theoretical stagess, feed stage at stage 9. (e) Temperature and compositions at each stage for Part (c). Effect on these results if reflux and stages of Part (d) are used. Analysis: Use the Flash, Shortcut, and TOWER models of Chemcad to make the calculations. (a) Assume a pressure at the top of 314 psia, with 316 psia at the bottom and 315 psia at the feed location. Using the Flash model, run a bubble-point temperature calculation of the assumed distillate in lbmol/h of 26 C1, 9 C2, 24.6 C3, and 0.3 nC4 at 314 psia for both 0% (total condenser case) and 100% vaporized (partial condenser case). For a total condenser, the distillate temperature is found to be -100oF, which is infeasible. For a partial condenser, the distillate temperature is 73.1oF, which is possible, but very close to the cooling water temperature of 65oF. A bubble point calculation on the assumed bottoms in lbmol/h of 0.4 C3, 16.7 nC4, 11 nC5, and 12 nC6, at 316 psia is 301.9oF. The saturation temperature of steam at 315 psia is 421oF. Therefore, the steam is adequate, but may have to be throttled to a lower pressure to avoid film boiling in the reboiler. (b) When the feed is flashed at 315 psia to 66 mol% vapor, the temperature = 183.3oF. The computed K-values and relative volatilities (relative to the HK, nC4) are: Component K-value referred to nC4 Methane 9.250 15.29 Ethane 3.081 5.09 Propane, LK 1.405 2.32 n-Butane, HK 0.605 1.00 n-Pentane 0.274 0.453 n-Hexane 0.126 0.208 Analysis (continued) Exercise 10.30 (continued) (c) The recovery of the LK, C3, to the distillate = 24.6/25 = 0.984. The recovery of the HK, nC4, to the distillate = 0.3/17 = 0.01765. Using the Shortcut (FUG) model of Chemcad, the following results are obtained: Minimum equilibrium stages = 10.5 Minimum reflux ratio = 0.87 Actual reflux ratio for 1.3 times minimum = 1.131 Number of equilibrium stages = 23.5 Feed stage from top = 11.5 Therefore, use, say, 10 theoretical plates in rectifying section plus a partial condenser, and 12 theoretical plates in the stripping section plus a partial reboiler. (d) For the rigorous TOWER model of chemcad, first use the following specifications for the case of a reflux ratio = 1.3 times minimum: Partial condenser Top pressure = 314 psia No condenser pressure drop. Column pressure drop = 2 psi Therefore, bottom pressure = 316 psia Number of stages = 24 (includes partial condenser and partial reboiler) Feed stage = 12 from the top At the top, reflux ratio = 1.131 At the bottom, bottoms flow rate = 40.1 lbmol/h Estimated distillate rate = 59.9 lbmol/h (actual rate) Estimated reflux rate = 59.9(1.131) = 67.75 lbmol/h (actual rate) Estimated temperature of stage 1 = 73oF Estimated temperature of bottom stage = 302oF Estimated temperature of stage 2 = 83oF The TOWER model converged in 4 iterations, with the following results, including the temperatures and concentrations at each stage, as requested in Part (e): Condenser duty = 417,600 Btu/h Reboiler duty = 482,500 Btu/h Top temperature = 72.9oF compared to assumed value of 73oF Bottom temperature = 302oF compared to assumed value of 302oF Stage 2 temperature = 105.2oF compared to assumed value of 83oF Note that the desired split of the two key components is achieved with: 0.2747 lbmol/h of nC4 (HK) in the distillate compared to a specification of 0.3 lbmol/h 0.3747 lbmol/h of C3 (LK) in the bottoms compared to a specification of 0.4 lbmol/h Analysis: (d) (continued) Exercise 10.30 (continued) Analysis: (d) (continued) Exercise 10.30 (continued) Analysis: (d) (continued) Exercise 10.30 (continued) For the rigorous TOWER model of chemcad, now use the following specifications for the case of a reflux ratio = 1.5 times minimum, with 15 theoretical plates and feed to plate 9 (assume from the top): Partial condenser Top pressure = 314 psia No condenser pressure drop. Column pressure drop = 2 psi Therefore, bottom pressure = 316 psia Number of stages = 17 (includes partial condenser and partial reboiler) Feed stage = 10 from the top At the top, reflux ratio = 1.5(0.87) = 1.305 At the bottom, bottoms flow rate = 40.1 lbmol/h Estimated distillate rate = 59.9 lbmol/h (actual rate) Estimated reflux rate = 59.9(1.305) = 78.17 lbmol/h (actual rate) Estimated temperature of stage 1 = 73oF Estimated temperature of bottom stage = 302oF Estimated temperature of stage 2 = 83oF The TOWER model converged in 3 iterations, with the following results for the products: Note that the desired split of the two key components is not achieved: 0.6269 lbmol/h of nC4 (HK) in the distillate compared to a specification of 0.3 lbmol/h 0.7256 lbmol/h of C3 (LK) in the bottoms compared to a specification of 0.4 lbmol/h Analysis: (e) (continued) The calculations with the TOWER model of Chemcad are repeated here for the case of a reflux ratio = 1.3 times the minimum, with 24 total stages (including the partial condenser and partial reboiler) and a variable feed stage in an attempt to make the best separation. We compare the results on the basis of the lbmol/h of nC4 (HK) in the distillate product as follows: Feed stage from top 7 8 9 10 11 12 13 14 15 lbmol/h of nC4 in distillate 0.509 0.364 0.276 0.244 0.246 0.275 0.343 0.439 0.559 Exercise 10.30 (continued) From these results, it is seen that the optimal feed location is stage 11 (stage 10 in the column), with stage 10 giving almost the same separation. For the case of a reflux ratio of 1.5 time the minimum, the desired separation could not be achieved for any feed stage location. Exercise 10.31 Subject: Separation of toluene from biphenyl (diphenyl) by distillation Given: The following feed at 264oF and 37.1 psia, and the desired products: lbmol/h: Component Feed Distillate Bottoms Benzene 3.4 3.4 0.0 Toluene 84.6 82.5 2.1 Biphenyl 5.1 1.0 4.1 Total 93.1 86.9 6.2 Reflux ratio = 1.3 times minimum with total condenser and partial reboiler. Pressure at top = 36 psia. Pressure at bottom = 38.2 psia. Assumptions: SRK equation of state for K-values and enthalpies. Find: (a) Actual reflux, number of equilibrium stages, and feed stage location. (b) Separation for a D/F = 86.9/93.1 = 0.9334 (c) Specified separation by adjusting reflux ratio in Part (b) Analysis: (a) The recovery of the LK, toluene, to the distillate = 82.5/84.6 = 0.9752. The recovery of the HK, biphenyl, to the distillate = 1.0/5.1 = 0.1980. Using the Shortcut (FUG) model of Chemcad, the following results are obtained: Minimum equilibrium stages = 2.5 Therefore, assume that a rectifying section is not needed. Try a reboiled stripper with 4 stages plus a partial reboiler. (b) For the rigorous TOWER model of chemcad, use the following specifications: No condenser Top pressure = 36 psia No condenser pressure drop. Column pressure drop = 2.2 psi Therefore, bottom pressure = 38.2 psia Number of stages = 5 (includes partial reboiler) Feed stage = top stage At the bottom, bottoms flow rate = 6.2 lbmol/h (corresponds to the specified D/F ratio. Estimated distillate rate = 86.9 lbmol/h (desired rate) Estimated reflux rate = feed rate = 93.1 lbmol/h) Estimated temperature of stage 1 = 300oF Estimated temperature of bottom stage = 400oF Estimated temperature of stage 2 = 320oF The TOWER model converged in 5 iterations, with the following results: Analysis: (b) (continued) Exercise 10.31 (continued) Reboiler duty = 1,379,300 Btu/h Note that the desired split of the two key components is better than desired with: 1.204 lbmol/h of toluene (LK) in the bottoms compared to a specification of 2.1 lbmol/h 0.106 lbmol/h of diphenyl (HK) in the distillate compared to a specification of 1.0 lbmol/h (c) Because the column has no reflux, the flow rate of the toluene in the bottoms was specified on the basis that the reboiler duty would be adjusted to obtained that flow rate. The TOWER model converged in 4 iterations with the following results, which did give the desired toluene rate in the bottoms, but much less than desired diphenyl in the distillate: Reboiler duty = 1,354,500 Btu/h Exercise 10.32 Subject: Comparison of two sequences of two distillation columns each for the separation of a mixture into three products. Given: Feed at 100oF and 480 psia and desired products with the following compositions in component flow rates: lbmol/h: Component Feed Product 1 Hydrogen 1.5 1.5 Methane 19.3 19.2 Benzene 262.8 1.3 Toluene 84.7 Biphenyl 5.1 Total: 373.4 22.0 Product 2 0.1 258.1 0.1 258.3 Product 3 3.4 84.6 5.1 93.1 In Sequence 1, the keys in Column 1 are methane and benzene, and the keys in Column 2 are benzene and toluene. For Sequence 2, the keys in Column 1 are benzene and toluene, and the keys in Column 2 are methane and benzene. Use partial condensers when the bulk of the methane is taken overhead and set condenser pressures to at least 20 psia, with distillate temperatures of about 130oF. Assumptions: SRK for K-values and enthalpies. Isobaric operation of columns. Find: Column pressures, type condensers, and reflux and stage requirements for columns in each sequence, using a reflux ratio of 1.3 times minimum reflux ratio. Analysis: First determine column pressures and type condensers, using Chemcad. Sequence 1, Column 1: Distillate is Product 1. The computed dew-point pressure at 130oF = 118 psia. Set top pressure to 118 psia with a partial condenser. Sequence 1, Column 2: Distillate is Product 2. The computed bubble-point pressure at 130oF = 8.4 psia. This is too low. Therefore, set top pressure to 20 psia with a total condenser. Distillate temperature at this pressure = 187oF. Sequence 2, Column 1: Distillate is the sum of Products 1 and 2. The computed dew-point pressure at 130oF = 6.7 psia. This is too low. Therefore, set top pressure to 20 psia with a partial condenser. Distillate temperature at this pressure = 190oF. Sequence 2, Column 2: Distillate is Product 1. Therefore, again set pressure to 118 psia with a partial condenser. Analysis: (continued) Exercise 10.32 (continued) Now, use the Shortcut model in Chemcad to estimate stage and reflux ratios using R = 1.3Rmin. Results are as follows, where recoveries are of the column feed to the distillate: Sequence 1 Sequence 1 Sequence 2 Sequence 2 Column 1 Column 2 Column 1 Column 2 Pressure, psia 118 20 20 118 Light key, LK Methane Benzene Benzene Methane Heavy key, HK Benzene Toluene Toluene Benzene Recovery of LK 0.9948 0.9870 0.9870 0.9948 Recovery of HK 0.00495 0.00118 0.00118 0.005012 Minimum stages 2.04 14.37 13.28 2.01 Min. reflux ratio 0.0685 1.188 0.930 5.628 Reflux ratio 0.0822 1.425 1.209 7.317 No. of stages 8.97 32.35 29.23 4.55 Feed stage from top 4.60 23.99 21.34 2.26 Now design the two sequences using the rigorous Tower model of Chemcad. For each column, the top specification is the flow rate of the LK in the distillate and the bottom specification is the flow rate of the HK in the bottoms. The stage specifications and estimated reflux and distillate rates are taken from the above shortcut results, with some modest changes. The rigorous Tower model then computes the required reflux ratio and distribution of nonkey components. The results are as follows, where the total stages includes the condenser (total or partial) and the reboiler. The feed stage is counted from the condenser as stage 1. Sequence 1 Sequence 1 Sequence 2 Sequence 2 Column 1 Column 2 Column 1 Column 2 Top pressure, psia 118 20 20 118 Pressure drop, psia 2 2 2 2 Total stages 3 30 30 5 Feed stage 2 22 21 2 Reflux ratio 2.2 1.363 1.101 0.0042 Condenser duty, Btu/h 892,100 8,122,800 4,074,800 1,900 Reboiler duty, Btu/h 4,227,300 6,135,800 8,843,100 2,748,100 T of bottoms, oF 347.4 258.8 258.8 404.6 Lbmol/h of toluene in Product 1 0.063 0.0002 Sequence 2 requires an expensive two-stage compression system to take the vapor distillate from Column 1 at 20 psia and increase its pressure to 119 psia to deliver it to Column 2. Note that Column 2 uses such a small amount of reflux, that it is essentially just a reboiled stripper. Sequence 1 is simpler and preferred over Sequence 2. However, the first column in Sequence 1 would benefit with more reflux and stages to reduce the loss of toluene. Exercise 10.33 Subject: Separation of propylene from propane by distillation using a rigorous equilibriumstage model Given: Feed of 360 lbmol/h of propylene and 240 lbmol/h of propane at 125.7oF and 294 psia. Distillate is to be 99 mol% propylene, and bottoms is to be 95 mol% propane. Column is equipped with a partial condenser and a partial reboiler. Pressure is 280 psia in the reflux drop. Condenser pressure drop is 2 psia, and tray pressure drop is 0.1 psi/tray. Tray efficiency is 100%. Column is to contain 200 trays. Therefore, bottoms pressure = 280 + 2 + 20 = 302 psia. Reflux ratio is approximately 15.9. Feed-stage location is approximately tray 140 from the top tray. Assumptions: SRK equation of state for K-values and enthalpies. Find: Reflux ratio for the optimal feed tray location. Analysis: The calculations were made with the SCDS Column model (Simultaneous-correction method) because of the small number of components and the large number of stages. It was understood that the Chemcad SCDS model would adjust the reflux rate and distillate rate to achieve the specified distillate and bottoms purities. Thus, the initial specifications for the SCDS Column model of the Chemcad program were: Partial condenser Top pressure = 280 psia Condenser pressure drop = 2 psia. Column pressure drop = 20 psia. Number of stages = 202 (includes partial condenser and partial reboiler) Feed stage =140 from the top as a first estimate of the optimal feed stage At the top, distillate purity = 99 mol% propylene At the bottom, bottoms purity = 95 mol% propane Estimated distillate rate = 350 lbmol/h Estimated reflux rate = 5582 lbmol/h Estimated temperature of stage 1 = 120oF Estimated temperature of bottom stage = 140oF Estimated temperature of stage 2 = 122oF Estimated temperature of next to bottom stage = 138oF When flashed, the feed was computed to be 69.11 mol% vaporized. Convergence of the calculations was achieved rapidly in 6 iterations, with a converged reflux ratio of 14.0047 compared to the initial estimate of 5582/350 = 15.95. The effect of feed-stage location on the reflux ratio was then studied with the results shown below. Analysis: (continued) Exercise 10.33 (continued) Feed stage from top 145 140 135 130 125 120 119 118 117 116 115 110 Reflux ratio 14.269 14.005 13.824 13.711 13.644 13.596 13.593 13.592 13.592 13.612 13.615 13.646 As shown above, the optimal feed stage is in the range of stages 117 to 119. For the feed stage at 117 (116 in the column), the product distribution is as follows, accompanied by: Condenser duty = 25,727,400 Btu/h Reboiler duty = 25,434,800 Btu/h The vapor and liquid compositions in the region of the feed stage, shown below, prove that a pinch condition does not exist, but the mole fractions are not changing rapidly. Stage from top 115 116 117 118 119 Mole fraction C3 in liquid 0.41794 0.41814 0.41931 0.42000 0.42073 Exercise 10.34 Subject: Use of a stabilizer to remove hydrogen and methane from an aromatic mixture. Given: A feed mixture of composition given below at 240oF and 275 psia. Column to contain about 10 equilibrium stages, with a partial condenser and a partial reboiler. Condenser outlet temperature = 100oF. Bottoms to contain no more than 0.05 mol% methane plus ethane. Tray efficiency = 70%. Column top pressure = 128 psia. Column bottom pressure = 132 psia. Suggested estimates of 2 for the reflux ratio and 49.4 lbmol/h Assumptions: SRK for K-values and enthalpies Find: A suitable design with number of trays, feed location, and reflux ratio. Analysis: A valve is used to drop the feed pressure to column pressure. The calculations were made with the Tower model (Inside-out method) of Chemcad. It was understood that the Chemcad TOWER model would adjust the reflux rate and distillate rate to achieve the specified distillate temperature and bottoms impurities. Thus, the initial specifications for the TOWER model of the Chemcad program were as follows, where it was assumed that the main impurity in the bottoms product would be ethane. Partial condenser Top pressure = 128 psia Condenser pressure drop = 0 psi. Column pressure drop = 4 psi. Number of stages = 12 (includes partial condenser and partial reboiler) Feed stage =3 from the top as a first estimate. At the top, distillate temperature = 100oF. At the bottom, bottoms purity = 0.0005 mole fraction of ethane Estimated distillate rate = 50 lbmol/h Estimated reflux rate = 100 lbmol/h Estimated temperature of stage 1 = 100oF (actual specification) Estimated temperature of bottom stage = 400oF Estimated temperature of stage 2 = 300oF A flash of the feed at feed conditions upstream of the feed valve showed it to be liquid. Convergence of the calculations was achieved rapidly in 6 iterations, with a converged reflux ratio, however of only 0.186, with a distillate rate of 48.715 lbmol/h. Thus, the reflux rate to the top stage was only 9.06 lbmol/h compared to 1748.4 lbmol/h of feed. This is considered to be unsatisfactory. Therefore, a second run was made with the feed entering the top stage (Stage 2) and reducing the number of stages to 7 (including the partial condenser and partial reboiler). A converged result shown on the next two pages was obtained in 5 iterations. This design is considered more satisfactory and achieves the specifications. The impurity in the bottoms is mainly ethane, as assumed. Other results are: reflux ratio = 0.0.211, condenser duty = 2,515,000 Btu/h, and reboiler duty = 12,976,000 Btu/h. The column will contain 5/0.7 = 8 trays. Analysis: (continued) Exercise 10.34 (continued) Exercise 10.34 (continued) Analysis: (continued) Exercise 10.35 Subject: Simulation of a four-stage, nearly isothermal distillation process by flash calculations. Given: Problem specifications in Figure 10.36, consisting of the feed conditions and temperature and pressure for each of four isothermal flashes. Assumptions: Grayson-Streed correlation for K-values and Redlich-Kwong equation for enthalpies. Find: Component flow rates for all equilibrium vapor and liquid streams. Analysis: The Chemcad simulator was used with its FLASH model in the isothermal flash mode. Referring to Figure 10.36, the recycle process was simulated with L1, V3, and V4 as the cut (tear) streams, with the sequence of calculations being Stages 2, 1, 3, and 4. No initial estimates were given for the cut streams. Convergence was achieved in 19 iterations using direct substitution. The resulting stream compositions are given below: Analysis (continued) Exercise 10.35 (continued) Exercise 10.36 Subject: Effect of number of stages, absorbent flow rate, and absorbent temperature on the absorption of a paraffin hydrocarbon gas. Given: 2,000 lbmol/h of a gas mixture containing in mole fractions, 0.830 C1, 0.084 C2, 0.048 C3, 0.026 nC4, and 0.012 nC5, at 75 psia and 60oF. Absorbent oil of MW = 161. Assumptions: SRK for K-values and enthalpies. Find: For the following four cases, compute product flow rates and compositions, and interstage vapor and liquid flowrates and compositions: Case (a) (b) (c) (d) No. of Stages 6 12 6 6 Absorbent rate, lbmol/h 500 500 1,000 500 Absorbent temperature, oF 90 90 90 60 Analysis: Chose nC11 , MW = 156, for the absorbent oil The calculations were made with the Tower model (Inside-out method) of Chemcad. For Case (a), the specs are: No condenser with top pressure = 75 psia Column pressure drop = 0 psi. Number of stages = 6 Feed gas to stage 6 Absorbent to stage 1 At the top, no condenser At the bottom, no reboiler Estimated distillate rate = 2,000 lbmol/h Estimated reflux rate = 500 lbmol/h Estimated temperature of stage 1 = 100oF Estimated temperature of bottom stage = 100oF Estimated temperature of stage 2 = 100oF Specifications for the other 3 cases differed only in the number of equilibrium stages and the flow and temperature of the absorbent oil, plus changes in the estimates. Convergence of each case was achieved in 3 or 4 iterations. The results for each case are given on the following s, where it is observed that 1. Doubling the number of stages, only increases the absorption of nC5. It is the only feed gas component with an absorption factor, A = L/KV, > 1. 2. Doubling the absorbent rate almost doubles the amount absorbed of each feed gas component. 3. Lowering the absorbent temperature increases somewhat the absorption of lighter components Analysis: (a) (continued) Exercise 10.36 (continued) Exercise 10.36 (continued) Analysis: (a) (continued) Exercise 10.36 (continued) Analysis: (b) (continued) Exercise 10.36 (continued) Analysis: (b) (continued) Exercise 10.36 (continued) Analysis: (c) (continued) Exercise 10.36 (continued) Analysis: (c) (continued) Exercise 10.36 (continued) Analysis: (d) (continued) Exercise 10.37 Subject: Absorption of a hydrocarbon gas with n-decane. Given: 1000 lbmol/h of a paraffin hydrocarbon gas, of composition below, at 90oF and 75 psia. 1000 lbmol/h of nC10 absorbent at the same temperature and pressure. Absorber has 4 equilibrium stages and operates at 75 psia. Assumptions: SRK for K-values and enthalpies. Find: Product rates and compositions, stage temperatures, and interstage vapor and liquid flow rates and compositions. Analysis: As shown below, when the gas feed was flashed, it was found to contain 7.2 mol% liquid. The calculations were made with the Tower model (Inside-out method) of Chemcad, with the following specifications: No condenser Column pressure drop = 0 psi. Number of stages = 4 Feed gas to stage 4 Absorbent to stage 1 At the top, no condenser At the bottom, no reboiler Estimated distillate rate = 1,000 lbmol/h Estimated reflux rate = 1,000 lbmol/h Estimated temperature of stage 1 = 90oF Estimated temperature of bottom stage = 90oF Estimated temperature of stage 2 = 90oF Convergence was achieved only with great difficulty after 29 iterations. The problem was also solved with the SCDS Column model (simultaneous correction method) in only 8 iterations. The percent absorption of each component is as follows: Component Methane Ethane Propane n-Butane n-Pentane % Absorbed 4.8 21.0 60.6 97.3 99.95 The other results are presented on the next two pages. Exercise 10.37 (continued) Analysis: (continued) Exercise 10.37 (continued) Analysis: (continued) Exercise 10.38 Subject: Effect of intercooler location on the absorption of a light hydrocarbon mixture. Given: Feed gas and absorbent oil of Figure 10.17 of Example 10.4. Incooler with a duty of 500,000 Btu/h. Assumptions: SRK for K-values and enthalpies. Absorbent oil is nC12. Find: Effectiveness of the intercooler and its preferred location. Possible change in duty if intercooler stage temperature is at least 100oF. Analysis: The calculations of Example 10.4 were first repeated using the TOWER model of Chemcad. The following results differ from those of Example 10.4 because of the different methods used for computing K-values and enthalpies: Component C1 C2 C3 nC4 nC5 nC12 Total: Temp., oF lbmol/h: Feed gas 160.00 370.00 240.00 25.00 5.00 0.00 800.00 105 Absorbent 0.00 0.00 0.00 0.05 0.78 164.17 165.00 90 Gas out 146.25 268.83 96.41 1.11 0.21 0.15 512.96 144.9 Liquid out 13.75 101.17 143.59 23,94 5.57 164.02 452.04 144.4 Compared to results of Example 10.4, more is absorbed, but the temperature rise is less. For the effect of the intercooler, the calculations were also made with the Tower model (Insideout method) of Chemcad. The specifications for an intercooler of -500,000 Btu/h located at Stage 2 from the top were as follows, where it should be noted that the intercooler is treated as a feed between the absorbent and feed gas. No condenser Top pressure = 400 psia Column pressure drop = 0 psi. Number of stages = 6 Absorbent to stage 1 Intercooler with -0.500 MMBtu/h enthalpy (as a feed) at stage 2 Feed gas to stage 6 At the top, no condenser At the bottom, no reboiler Estimated distillate rate = 500 lbmol/h Estimated reflux rate = 200 lbmol/h Exercise 10.38 (continued) Analysis: (continued) Estimated temperature of stage 1 = 120oF Estimated temperature of bottom stage = 130oF Estimated temperature of stage 2 = 120oF Convergence was achieved in 6 iterations. The calculations were repeated for the intercooler located at stages 3, 4, and 5. The following results show that the outlet gas temperature was reduced and the amount of absorption was considerably increased, but the effect of stage location is not significant, although a location at stage 4 or 5 appeared to be very slightly best. Case Base - no intercooler (a) - intercooler stage 2 (b) - intercooler stage 3 (c) - intercooler stage 4 (d) - intercooler stage 5 Liquid out, lbmol/h 452.04 498.69 500.53 500.74 500.74 Temp., gas out, oF 144.9 133.8 137.3 138.6 138.6 Because the temperature was still significantly higher than the lower limit of 100oF, a higher intercooler heat duty of -2,000,000 Btu/h was tried at stage 5. However, the temperature of that stage dropped to 84oF. For a duty of -1,500,000 Btu/h, the temperature of stage 5 was an acceptable value of 102oF. The liquid out increased to 599.68 lbmol/h and the temperature of the gas out decreased to 127.5oF. The corresponding component absorption rates compared to those of the base case (no intercooler) were as follows: lbmol/h: Component No intercooler -1,500,000 Btu/h, stage 5 C1 13.75 21.09 C2 101.17 171.52 C3 143.59 212.32 nC4 23.94 24.98 nC5 5.57 5.66 Total: 288.02 435.57 Thus, the effect of the intercooler is very significant. The corresponding stage temperatures compared to those of the base case are also of interest: o F: 2 153.1 134.1 3 156.1 135.1 4 157.3 128.5 5 154.6 102.1 6 144.3 114.9 Stage No intercooler -1,500,000 Btu/h - stage 5 1 144.9 127.5 Exercise 10.39 Subject: Design of a an absorber with two absorbent oil feeds. Given: Feed gas at 90oF and 400 psia with the composition given below. Column has 8 equilibrium stages. Lean oil, assumed to be nC12, at 80oF and 400 psia at 250 lbmol/h enters the top stage. Secondary oil at 80oF and 400 psia, of the composition given below, enters stage 7 from the top. 150,000 Btu/h (-0.15 MMBtu/h) is removed by intercooling from stage 7. The feed gas enters stage 8. Assumptions: SRK for K-values and enthalpies. Find: Product rates and compositions, stage temperatures, and interstage vapor and liquid flow rates and compositions. Analysis: The calculations were made with the Tower model (Inside-out method) of Chemcad, with the following specifications: No condenser Top pressure = 400 psia Column pressure drop = 0 psi. Number of stages = 8 Lean oil (1st feed) to stage 1 Secondary oil (2nd feed) to stage 4 Intercooler (3rd feed) at stage 7 Feed gas (4th feed) to stage 8 At the top, no condenser At the bottom, no reboiler Estimated distillate rate = 450 lbmol/h Estimated reflux rate = 350 lbmol/h Estimated temperature of stage 1 = 110oF Estimated temperature of bottom stage = 100oF Estimated temperature of stage 2 = 100oF Note that the intercooler is treated as the 3rd feed with a specification as a stream of -0.15 MM Btu/h. Convergence was achieved in 4 iterations. The results are shown on the following two pages. Note that the degree of absorption was good with following results: Component C1 C2 C3 nC4 nC5 % in Feed gas absorbed 15.2 73.4 99.2 99.98 100.00 Exercise 10.39 (continued) Analysis: (continued) Exercise 10.39 (continued) Analysis: (continued) Exercise 10.40 Subject: Reboiled absorber for the separation of a hydrocarbon mixture. Given: Absorbent oil of nC9 at 60oF and 230 psia and 40 lbmol/h. Feed at 120oF and 230 psia with the composition given below. Column has 15 equilibrium stages plus a partial reboiler. Absorbent oil enters stage 1 (top stage). Feed enters stage 9. Interreboiler with a heat duty of 1.0 MMBtu/h at stage13. Column operates at 230 psia. Overhead flow rate = 103 lbmol/h. Assumptions: SRK for K-values and enthalpies. Find: Product rates and compositions, stage temperatures, and interstage vapor and liquid flow rates and compositions. Reboiler duty with and without the interreboiler. Whether an interreboiler is worthwhile and if an intercooler should be considered. Analysis: The calculations were made with the Tower model (Inside-out method) of Chemcad, with the following specifications: No condenser Top pressure = 230 psia Column pressure drop = 0 psi. Number of stages = 15 plus a partial reboiler Absorbent oil (1st feed) to stage 1 Feed (2nd feed) to stage 9 Interreboiler (3rd feed) to stage 13 At the top, no condenser At the bottom, bottoms flow rate = 459 - 103 = 356 lbmol/h Estimated distillate rate = 103 lbmol/h (specified rate) Estimated reflux rate = 40 lbmol/h Estimated temperature of stage 1 = 80oF Estimated temperature of bottom stage = 300oF Estimated temperature of stage 2 = 90oF Note that the interreboiler is treated as the 3rd feed with a specification as a stream of 1.00 MMBtu/h. Convergence was achieved in 7 iterations. The results are shown on the following two pages. The column feed is 14.2% vaporized. The distillate rate is set to send the methane and ethane overhead, while a portion of the propane and all heavier components to the bottoms. The temperature of the interreboiler stage is 202.6oF compared to a bottoms temperature of 302.4oF. The heat duty of the reboiler is 3.164 MMBtu/h. Thus, almost 25% of the total required heat duty can be supplied by waste heat.. Results without the interreboiler, on a subsequent page, show only a small effect on the separation, with the major effect being 17% more ethane leaving in the overhead vapor and a reboiler duty of 4.177 MMBtu/h. If an intercooler of a heat duty of 100,000 MMBtu/h is added at stage 4, along with the interreboiler at stage 13, as shown below, less ethane leaves in the overhead vapor. Thus, the use of an intercooler is not desirable. Exercise 10.40 (continued) Analysis: (continued) With interreboiler Exercise 10.40 (continued) Analysis: (continued) With interreboiler Exercise 10.40 (continued) Analysis: (continued) Without interreboiler Exercise 10.40 (continued) Analysis: (continued) With interreboiler and intercooler Exercise 10.41 Subject: Design of a reboiled stripper. Given: Feed of composition below at 39.2oF and 150 psia. Column with 7 equilibrium stages plus a partial reboiler, operating at 150 psia. Bottoms flow rate of 99.33 lbmol/h. Assumptions: SRK for K-values and enthalpies. Find: Product compositions, stage temperatures, interstage flow rates and compositions, and reboiler duty. Analysis: : The calculations were made with the Tower model (Inside-out method) of Chemcad, with the following specifications: No condenser Top pressure = 230 psia Column pressure drop = 0 psi. Number of stages = 7 plus a partial reboiler Feed to stage 1 At the top, no condenser At the bottom, bottoms flow rate = 99.33 lbmol/h Estimated distillate rate = 450 lbmol/h (actual rate = 551.59 - 99.33 = 452.26 lbmol/h) Estimated reflux rate = 550 lbmol/h Estimated temperature of stage 1 = 50oF Estimated temperature of bottom stage = 250oF Estimated temperature of stage 2 = 80oF Convergence was achieved in 4 iterations. The results are shown on the following two pages. The column feed is 14.5 mol% vaporized. The heat duty of the reboiler is 4.4674 MMBtu/h. The % stripping of each component in the feed is as follows: Component N2 C1 C2 C3 nC4 nC5 nC6 % Stripped 100.00 100.00 100.00 99.92 86.26 23.61 8.03 Exercise 10.41 (continued) Analysis: (continued) Analysis: (continued) Exercise 10.41 (continued) Exercise 10.42 Subject: Liquid-liquid extraction of a mixture of cyclohexane (1) and cyclopentane (2) with methanol (3). Given: Feed of 700 lbmol/h of cyclohexane and 300 lbmol/h of cyclopentane. Methanol solvent of 1,000 lbmol/h. Extractors with various numbers of equilibrium stages, operating at 25oC. Assumptions: Phase equilibria predicted by NRTL or UNIQUAC equations. Find: Product rates and compositions, and interstage flow rates and compositions for 1, 2, 5, and 10 equilibrium stages. Analysis: Reference 55 in Chap. 2 presents L/L equilibria data for the ternary system at 25oC (77oF). These data are fitted to the NRTL and UNIQUAC equations in "Liquid-Liquid Equilibrium Data Collection - Ternary Systems", Vol. V, Part 2 of the Chemical Data Series, by J. M. Sorenson and W. Arlt, DECHEMA (1980). Their fit to the simpler NRTL equation is almost as good as the UNIQUAC fit. Therefore, the NRTL equation is used here, with the following binary interaction parameters for use in the Chemcad program, where from Eq. (2-94), ij = gij - g jj RT = Bij T , where T is in Kelvins I 1 1 2 J 2 3 3 BIJ -97.419 544.82 361.38 BJI -56.868 318.02 234.33 IJ 0.2 0.2 0.2 Using the Binodal Plot feature of the Chemcad program, a predicted ternary liquid-liquid equilibrium diagram is obtained, as shown on the next page, along with a corresponding table of predicted tie lines. It is seen that at 25oC, cyclopentane is completely soluble in methanol, but cyclohexane is only partially soluble in methanol. Is methanol a good solvent for the separation of cyclopentane from cyclohexane? From the slope of the tie lines, the answer is no; Cyclopentane distributes preferably to the cyclohexane-rich phase. For example, for one equilibrium stage, using the LLVF model of Chemcad, the following two liquid phases (raffinate and extract) are calculated when the given feed and solvent are mixed and brought to equilibrium at 25oC and 30 psia, where the pressure exceeds the bubble-point pressures of the feed (2.69 psia) and solvent (2.44 psia) so that a vapor phase does not form. Very little separation of (1) and (2) occurs. Exercise 10.42 (continued) Analysis: (continued) Analysis: (continued) Exercise 10.42 (continued) The Extract model of the Chemcad program was used to calculate a 10-stage extraction unit for the given conditions. The model converged in 12 iterations, with the results given below. The product compositions differ little from those computed for 1 stage. Also, the interstage flow rates are almost constant for stages 1 to 7. Thus, a significant separation is not possible for any number of stages. The reason for this is clear if a line on the above ternary diagram is drawn from the feed point [30 mol% (2) at 0 mol% (3)] to the solvent point [100 mol% (3)]. This line is almost collinear with a tie line. Regardless of the number of stages and the number of equilibrium stages, a significant separation of cyclopentane and cyclohexane can not be made. Exercise 10.43 Subject: Liquid-liquid extraction of acetic acid from water by ethyl acetate at 100oF Given: Feed of 6,660 lb/h of acetic acid and 23,600 lb/h of water. Solvent of 68,600 lb/h of ethyl acetate and 2,500 lb/h of water. Ccountercurrent-flow extractor. Assumptions: NRTL equation for activity coefficients using the following parameters: I Ethyl acetate Ethyl acetate Water Find: J Water Acetic acid Acetic acid BIJ 166.36 643.30 -302.63 BJI 1190.1 -702.57 -1.683 IJ 0.2 0.2 0.2 Number of equilibrium stages to extract 99.5% of the acetic acid. Analysis: The Extract model of Chemcad was used starting with 2 equilibrium stages and increasing by one until 99.5% extraction was achieved. For each run, convergence was achieved in 12 to 13 iterations. The results were: No. of stages % Extraction of Acetic acid 2 90.59 3 96.68 4 98.80 5 99.56 6 99.86 The results show that 5 stages are needed, but 6 stages is safer. Detailed results for 6 stages are: Subject: Azeotropes and approximate residue curve map for the system n-hexane methanol - methyl acetate at 1 atm Given:. Ternary system n-hexane - methanol - methyl acetate at 1 atm. Find: All binary and ternary azeotropes from suitable references. Approximate residue curve map showing distillation boundaries. Type of node for each azeotrope and pure component. Analysis: The following azeotrope data were obtained from the "Handbook of Chemistry and Physics" or the 6th edition of "Perry's Chemical Engineers' Handbook": At 1 atm, the molecular weights and boiling points of the three components are: Component Molecular weight Normal b. pt., oC n-Hexane 86.17 69.0 Methanol 32.04 64.7 Methyl acetate 74.08 57.0 Binary azeotropes (all minimum boiling): Mixture, A/B Boiling pt., oC n-Hexane/Methanol 50.6 n-Hexane/Methyl acetate 51.8 Methanol/Methyl acetate 53.5 Exercise 11.1 Wt% A/B 39.3/60.7 19.0/81.0 Mol% A/B 49.0/51.0 35.8/64.2 35.2/64.8 Ternary azeotrope (minimum boiling): Mixture, A/B/C Boiling pt., oC Wt% A/B/C n-Hexane/Methanol/ Methyl acetate 47.4 48.6/14.6/36.8 Mol% A/B/C 37.2/30.1/32.7 An approximate residue curve map on a right triangular diagram, of the type developed in Fig. 11.7, is shown on the next page. From that diagram, which includes approximate distillation boundaries, the following types of nodes are determined using Fig. 11.6: Component or Azeotrope Normal boiling pt., oC Type node n-Hexane 69.0 stable Methanol 64.7 stable Methyl acetate 57.0 stable 50.6 saddle n-Hexane/Methanol n-Hexane/Methyl acetate 51.8 saddle Methanol/Methyl acetate 53.5 saddle n-Hexane/Methanol/Methyl acetate 47.4 unstable A computed residue curve map is given by van Dongen and Doherty, IEC Fundamentals, Vol. 24, 454 (1985). Analysis: (continued) Exercise 11.1 (continued) Subject: acetate. Exercise 11.2 Calculation of a residue curve for the system n-hexane - methanol - methyl Given:. Ternary system n-hexane - methanol - methyl acetate at 1 atm. Assumptions: UNIFAC method for computing K-values. Find: Portion of a residue curve starting from a bubble-point liquid of 20 mol% n-Hexane (1), 60 mol% methanol (2), and 20 mol% methyl acetate (3), at 1 atm. Analysis: The residue curve is computed numerically, in the same manner as in Example 11.1, except that the bubble-point calculations at each step are made with the Chemcad FLASH model. To begin, a bubble point is run with Chemcad on the above initial liquid composition, xi(0), at 1 atm. The computed vapor composition, yi(0), is then used with the Euler form of Eqs. (11-5) and (11-6) to compute xi(1): xi(1) = xi( 0) + xi( 0) - yi( 0) , i = 1 to C -1 The procedure is repeated in steps of = 0.1 in the forward direction from = 0 to = 1, and = -0.1in the backward direction from = 0 to = -1. The results are as follows, with a plot on the following page. -1.0 -0.9 -0.8 -0.7 -0.6 -0.5 -0.4 -0.3 -0.2 -0.1 0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.0 x (1) 0.3277 0.3203 0.3120 0.3027 0.2923 0.2806 0.2676 0.2531 0.2371 0.2194 0.2000 0.1806 0.1597 0.1373 0.1137 0.0894 0.0652 0.0426 0.0233 0.0097 0.0028 x (2) 0.4560 0.4647 0.4744 0.4852 0.4970 0.5102 0.5248 0.5409 0.5587 0.5784 0.6000 0.6216 0.6451 0.6704 0.6974 0.7260 0.7556 0.7853 0.8137 0.8389 0.8598 x (3) 0.2163 0.2150 0.2136 0.2121 0.2107 0.2092 0.2076 0.2060 0.2042 0.2022 0.2000 0.1978 0.1952 0.1923 0.1889 0.1846 0.1792 0.1721 0.1630 0.1514 0.1374 y (1) 0.3933 0.3942 0.3950 0.3959 0.3966 0.3973 0.3978 0.3979 0.3975 0.3964 0.3940 0.3900 0.3835 0.3730 0.3566 0.3311 0.2920 0.2350 0.1591 0.0797 0.0260 y (2) 0.3770 0.3773 0.3776 0.3779 0.3782 0.3786 0.3790 0.3796 0.3804 0.3818 0.3840 0.3871 0.3920 0.3995 0.4114 0.4300 0.4588 0.5019 0.5617 0.6300 0.6866 y (3) 0.2297 0.2285 0.2274 0.2262 0.2252 0.2241 0.2232 0.2225 0.2221 0.2218 0.2220 0.2229 0.2245 0.2275 0.2320 0.2389 0.2492 0.2631 0.2792 0.2903 0.2874 T, oC 47.88 47.88 47.88 47.88 47.88 47.88 47.88 47.89 47.90 47.93 47.99 48.08 48.24 48.51 48.96 49.68 50.80 52.44 54.59 56.83 58.46 Analysis: (continued) Exercise 11.2 (continued) Exercise 11.3 Subject: : Calculation of a portion of a distillation curve for the system n-hexane - methanol methyl acetate. Given:. Ternary system n-hexane - methanol - methyl acetate at 1 atm. Assumptions: UNIFAC method for computing K-values. Find: Portion of a distillation curve starting from a bubble-point liquid of 20 mol% n-Hexane (1), 60 mol% methanol (2), and 20 mol% methyl acetate (3), at 1 atm, using a simulation program. Analysis: The distillation curve is computed stagewise using the FLASH model in Chemcad, in the same manner as in Example 11.2, except that here the calculation is made in both directions. To begin, a bubble point is run with Chemcad on the above initial liquid composition, xi(1), at 1 atm. The computed vapor composition, yi(1), is set equal to xi(2), according to Eq. (11-14) so that the compositions of passing vapor and liquid streams are made equal, and the bubble-point calculation is repeated. In this manner, we step in the first direction. The sequence is repeated, starting again from xi(1), but this time running a sequence of dew points to step in the opposite direction. The results are as follows, with a plot on the following page. x (1) 0.3753 0.3783 0.3822 0.3873 0.3939 0.2000 0.0272 0.0022 0.0002 0.0000 0.0000 0.0000 x (2) 0.3594 0.3635 0.3689 0.3757 0.3840 0.6000 0.8662 0.9586 0.9869 0.9958 0.9987 0.9996 x (3) 0.2653 0.2582 0.2489 0.2370 0.2221 0.2000 0.1066 0.0392 0.0129 0.0042 0.0013 0.0004 y (1) 0.3731 0.3753 0.3783 0.3822 0.3873 0.3939 0.2000 0.0272 0.0022 0.0002 0.0000 0.0000 y (2) 0.3562 0.3594 0.3635 0.3689 0.3757 0.3840 0.6000 0.8662 0.9586 0.9869 0.9958 0.9987 y (3) 0.2707 0.2653 0.2582 0.2489 0.2370 0.2221 0.2000 0.1066 0.0392 0.0129 0.0042 0.0013 T, oC 47.84 47.84 47.85 47.86 47.89 47.99 55.01 61.89 63.74 64.25 64.41 64.46 In the plot on the next page, the distillation curve is seen to take a sharp turn at a n-hexane mole fraction of approximately 0.40. This may be due to the drastic changes that occur to the K-values along the distillation curve for this very non-ideal system. Starting from almost pure methyl alcohol and moving along the distillation curve to right, the K-values change as follows: Analysis: (continued) Exercise 11.3 (continued) K-values: 14.29 1.00 3.16 Component nC6 MeOH MeAc 12.20 0.90 2.73 7.37 0.69 1.88 1.97 0.64 1.11 0.98 0.98 1.07 0.99 0.99 1.01 Note how the final K- values are tending toward the azeotropic condition. This may explain the sharp turn of the distillation curve on the plot below. Exercise 11.4 Subject: Azeotropes and approximate residue curve map for the system n-hexane - methanol methyl acetate at 1 atm Given:. Ternary system acetone - benzene - n-heptane at 1 atm. Find: All binary and ternary azeotropes from suitable references. Approximate distillation curve map showing distillation boundaries. Type of node for each azeotrope and pure component. Analysis: The following azeotrope data were obtained from the "Handbook of Chemistry and Physics". At 1 atm, the molecular weights and boiling points of the three components are: Component Molecular weight Normal b. pt., oC n-Heptane 100.2 98.4 Benzene 78.11 80.1 Acetone 58.08 56.5 Binary azeotropes (all minimum boiling): Mixture, A/B Boiling pt., oC n-Heptane/Benzene 80.1 n-Heptane/Acetone 55.9 Benzene/Acetone No azeotrope No ternary azeotrope An approximate distillation curve map on a equilateral triangular diagram, of the type in Fig. 11.3, is shown on the next page. However, it is not a sketch, but is a residue curve map drawn by the Aspen Plus program and modified to represent a distillation curve map. The distillation boundary extends from one binary azeotrope to the other, but is barely discernible. From that diagram, the following types of nodes are determined using Fig. 11.6: Component or Azeotrope n-Heptane Benzene Acetone n-Heptane/Benzene n-Heptane/Acetone Normal boiling pt., oC 98.4 80.1 56.5 80.1 55.9 Type node stable stable saddle saddle unstable Wt% A/B 0.7/99.3 10.5/89.5 Mol% A/B 0.5/99.5 6.5/93.5 Analysis: (continued) Exercise 11.4 (continued) Exercise 11.5 Subject: Calculation of a residue curve for the system acetone - benzene - n-heptane. Given:. Ternary system acetone - benzene - n-heptane at 1 atm. Assumptions: UNIFAC method for computing K-values. Find: Portion of a residue curve starting from a bubble-point liquid of 20 mol% acetone (1), 60 mol% benzene (2), and 20 mol% n-heptane (3), at 1 atm. Analysis: The residue curve is computed numerically, in the same manner as in Example 11.1, except that the bubble-point calculations at each step are made with the Chemcad FLASH model. To begin, a bubble point is run with Chemcad on the above initial liquid composition, xi(0), at 1 atm. The computed vapor composition, yi(0), is then used with the Euler form of Eqs. (11-5) and (11-6) to compute xi(1): xi(1) = xi( 0) + xi( 0) - yi( 0) , i = 1 to C -1 The procedure is repeated in steps of = 0.1 in the forward direction from = 0 to = 1, and = -0.1in the backward direction from = 0 to = -1. The results are as follows, with a plot on the following page. -1.0 -0.9 -0.8 -0.7 -0.6 -0.5 -0.4 -0.3 -0.2 -0.1 0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.0 x (1) 0.4458 0.4224 0.3985 0.3740 0.3491 0.3239 0.2986 0.2733 0.2483 0.2238 0.2000 0.1762 0.1534 0.1318 0.1118 0.0935 0.0770 0.0625 0.0500 0.0395 0.0307 x (2) 0.4248 0.4425 0.4604 0.4786 0.4969 0.5151 0.5332 0.5509 0.5681 0.5845 0.6000 0.6155 0.6298 0.6426 0.6539 0.6633 0.6708 0.6763 0.6797 0.6811 0.6807 x (3) 0.1294 0.1351 0.1411 0.1474 0.1540 0.1610 0.1682 0.1758 0.1836 0.1917 0.2000 0.2083 0.2169 0.2255 0.2343 0.2432 0.2522 0.2612 0.2703 0.2794 0.2886 y (1) 0.6723 0.6558 0.6381 0.6189 0.5981 0.5758 0.5518 0.5259 0.4983 0.4690 0.4381 0.4043 0.3689 0.3322 0.2951 0.2581 0.2218 0.1874 0.1556 0.1272 0.1019 y (2) 0.2519 0.2658 0.2807 0.2968 0.3141 0.3326 0.3525 0.3738 0.3964 0.4202 0.4452 0.4726 0.5010 0.5303 0.5596 0.5884 0.6162 0.6420 0.6651 0.6851 0.7020 y (3) 0.0758 0.0784 0.0812 0.0843 0.0878 0.0916 0.0957 0.1003 0.1053 0.1108 0.1167 0.1231 0.1301 0.1375 0.1453 0.1535 0.1620 0.1706 0.1793 0.1877 0.1961 T, oC 63.02 63.53 64.07 64.67 65.30 65.99 66.73 67.52 68.35 69.23 70.14 71.13 72.14 73.18 74.21 75.23 76.21 77.13 77.98 78.74 79.42 Analysis: (continued) Exercise 11.5 (continued) Exercise 11.6 Subject: : Calculation of a portion of a distillation curve for the system acetone - benzene n-heptane. Given:. Ternary system acetone - benzene - n-heptane at 1 atm. Assumptions: UNIFAC method for computing K-values. Find: Portion of a distillation curve starting from a bubble-point liquid of 20 mol% Acetone (1), 60 mol% Benzene (2), and 20 mol% n-Heptane (3), at 1 atm, using a simulation program. Analysis: The distillation curve is computed stagewise using the FLASH model in Chemcad, in the same manner as in Example 11.2, except that here the calculation is made in both directions. To begin, a bubble point is run with Chemcad on the above initial liquid composition, xi(1), at 1 atm. The computed vapor composition, yi(1), is set equal to xi(2), according to Eq. (11-14) so that the compositions of passing vapor and liquid streams are made equal, and the bubble-point calculation is repeated. In this manner, we step in the first direction. The sequence is repeated, starting again from xi(1), but this time running a sequence of dew points to step in the opposite direction. The results are as follows, with a plot on the following page. x (1) 0.0000 0.0000 0.0003 0.0024 0.0146 0.0633 0.2000 0.4384 0.6643 0.7996 0.8702 0.9076 x (2) 0.0445 0.0967 0.1953 0.3464 0.5125 0.6196 0.6000 0.4447 0.2695 0.1510 0.0879 0.0526 x (3) 0.9555 0.9033 0.8044 0.6512 0.4729 0.3171 0.2000 0.1169 0.0698 0.0494 0.0419 0.0397 y (1) 0.0000 0.0003 0.0024 0.0146 0.0633 0.2000 0.4384 0.6643 0.7996 0.8702 0.9076 0.0000 y (2) 0.0967 0.1953 0.3464 0.5125 0.6196 0.6000 0.4447 0.2695 0.1510 0.0879 0.0526 0.9996 y (3) 0.9033 0.8044 0.6512 0.4729 0.3171 0.2000 0.1169 0.0698 0.0494 0.0419 0.0397 0.0004 T, oC 96.65 94.59 91.23 86.98 82.59 77.39 70.19 70.20 63.31 59.44 57.73 56.94 Analysis: (continued) Exercise 11.6 (continued) Exercise 11.7 Subject: Regions of feasible product composition on a ternary diagram Given: Residue curve map of Fig. 11.13, which shows a highly curved distillation boundary Find: Feasible product composition regions for Feed F1. Analysis: The solution to this exercise is given in detail by Widagdo and Seider in Ref. 19 of Chapter 11. Their figure 20, shown below, shows the bow-tie regions. Because the distillation boundary is highly curved, as discussed in Section 11.1, feasible product compositions lie on both sides of the distillation boundary when the feed is at F1. The two dark-shaded feasible regions are on the feed side of the distillation boundary, while the two lightly shaded regions are on the other side of the distillation boundary. As stated by Widagdo and Seider, "The region of feasible compositions on the feed side is bounded by the distillation line p that passes through the feed, the distillation boundary (curve q), and the two overall mass-balance lines through pure species H and L. The bottoms products lie in the shaded portion of the F1RKHF1 region, and the distillates lie in the shaded portion of the F1PLF1 region. On the other side of the distillation boundary, the region of feasible bottoms-product compositions is bounded by the distillation boundary (curve q) and the overall mass-balance line that passes through pure species L (line LF1M); that is, the lightly shaded region RMKR. The region of distillate compositions is also bounded by the distillation boundary (curve q). In addition, it is bounded by the end points of the chords (BiDi) that are tangent to the distillation lines at the tangent points (K, B1,..., B4 ,..., I) forming the dashed curve in the figure below. Because the distillate and bottoms-product compositions must lie on the same distillation line, for the feed composition at F1, the distillate compositions must lie to the left of the overall mass-balance lines through the tangent points at Bi (Di F1 Bi). It follows that, for the feed at F1, the regions of feasible distillate and bottoms-product compositions, are the shaded portions of the union of F1PLF1 and LD0 ... D3L regions, and the F1MHF1 regions, respectively." Analysis: (continued) Exercise 11.7 (continued) Exercise 11.8 Subject: Regions of feasible product composition on a ternary diagram Given: Residue curve map and distillation curve map of Fig. 11.10 for the ternary system, acetone, methanol, and chloroform, which has one maximum-boiling binary azeotrope, two minimum-boiling azeotropes and one ternary azeotrope, all homogeneous. Find: Feasible product composition regions for a feed of 50 mol% chloroform, 25 mol% acetone, and 25 mol% methanol. Analysis: In the right-triangle diagram on the next page, the feed composition is shown at F. Thus, the feed is in Region 1, which is bounded by distillation boundaries A and B, the binary mixture of chloroform and acetone from pure methanol to the azeotrope, and the binary mixture of chloroform and methanol from pure chloroform to the binary azeotrope. For Region 1, the minimum boiling product is AZ1, the azeotrope of chloroform and methanol, boiling at 53.9oC, while the maximum boiling product is AZ2, the azeotrope of chloroform and acetone, boiling at 65.5oC. To determine the feasible product compositions, one straight line is drawn from AZ1 to its intersection with distillation boundary B. Another straight line is drawn from AZ2 to its intersection with distillation boundary A. These two lines define an approximate bow-tie region through which a line from the feed point F to either pure chloroform and the ternary azeotrope (both infeasible products) does not pass. This approximate bow tie region is refined by adding a residue curve (or distillation curve) that passes through the feed point F. The final feasible product composition region is confined to the convex side of this added residue curve, as shown by the shaded bow-tie region. Typical bow-tie regions are shown for this system in the equilateral triangle diagram of Figure 11.12b. For the feed of this exercise, the feasible product composition regions are shown on an equilateral triangle residue curve map diagram, drawn by Aspen Plus, on the last page of this exercise. Analysis: (continued) Exercise 11.8 (continued) Exercise 11.8 (continued) Analysis: (continued) Exercise 11.9 Subject: Separation of acetone from methanol by extractive distillation with ethanol. Given: Bubble-point feed of 30 mol/s acetone (A) and 10 mol/s methanol (M) at 1 atm. Ethanol (E) as the solvent. Results of Example 11.3 with a solvent of water. Assumptions: Because the feed is close to the azeotropic composition of 22 mol% methanol, use a two-column system, with the first column being extractive distillation and the second ordinary distillation to recover the solvent. Column operations at 1 atm. UNIFAC for K-values. Negligible tray pressure drop. 100% tray efficiency. Find: Suitable column designs to obtain an acetone product of at least 95 mol%, a methanol product of at least 98 mol%, and ethanol of 99.9 mol% purity for recycle. Analysis: Extractive Distillation Column: First compute relative volatilities over a composition range to determine: (1) which of the two feed components is the most volatile in the presence of ethanol and (2) potential ease or difficulty of the separations. This is done with the LLVF threephase flash model of Chemcad, which also detects any formation of a second liquid phase, by conducting a series of flashes at 1 atm and a V/F = 0.0001 (a near bubble-point condition), for different mole fractions of the two feed components. For all cases, the mole fraction of the solvent, ethanol, is fixed at 0.5 so as to have a strong effect on the liquid-phase activity coefficients of acetone and methanol. The flash model computes the K-values, from which the relative volatilities can be obtained. Since acetone and methanol boil at 56.5oC and 64.7oC, respectively, and methanol is closer in structure to ethanol than is acetone, it is expected that the volatility of acetone to methanol will be increased in the presence of ethanol The following results confirm this expectation: xA 0.475 0.375 0.250 0.125 0.025 KA 1.444 1.579 1.832 2.234 2.746 KM 1.043 1.019 1.032 1.105 1.233 KE 0.576 0.561 0.568 0.613 0.692 A-M 1.38 1.55 1.78 2.02 2.23 A-E 2.51 2.81 3.23 3.64 3.97 Although these relative volatilities are favorable, they are not as high as those obtained when water is used as the solvent. Nevertheless, an initial run was made using the same arrangement of stages as in Example 11.3, a solvent rate equal to 60 mol/s, entering as a liquid at 60oC, and a reflux ratio of 4. The following results were obtained using the SCDS model of Chemcad, with an assumed bottoms rate of 68.774 mol/s, as computed in Example 11.3: Exercise 11.9 (continued) Analysis (continued) Component Acetone, A Methanol, M Ethanol, E Mole fractions: Distillate Bottoms 0.8801 0.0366 0.0379 0.1282 0.0820 0.8352 The desired separation is not achieved. The distillate contains too much methanol and ethanol, while the bottoms contains too much acetone. The acetone recovery is only 70.3%. In Example 11.3, it is 99.5%. Therefore, the stage, reflux, and solvent rate were increased in increments until the specification for acetone purity, and a high acetone recovery were met. The following Chemcad SCDS results, for the listed conditions, satisfy the product specifications, but other combinations of stages, reflux, and solvent rate could also give satisfactory results: Number of stages = 57 (includes total condenser and partial reboiler) Solvent enters stage 12 from the top Feed enters stage 36 from the top Reflux ratio = 5 Bottoms flow rate = 78.774 mol/s Solvent flow rate = 70 mol/s Initial guesses: Distillate rate = 108 kmol/h Reflux rate = 540 kmol/h o Temperatures: Stage 1, 55 C Stage 57, 75oC Stage 2, 60oC Stage 56, 70oC Convergence was achieved in 34 iterations with the following material balance: o Temperature, C Flow rates, mol/s: Acetone Methanol Ethanol Total: Solvent 60 0 0 70 70 Feed 55.2 30 10 0 40 Distillate 56.3 29.87 0.11 1.24 31.22 Bottoms 75.4 0.13 9.89 68.76 78.78 The liquid phase composition profile is shown on the next page. The extractive distillation column was sized with Chemcad for sieve trays, with: 24-inch tray spacing 12% downcomer area 10% hole area 1/4-inch holes 80% of flooding The resulting column inside diameter was determined to be 7 feet, which gives 79.5% of flooding at the top tray and 72.2% of flooding at the bottom. The tray pressure drop averages 0.094 psi/tray, for a total of 5.2 psi or 0.35 atm for the 55 trays. Thus, the assumption of negligible tray pressure drop is not satisfactory. For a final design, the computations should be repeated taking into account tray pressure drop and tray efficiency. Analysis (continued) Exercise 11.9 (continued) Exercise 11.9 (continued) Analysis (continued) Ordinary Distillation Column: This column separates the bottoms from the extractive distillation column to obtain 98 mol% methanol and 99.9 mol% ethanol. A starting point for this separation is obtained by using the ShortCut Column (FUG) model of Chemcad with an operating reflux ratio of 1.3 times the minimum ratio. The feed is assumed to be at the bubble point. To achieve the purity specifications, a preliminary material balance gives: mol/s: Component Feed Distillate Bottoms Acetone 0.13 0.13 0.00 Methanol 9.89 9.82 0.07 Ethanol 68.76 0.07 68.69 Total: 78.78 10.02 68.76 The specifications for the ShortCut Column model are 0.982/0.989 = 0.993 for the LK (methanol) and 0.07/68.76 = 0.001 for the HK (ethanol). The results are: Minimum number of equilibrium stages = 24.6 Minimum reflux ratio = 11.9 Reflux ratio = 1.3(11.9) = 15.5 Number of equilibrium stages = 44.4 Feed stage location (Kirkbride equation) = 31 from the top The SCDS Column model of the Chemcad program was then used to make a rigorous calculation based on the following input data: Pressure constant at 1 atm Number of stages = 46 (includes the total condenser and the partial reboiler) Feed stage = 32 Methanol purity in the distillate = 98 mol% Ethanol purity in the bottoms = 99.9 mol% Estimated distillate rate = 36.1 kmol/h (10.03 mol/s) Estimated reflux rate = 560 kmol/h (155.6 mol/s or a reflux ratio of 15.5) Estimated temperatures: Stage 1, 64oC Stage 46, 78oC Stage 2, 65oC Stage 45, 75oC A converged result was obtained in 6 iterations, with a computed reflux ratio of 14.9, compared to the above 15.5 estimate by the Underwood equations. However, as shown in the composition profile on the next page, a pinch region (where no composition changes occur) of approximately 7 stages is evident just above the feed stage. Accordingly, the calculations with the SCDS Column model were repeated by removing 7 stages above the feed, resulting in the following new input: Number of stages = 39 (includes the total condenser and the partial reboiler) Feed stage = 25 A converged result was obtained in 5 iterations, with a computed reflux ratio of 15.0, just slightly greater. The composition profile for this run also appears on the next page. No pinch is evident. Analysis (continued) Exercise 11.9 (continued) Analysis (continued) Exercise 11.9 (continued) The ordinary distillation column for separating methanol from ethanol was sized with Chemcad for sieve trays, with: 24-inch tray spacing 12% downcomer area 10% hole area 1/4-inch holes 80% of flooding The resulting column inside diameter was determined to be 6.5 feet, which gives 60.0% of flooding at the top tray and 72.4% of flooding at the bottom. The tray pressure drop averages 0.091 psi/tray, for a total of 3.4 psi or 0.23 atm for the 37 trays. Thus, the assumption of negligible tray pressure drop is not satisfactory. For a final design, the computations should be repeated taking into account tray pressure drop and tray efficiency. The final material balance for the ordinary distillation column was as follows: Component Acetone Methanol Ethanol Total: mol/s: Feed 0.13 9.89 68.76 78.78 Distillate 0.13 9.82 0.07 10.02 Bottoms 0.00 0.07 68.69 68.76 Because the flow rate of ethanol in the bottoms is only 68.69 mol/s and 70.00 is needed for the extractive distillation column, the makeup ethanol rate is 70.00 - 68.69 = 1.31 mol/s. Comparison with Example 11.3: In Example 11.3, water is used as the solvent. It is much more effective than the ethanol solvent used in this exercise, as shown by the following comparison: Item Solvent flow rate, mol/s Solvent makeup rate, mol/s Equilibrium stages in extractive distillation column Equilibrium stages in ordinary distillation column Diameter of extractive distillation column, feet Diameter of ordinary distillation column, feet Water solvent 60 1.4 28 16 6 2.5 Ethanol solvent 70 1.31 56 36 7 6.5 Exercise 11.10 Subject: Separation of acetone from methanol by extractive distillation with methyl-ethylketone (MEK) Given: Bubble-point feed of 30 mol/s acetone (A) and 10 mol/s methanol (M) at 1 atm. MEK (K) as the solvent. Results of Example 11.3 with a solvent of water. Assumptions: Because the feed is close to the azeotropic composition of 22 mol% methanol, use a two-column system, with the first column being extractive distillation and the second ordinary distillation to recover the solvent. UNIFAC for K-values. Negligible tray pressure drop. 100% tray efficiency. Find: Suitable column designs to obtain an acetone product of at least 95 mol%, a methanol product of at least 98 mol%, and MEK of 99.9 mol% purity for recycle. Analysis: Extractive Distillation Column: First compute relative volatilities over a composition range to determine: (1) which of the two feed components is the most volatile in the presence of MEK and (2) potential ease or difficulty of the separations. This is done with the LLVF threephase flash model of Chemcad, which also detects any formation of a second liquid phase, by conducting a series of flashes at 1 atm and a V/F = 0.0001 (a near bubble-point condition), for different mole fractions of the two feed components. For all cases, the mole fraction of the solvent, MEK, is fixed at 0.5 so as to have a strong effect on the liquid-phase activity coefficients of acetone and methanol. The flash model computes the K-values, from which the relative volatilities can be obtained as given below. Since acetone and methanol boil at 56.5oC and 64.7oC, respectively, but acetone is closer in structure to MEK than is methanol, it is difficult to predict which of the two feed components will be the more volatile. At high concentrations of acetone, methanol is more volatile, while at high concentrations of methanol, acetone is more volatile. xA 0.475 0.375 0.250 0.125 0.025 KA 1.345 1.310 1.331 1.413 1.524 KM 1.877 1.616 1.415 1.301 1.255 KK 0.629 0.613 0.627 0.671 0.731 M-A 1.40 1.23 1.06 0.92 0.82 M,K 2.98 2.64 2.26 1.94 1.72 Of greater importance, however, is the fact that UNIFAC predicts, at 1 atm, that MEK forms a minimum-boiling azeotrope with methanol at 64.0oC, with a mole fraction of methanol equal to 0.844. Thus, MEK is a questionable solvent for extractive distillation because on the ternary diagram for 1 atm pressure, a distillation boundary will connect the A-M azeotrope with the M-K azeotrope, similar to Fig. 11.5, where a distillation boundary connects the benzene-isopropanol Exercise 11.10 Analysis: (Extractive Distillation Column continued) azeotrope with the benzene-n-propanol azeotrope. For extractive distillation, a residue curve map similar to that in Fig. 11.14 is needed. As shown in Fig. 11.22, some azeotropes are pressure sensitive, such that if the pressure is changed sufficiently the azeotrope disappears. As shown in the following table, based on the UNIFAC equation, the M-K azeotrope shows a favorable shift with pressure, while the A-M azeotrope does not. Above about 6 atm, the methanol-MEK azeotrope disappears. Pressure, atm A-M azeotrope M-K azeotrope 0.3 1 3 5 10 15 Mole% methanol in azeotrope: 6 22 40 50 62 70 74 84 94 99 25 80 - Thus, as shown by Knapp and Doherty in Ref. 34 of Chapter 11, a mixture of acetone and methanol can be separated by extractive distillation with MEK by operating the extractive distillation column (Column 1) at an elevated pressure so as eliminate the M-K azeotrope and to obtain, as will be shown, a distillate of 98 mol% methanol, followed by sending the bottoms to ordinary distillation (Column 2) at 1 atm to obtain a distillate of 95 mol% acetone and a bottoms of 99.9 mol% MEK for recycle. The two columns were designed with the SCDS (simultaneous correction) model of Chemcad leading to the following final material balance, which satisfies the product purity specifications: Component Acetone Methanol MEK Total: Temp, oC Mol/s: Feed Solvent 30 10 40 144.9 Column 1 distillate 0.016 8.526 0.158 8.700 144.6 Column 1 bottoms 29.984 1.474 59.842 91.300 169.5 Column 2 distillate 29.924 1.474 0.101 31.499 55.8 Column 2 bottoms 0.060 0.000 59.741 59.801 79.1 60 60 145 The pressure of the extractive distillation column (Column 1) was set at 12 atm, with a solvent rate of 60 mol/s at 145oC. The following Chemcad SCDS results, for the listed conditions, was obtained after 11 iterations. The resulting product compositions are those in the above table. Other combinations of stages, reflux, and solvent rate could also give satisfactory results: Number of stages = 60 (includes total condenser and partial reboiler) Solvent enters stage 20 from the top Feed enters stage 40 from the top Acetone mole fraction in distillate = 0.98 Bottoms flow rate = 91.3 mol/s Solvent flow rate = 60 mol/s at 145oC Initial guesses: Distillate rate = 8 mol/s Reflux rate = 80 mol/s Temperatures: Stage 1, 145oC Stage 60, 170oC Stage 2, 145oC Stage 59, 160oC Exercise 11.10 Analysis: (Extractive Distillation Column continued) The reflux ratio was computed to be 8.083. The liquid phase composition profile is shown below. The extractive distillation column was sized with Chemcad for sieve trays, with: 24-inch tray spacing 12% downcomer area 10% hole area 1/4-inch holes 80% of flooding The resulting column inside diameter was determined to be 5 feet. For a final design, the computations should be repeated taking into account tray pressure drop, tray efficiency, and impurity in the MEK solvent. Exercise 11.10 (continued) Analysis: (continued) Ordinary Distillation Column: This column, at a pressure of 1 atm, separates the bottoms from the extractive distillation column to obtain 95 mol% acetone and 99.9 mol% MEK. A starting point for this separation is obtained by using the ShortCut Column (FUG) model of Chemcad with an operating reflux ratio of 1.3 times the minimum ratio. The feed is assumed to be at the bubble point. Thus, the bottoms from Column 1 at 169.5oC and 12 atm is cooled and decompressed to 68.7oC and 1 atm. Using the above material balance table, the specifications for the ShortCut Column model are 29.924/29.984 = 0.998 for the recovery of the LK (methanol) to the distillate and 0.101/59.842 = 0.0169 for the recovery of the HK (MEK). The results are: Minimum number of equilibrium stages = 17.7 Minimum reflux ratio = 2.52 Reflux ratio = 1.3(2.52) = 3.28 Number of equilibrium stages = 34.1 Feed stage location (Kirkbride equation) = 18 from the top The SCDS Column model of the Chemcad program was then used to make a rigorous calculation based on the following input data: Pressure constant at 1 atm Number of stages = 35 (includes the total condenser and the partial reboiler) Feed stage = 20 from the top Acetone purity in the distillate = 95 mol% MEK purity in the bottoms = 99.9 mol% Estimated distillate rate = 31.4 mol/s Estimated reflux rate = 109.9 mol/s (reflux ratio of 3.5) Estimated temperatures: Stage 1, 64oC Stage 35, 78oC Stage 2, 65oC Stage 34, 75oC A converged result was obtained in 5 iterations, with a computed reflux ratio of 3.69, compared to the above 3.5 estimate by the Underwood equations. The composition profile appears on the next page. The ordinary distillation column for separating acetone from MEK was sized with Chemcad for sieve trays, with: 24-inch tray spacing 12% downcomer area 10% hole area 1/4-inch holes 80% of flooding The resulting column inside diameter was determined to be 7.0 feet. For a final design, the computations should be repeated taking into account tray pressure drop and tray efficiency. From the above material balance table, the flow rate of MEK in the bottoms of Column 2 is only 59.80 mol/s, while 60.00 is needed for the extractive distillation column, the makeup MEK rate is 60.00 - 59.74 = 0.26 mol/s. Analysis: (continued) Exercise 11.10 (continued) Comparison with Example 11.3: In Example 11.3, water is used as the solvent. The following comparison shows that water is preferred over MEK. Item Water solvent MEK solvent Solvent flow rate, mol/s 60 60 Solvent makeup rate, mol/s 1.4 0.26 Equilibrium stages in extractive distillation column 28 59 Equilibrium stages in ordinary distillation column 16 34 Diameter of extractive distillation column, feet 6 5 Diameter of ordinary distillation column, feet 2.5 7 Operating pressure in extractive distillation 1 12 Exercise 11.11 Subject: Separation of acetone from methanol by extractive distillation with toluene Given: Bubble-point feed of 30 mol/s acetone (A) and 10 mol/s methanol (M) at 1 atm. Toluene (T) as the solvent. Results of Example 11.3 with a solvent of water. Assumptions: Because the feed is close to the azeotropic composition of 22 mol% methanol, use a two-column system, with the first column being extractive distillation and the second ordinary distillation to recover the solvent. UNIFAC for K-values. Negligible tray pressure drop. 100% tray efficiency. Find: Suitable column designs to obtain an acetone product of at least 95 mol%, a methanol product of at least 98 mol%, and toluene of 99.9 mol% purity for recycle. Analysis: Of great importance here is the fact that UNIFAC predicts, at 1 atm, that toluene forms a minimum-boiling azeotrope with methanol at 63.8oC, with a mole fraction of methanol equal to 0.891. Thus, toluene is a questionable solvent for extractive distillation because on the residue curve map for 1 atm pressure, shown on the next page, a distillation boundary connects the A-M azeotrope with the M-T azeotrope, similar to Fig. 11.5, where a distillation boundary connects the benzene-isopropanol azeotrope with the benzene-n-propanol azeotrope. For extractive distillation, a residue curve map similar to that in Fig. 11.14 is needed. As shown in Fig. 11.22, some azeotropes are pressure sensitive, such that if the pressure is changed sufficiently the azeotrope disappears. Unfortunately, over the pressure range of 0.3 atm to 25 atm, neither the A-M azeotrope nor the M-T azeotrope disappears. Therefore, it is concluded that the use of toluene can not produce a distillate of either acetone or methanol by extractive distillation. To confirm this, two runs were made with the SCDS model of Chemcad for the conditions below. The runs differed only in the bottoms mole flow specification, with Case 1 being the sum of the toluene solvent and the acetone in the feed, in an attempt to obtain a distillate of methanol. For Case 2 the bottoms mole flow specification was the sum of the toluene and the methanol in the feed, in an attempt tot obtain a distillate of acetone. The results, given on the next page, confirm the effect of the distillation boundary. In Case 1, the distillate is not methanol, but is close to the distillation boundary near the M-T azeotrope. In Case 2, the distillate is not acetone, but is close to the distillation boundary and near the A-M azeotrope. Solvent (toluene) flow rate = 60 mol/s 60oC Pressure constant at 1 atm Number of stages = 46 (includes the total condenser and the partial reboiler) Solvent stage = 10 from the top Feed stage = 30 from the top Reflux ratio = 5 Bottoms mole flow rate: Case 1, 90 mol/s Case 2, 70 mol/s Estimated distillate rate: Case 1, 10 mol/s Case 2, 30 mol/s Estimated reflux rate: Case 1, 50 mol/s Case 2, 150mol/s (reflux ratio of 5) Estimated temperatures: Stage 1, 55oC Stage 46, 75oC Stage 2, 60oC Stage 45, 70oC Exercise 11.11 (continued) Analysis: (continued) Case 1 material balance: mol/s Component Solvent Acetone Methanol Toluene 60 Total: 60 Case 2 material balance: mol/s Component Solvent Acetone Methanol Toluene 60 Total: 60 Feed 30 10 40 Distillate 0.07 8.74 1.19 10.00 Bottoms 29.93 1.26 58.81 90.00 Feed 30 10 40 Distillate 19.98 10.00 0.02 30.00 Bottoms 10.02 0.00 59.98 70.00 For both cases, most of the methanol moves to the distillate because, in the presence of a substantial fraction of toluene, methanol is more volatile than acetone. Exercise 11.11 (continued) Analysis: (continued) Exercise 11.12 Subject: Separation of n-heptane from toluene by extractive distillation with phenol. Given: 400 lbmol/h of a feed of an equimolar mixture at 200oF and 20 psia of n-heptane and toluene. 1200 lbmol/h of phenol solvent at 20 psia and 220oF. Assumptions: . Column operations at 20 psia UNIFAC for K-values. Negligible tray pressure drop. 100% tray efficiency. Find: Suitable column designs to obtain reasonable product purities, with only a small loss of solvent. Analysis: Using the vapor pressure data in Chemcad, the boiling points at 20 psia are found to be: 229.1oF for n-heptane, 251.4oF for toluene, and 379.0oF for phenol. Thus, n-heptane appears to be the most volatile component. Although n-heptane and toluene do not form an azeotrope at 20 psia, their relative volatility at high concentrations of n-heptane is almost 1.0. Therefore, ordinary distillation of the mixture is not practical. Phenol does not form an azeotrope with either n-heptane or toluene, and with its higher boiling point is a possible solvent. Its selection is further enhanced because, as might be expected on molecular structure considerations, phenol at say a mole fraction in the liquid of 0.5 causes a larger liquid-phase activity coefficient for nheptane than for toluene. For example, for 50 mol% phenol in the liquid and at low concentrations of toluene, the relative volatility of n-heptane with respect to toluene is about 1.7. At low concentrations of n-heptane and 50 mol% phenol in the liquid, the relative volatility of nheptane with respect to toluene is even higher at about 2.8. Thus, phenol makes the separation of n-heptane and toluene quite favorable. Extractive Distillation Column: The following Chemcad results, using the Tower (inside-out) model, give a reasonable n-heptane distillate product. Convergence was achieved in 15 iterations. Other combinations of stages, reflux, and solvent and feed stage locations could also give satisfactory results. Number of stages = 21 (includes total condenser and partial reboiler) Solvent enters stage 6 from the top Feed enters stage 13 from the top Reflux ratio = 5 Bottoms flow rate = 1400 lbmol/h (based on toluene + phenol) Initial guesses: Distillate rate = 200 lbmol/h Reflux rate = 1000 lbmol/h o o Temperatures: Stage 1, 200 F Stage 21, 300 F Stage 2, 220oF Convergence was achieved in 34 iterations with the following material balance: Solvent Feed Distillate Bottoms Temperature, oF 220 200 229.2 323.4 Flow rates, lbmol/h: n-Heptane 0 200 198.2 1.8 Toluene 0 200 1.2 198.8 Phenol 1200 0 0.6 1199.4 Exercise 11.12 (continued) Analysis: Extractive Distillation Column (continued) The liquid phase composition profile is shown below. The condenser and reboiler duties are 15.99 million Btu/h and 24.30 milllion Btu/h, respectively. The extractive distillation column was sized with Chemcad for sieve trays, with: 24-inch tray spacing 12% downcomer area 10% hole area 1/4-inch holes 80% of flooding The resulting column inside diameter was determined to be 8 feet For a final design, the computations should be repeated taking into account tray pressure drop and tray efficiency. Exercise 11.12 (continued) Analysis: (continued) Ordinary Distillation Column to separate toluene from phenol: A starting point for this separation is obtained by using the ShortCut Column (FUG) model of Chemcad with an operating reflux ratio of 1.3 times the minimum ratio, with an operating pressure of 20 psia. The feed is assumed to be at the bubble point. To achieve reasonable product purities, the following material balance is assumed: Component n-Heptane Toluene Phenol Total: lbmol/h: Feed Distillate 1.8 1.8 198.8 196.8 1199.4 1.2 1400.0 199.8 Bottoms 0.00 2.0 1198.2 1200.2 The specifications for the ShortCut Column model are 196.8/198.8 = 0.99 for the recovery of toluene to the distillate, and 1.2/1199.4 = 0.001 for the recovery to the distillate of phenol. The results are: Minimum number of equilibrium stages = 7.7 Minimum reflux ratio = 1.50 Reflux ratio = 1.3(1.50) = 1.95 Number of equilibrium stages = 15.5 Feed stage location (Kirkbride equation) = 11.4 from the top The Tower model of the Chemcad program is then used to make a rigorous calculation based on the following input data, using the FUG results: Pressure constant at 20 psia. Number of stages = 16 (includes the total condenser and the partial reboiler) Feed stage = 12 from the top Reflux ratio = 2 Bottoms mole flow rate = 1200 lbmol/h Estimated distillate rate = 200 lbmol/h Estimated reflux rate = 400 lbmol/h Estimated temperatures: Stage 1, 200oF Stage 16, 300oF Stage 2, 220oF A converged result, which showed a slightly better separation than specified for the FUG method, was obtained in 17 iterations: However, as shown in the composition profile on the next page, a pinch region (where no composition changes occur) of approximately 3 stages is evident just above the feed stage. Accordingly, the calculations with the Tower model were repeated by removing 3 stages above the feed, resulting in the following new input: Number of stages = 13 (includes the total condenser and the partial reboiler) Feed stage = 9 Exercise 11.12 (continued) Analysis: Ordinary Distillation Column (continued) A converged result was obtained in 17 iterations, with little change in the separation. The composition profile for this run appears on the next page. No pinch is evident. The ordinary distillation column for separating methanol from ethanol was sized with Chemcad for sieve trays, with: 24-inch tray spacing 12% downcomer area 10% hole area 1/4-inch holes 80% of flooding The resulting column inside diameter was determined to be 5.5 feet. The condenser duty was computed to be 8.53 million Btu/h, with a reboiler duty of 11.70 million Btu/h. For a final design, the computations should be repeated taking into account tray pressure drop and tray efficiency. The final material balance for the ordinary distillation column was as follows: Component n-Heptane Toluene Phenol Total: Temp., oF lbmol/h: Feed Distillate 1.8 1.8 198.8 197.0 1199.4 1.2 1400.0 200.0 323.3 251.0 Bottoms 0.00 1.8 1198.2 1200.0 378.2 Because the flow rate of phenol in the bottoms is only 1198.2 lbmol/h and 1200.0 is needed for the extractive distillation column, the makeup phenol rate is 1200.0 - 1198.2 = 1.8 lbmol/h. Exercise 11.12 (continued) Analysis: Ordinary Distillation Column (continued) Exercise 11.12 (continued) Analysis: Ordinary Distillation Column (continued) Exercise 11.13 Subject: Separation of a mixture of ethanol and benzene by pressure-swing distillation. Given: Feed of 100 mol/s of 55 mol% ethanol and 45 mol% benzene at the bubble point at 101.3 kPa. Desired products are 99 mol% ethanol and 99 mol% benzene. Assumptions: UNIFAC method for estimating K-values. Find: Feasible design using pressure-swing distillation in the manner of Example 11.5 and Fig. 11.23(b). Analysis: Use the two-column system of Fig. 11.23(b). Ethanol and benzene form a minimumboiling azeotrope. Use the same column pressures as in Example 11.5. Therefore, in Column 1, the pressures are 26 kPa at the condenser outlet, 30 kPa at the top tray, and 40 kPa at the reboiler. The distillate product is the near azeotrope at 30 kPa and the bottoms product is 99 mol% ethanol. In Column 2, the pressures are 106kPa before the condenser, 101.3 kPa after the condenser, and 120 kPa at the reboiler. The distillate product is the near azeotrope at 101.3 kPa and the bottoms product is 99 mol% benzene. Using the Chemcad program with the UNIFAC method, the azeotrope compositions in mol% ethanol are 35.7 at 26 kPa, and 45.2 at 101.3 kPa. A major factor in the design of a pressure-swing distillation system is the recycle-to-feed ratio, D2/F in Fig. 11.23(b), which is related to compositions as follows: Let x be the mole fraction of ethanol in any stream. Referring to Fig. 11.23(b), an overall system material balance for total flows gives: F = B1 + B2 (1) (2) An overall system balance on ethanol gives: x F F = x B1 B1 + x B2 B2 A total material balance around Column 2 gives: D1 = D2 + B 2 An ethanol balance around Column 2 gives: x D1 D1 = x D2 D2 + x B2 B2 Substitute Eq. (1) into (2) to eliminate B1. After rearrangement, x B1 - x B2 F = B2 x B1 - x F Substitute Eq. (3) into (4) to eliminate D1. After rearrangement, x D1 - x B2 D2 = B2 x D2 - x D1 (3) (4) (5) (6) Combine Eqs. (5) and (6) to eliminate B1. After rearrangement, the recycle ratio is: x B1 - x F x D1 - x B2 D2 = (7) F x B1 - x B2 x D2 - x D1 In the limit, for pure bottoms products and azeotropic distillate products, Eq. (7) reduces to: xAz1 D2 = 1 - xF (8) F xAz 2 - xAz1 Exercise 11.13 (continued) Analysis: (continued) Eq. (8) shows that the recycle ratio is sensitive to the feed composition and the azeotropic compositions at the pressures of Columns 1 and 2. The smaller the difference between the two azeotropic compositions, the larger the recycle ratio. For this exercise, x F = 0.55, x B1 = 0.99, and x B2 = 0.01 Let x D1 = 0.37 , which is slightly greater than the azeotropic mole fraction of 0.357. Let x D2 = 0.44 , which is slightly smaller than the azeotropic mole fraction of 0.452. Substituting these 5 mole-fraction values into Eq. (7) gives D2/F = 2.309. Therefore D2 = 230.90 mol/s. Using Eqs. (1) to (6), the following material balance is obtained for the system: Stream: Flow rate, mol/s: Ethanol Benzene Total: mol% Ethanol: F 55 45 100 55 D2 101.60 129.30 230.90 44 F1 156.60 174.30 330.90 47.33 B1 54.55 0.55 55.10 99 D1 102.05 173.75 275.80 37 B2 0.45 44.45 44.90 1 The design of the two columns was made with the Chemcad program, using the Shortcut Column model (FUG method) to obtain initial estimates of stage and reflux requirements, followed by the SCDS Column or TOWER rigorous model to finalize the designs. Column 1: Using the above material balance, with a combined feed, F1, the specifications for the ShortCut Column model are 173.75/174.3 = 0.997 for the recovery of benzene to the distillate, and 102.05/156.6 = 0.652 for the recovery to the distillate of ethanol. Thus, it is expected that little rectification is necessary. The results are: Minimum number of equilibrium stages = 6.23 Minimum reflux ratio = very small and probably in error. Therefore, the Gilliland correlation was not reliable. The SCDS Column model of the Chemcad program was then used to make a rigorous calculation based initially on the following input data, using the FUG results, with N = 2 x Nmin. A guess of R is made: Pressures as stated above. Number of stages = 12 (includes the total condenser and the partial reboiler) Feed stage for recycle, D2 = 4 from the top Entry for feed, F = 6 from top Reflux ratio = 0.5 Bottoms mole flow rate = 55.1 mol/s Estimated distillate rate = 275.76 mol/s Estimated reflux rate = 138 mol/s Estimated temperatures: Stage 1, 35oC Stage 12, 40oC Stage 2, 36oC Stage 11, 39oC Exercise 11.13 (continued) Analysis: Column 1 (continued) A converged result, which gave a 99.8 mol% purity of ethanol in the bottoms, was obtained in 10 iterations. A purity of only 99 mol% was needed. This indicated that the reflux could be reduced. Accordingly, using the profiles from the first rigorous run, a second run was made with a column top specification of D = 173.75 mol/s in place of the reflux specification. A converged result, with the desired 99 mol% ethanol purity in B1, was obtained at a reflux ratio of 0.279 in 24 iterations. The material balance was in agreement with the above table. Although the feed locations could probably be improved, no pinch region is evident, as shown in the liquid-phase composition profile on the next page. Also included on the next page are the column profiles for temperature, liquid flow, and vapor flow. Column 1 was sized with Chemcad for sieve trays, with: 24-inch tray spacing 12% downcomer area 10% hole area 1/4-inch holes 80% of flooding The resulting column inside diameter was determined to be 12 feet. The condenser duty was computed to be 12.91 MW, with a reboiler duty of 11.58 MW. For the 10 theoretical plates, the total pressure drop was computed to be 8.2 kPa. For a final design, the computations should be repeated taking into account tray pressure drop and tray efficiency. Column 2: Using the above material balance table, the specifications for the ShortCut Column model are 101.60/102.05 = 0.9956 for the recovery of ethanol to the distillate, and 129.30/173.75 = 0.7442 for the recovery to the distillate of benzene. Thus, it is again expected that little rectification is necessary. The feed to Column 2 is the distillate from Column 1 after being pumped to a pressure of 120 kPa for entry into Column 2 and heating to the bubble point at that pressure. The results are: Minimum number of equilibrium stages = 5.0 Minimum reflux ratio = very small and probably in error. Therefore, the Gilliland correlation was not reliable. The SCDS Column model of the Chemcad program experienced convergence difficulties and was replaced with the TOWER model, which also experienced difficulty until the following specifications were made with a damping factor of 0.7. Pressures as stated above. Number of stages = 6 (includes the total condenser and the partial reboiler) Entry for feed = 2 from top (giving in effect a reboiled stripper with some reflux) Distillate flow rate of ethanol = 101.60 mol/s Bottoms mole flow rate = 44.9 mol/s Estimated distillate rate = 230.90 mol/s Estimated reflux rate = 23 mol/s (corresponding to a small reflux ratio of 0.1) Estimated temperatures: Stage 1, 67oC Stage 6, 80oC Stage 2, 70oC Analysis: Column 1 (continued) Exercise 11.13 (continued) Exercise 11.13 (continued) Analysis: Column 2 (continued) A converged result was obtained in 21 iterations, with a small reflux ratio of 0.16 or 36.9 mol/s. The material balance was in agreement with the above table. Although the feed location could probably be improved, no pinch region is evident, as shown in the liquid-phase composition profile below. Also included are the column profiles for temperature, liquid flow, and vapor flow. Column 2 was sized with Chemcad for sieve trays, with: 24-inch tray spacing 12% downcomer area 10% hole area 1/4-inch holes 80% of flooding The resulting column inside diameter was determined to be 9 feet. The condenser duty was computed to be 9.38 MW, with a reboiler duty of 9.30 MW. For the 4 theoretical plates, the total pressure drop was computed to be 3.9kPa. For a final design, the computations should be repeated taking into account tray pressure drop and tray efficiency. Exercise 11.13 (continued) Analysis: Column 2 (continued) Exercise 11.14 Subject: Separation of a mixture of ethanol and benzene by pressure-swing distillation. Given: Feed of 100 mol/s of 20 mol% ethanol and 80 mol% benzene at the bubble point at 101.3 kPa. Desired products are 99 mol% ethanol and 99 mol% benzene. Assumptions: UNIFAC method for estimating K-values. Find: Feasible design using pressure-swing distillation in the manner of Example 11.5. Analysis: Use the two-column system of Fig. 11.23(b). Ethanol and benzene form a minimumboiling azeotrope. Use the same column pressures as in Example 11.5. Therefore, in Column 1, the pressures are 26 kPa at the condenser outlet, 30 kPa at the top tray, and 40 kPa at the reboiler. The distillate product is the near azeotrope at 30 kPa and the bottoms product is 99 mol% ethanol. In Column 2, the pressures are 106kPa before the condenser, 101.3 kPa after the condenser, and 120 kPa at the reboiler. The distillate product is the near azeotrope at 101.3 kPa and the bottoms product is 99 mol% benzene. Using the Chemcad program with the UNIFAC method, the azeotrope compositions in mol% ethanol are 35.7 at 26 kPa, and 45.2 at 101.3 kPa. A major factor in the design of a pressure-swing distillation system is the recycle-to-feed ratio, D2/F in Fig. 11.23(b), which is related to compositions as follows: Let x be the mole fraction of ethanol in any stream. Referring to Fig. 11.23(b), an overall system material balance for total flows gives: F = B1 + B2 (1) (2) An overall system balance on ethanol gives: x F F = x B1 B1 + x B2 B2 A total material balance around Column 2 gives: D1 = D2 + B 2 An ethanol balance around Column 2 gives: x D1 D1 = x D2 D2 + x B2 B2 Substitute Eq. (1) into (2) to eliminate B1. After rearrangement, x B1 - x B2 F = B2 x B1 - x F Substitute Eq. (3) into (4) to eliminate D1. After rearrangement, x D1 - x B2 D2 = B2 x D2 - x D1 (3) (4) (5) (6) Combine Eqs. (5) and (6) to eliminate B1. After rearrangement, the recycle ratio is: x B1 - x F x D1 - x B2 D2 = (7) F x B1 - x B2 x D2 - x D1 In the limit, for pure bottoms products and azeotropic distillate products, Eq. (7) reduces to: xAz1 D2 = 1 - xF (8) F xAz 2 - xAz1 Exercise 11.14 (continued) Analysis: (continued) Eq. (8) shows that the recycle ratio is sensitive to the feed composition and the azeotropic compositions at the pressures of Columns 1 and 2. The smaller the difference between the two azeotropic compositions, the larger the recycle ratio. For this exercise, x F = 0.20, x B1 = 0.99, and x B2 = 0.01 Let x D1 = 0.37 , which is slightly greater than the azeotropic mole fraction of 0.357. Let x D2 = 0.44 , which is slightly smaller than the azeotropic mole fraction of 0.452. Substituting these 5 mole-fraction values into Eq. (7) gives D2/F = 4.1457. Therefore D2 = 414.57 mol/s. This a very high recycle ratio. The system is technically feasible, but may not be economically attractive. Using Eqs. (1) to (6), the following material balance is obtained for the system: F Stream: D2 F1 B1 D1 B2 Flow rate, mol/s: Ethanol 20 182.41 202.41 19.19 183.22 0.81 Benzene 80 232.16 312.16 0.20 311.96 79.80 Total: 100 414.57 514.57 19.39 495.18 80.61 mol% Ethanol: 20 44 39.3 99 37 1 The design of the two columns is made as follows with the Chemcad program, using the Shortcut Column model (FUG method) to obtain initial estimates of stage and reflux requirements, followed by the SCDS Column rigorous model to finalize the designs. Column 1: Using the above material balance, with a combined feed, F1, the specifications for the ShortCut Column model are 311.96/312.96 = 0.9968 for the recovery of benzene to the distillate, and 183.22/202.41 = 0.9052 for the recovery to the distillate of ethanol. Thus, it is expected that little rectification is necessary. The results are: Minimum number of equilibrium stages = 5.95 Minimum reflux ratio = very small and probably in error. Therefore, the Gilliland correlation was not reliable. The SCDS Column model of the Chemcad program was then used to make a rigorous calculation. Based on several runs where the number of stages, the feed stage location, and the reflux ratio were varied, it was found that the separation was almost insensitive to the reflux ratio, indicating that little rectification was needed. However, attempts to configure the column as a reboiled stripper met with failure to converge. The following input data resulted in almost satisfying the specifications. Because the separation depended mainly on the number of stages, the specified separation could not be satisfied exactly. Pressures as stated above. Number of stages = 7 (includes the total condenser and the partial reboiler) Feed stage for recycle, D2 = 3 from the top Entry for feed, F = 4 from top Reflux ratio = 0.05 Bottoms mole flow rate = 19.39 mol/s Exercise 11.14 (continued) Analysis: Column 1 (continued) Estimated distillate rate = 495.18 mol/s Estimated reflux rate = 24.75 mol/s Estimated temperatures: Stage 1, 35oC Stage 12, 55oC Stage 2, 40oC Stage 11, 50oC A converged result, which gave just slightly less than the desired 99 mol% purity of ethanol in the bottoms, was obtained in 4 iterations. The material balance was almost in agreement with the above table. Although the feed locations could probably be improved, no pinch region is evident, as shown in the liquid-phase composition profile below. Also included on the this page are the column profiles for temperature, liquid flow, and vapor flow. Column 1 was sized with Chemcad for sieve trays, with: 24-inch tray spacing 12% downcomer area 10% hole area 1/4-inch holes 80% of flooding The resulting column inside diameter was determined to be 14 feet. The condenser duty was computed to be 19.03 MW, with a reboiler duty of 16.42 MW. For the 5 theoretical plates, the total pressure drop was computed to be 4.9 kPa. For a final design, the computations should be repeated taking into account tray pressure drop and tray efficiency. Exercise 11.14 (continued) Analysis: Column 1 (continued) Exercise 11.14 (continued) Analysis: (continued) Column 2: Using the above material balance table, the specifications for the ShortCut Column model are 182.41/183.22 = 0.9956 for the recovery of ethanol to the distillate, and 232.16/173.75 = 0.7442 for the recovery to the distillate of benzene. Thus, it is again expected that little rectification is necessary. The feed to Column 2 is the distillate from Column 1 after being pumped to a pressure of 120 kPa for entry into Column 2 and heating to the bubble point at that pressure. The results are: Minimum number of equilibrium stages = 5.0 Minimum reflux ratio = very small and probably in error. Therefore, the Gilliland correlation was not reliable. The SCDS Column model of the Chemcad program experienced convergence difficulties and was replaced with the TOWER model, which also experienced difficulty until the following specifications were made with a damping factor of 0.7. Pressures as stated above. Number of stages = 6 (includes the total condenser and the partial reboiler) Entry for feed = 2 from top (giving in effect a reboiled stripper with some reflux) Distillate flow rate of ethanol = 182.41 mol/s Bottoms mole flow rate = 80.61 mol/s Estimated distillate rate = 414.57 mol/s Estimated reflux rate = 41 mol/s (corresponding to a small reflux ratio of 0.1) Estimated temperatures: Stage 1, 67oC Stage 6, 80oC Stage 2, 70oC A converged result was obtained in 20 iterations, with a small reflux ratio of 0.155 or 64.1 mol/s. The material balance was in agreement with the above table. Although the feed location could probably be improved, no pinch region is evident, as shown in the liquid-phase composition profile on the next page. Also included on the next page are the column profiles for temperature, liquid flow, and vapor flow. Column 2 was sized with Chemcad for sieve trays, with: 24-inch tray spacing 12% downcomer area 10% hole area 1/4-inch holes 80% of flooding The resulting column inside diameter was determined to be 12 feet. The condenser duty was computed to be 16.76 MW, with a reboiler duty of 16.62 MW. For the 4 theoretical plates, the total pressure drop was computed to be 4.0 kPa. For a final design, the computations should be repeated taking into account tray pressure drop and tray efficiency. Exercise 11.14 (continued) Analysis: Column 2 (continued) Exercise 11.15 Subject: Separation of a mixture of ethanol and water by pressure-swing distillation. Given: Feed of 100 mol/s of 30 mol% ethanol and 70 mol% water. Desired product is 99.8 mol% ethanol.. Assumptions: Feed at the bubble point at 120 kPa. Use of NRTL equation to obtain K-values using modified coefficients given below. Find: Feasible design using pressure-swing distillation in the manner of Fig. 11.23(b). Analysis: Use the two-column system of Fig. 11.23(b), except that for this system, the bottoms from Column 1 is nearly pure B (water), while the bottoms from Column 2 is nearly pure A (ethanol). Ethanol and water form a minimum-boiling azeotrope. Using Chemcad, the NRTL equation and the following binary interaction parameters, which were entered to override the stored parameters (I = ethanol, J= water, Bij = 246, Bji = -586, alpha = 0.3, Aij = -0.8, and Aji = 3.45), the predicted azeotropes are 89.3 mol% ethanol at 1 atm, 85.6 mol% ethanol at 300 kPa, and 94.3 mol% ethanol at 13 kPa. Agreement with experimental data is excellent at 1 atm, but not accurate at 13 kPa torr, where the azeotrope composition is about 97 mol%. Thus, the following calculations using the NRTL equation will not be as accurate as would be desired. For Column 1, set the pressures at 13 kPa at the condenser outlet, 15 kPa at the top tray, and 30 kPa at the reboiler. Specify a Column 1 distillation composition of 92.5 mol% ethanol compared to the predicted azeotropic composition of 94.3 mol%. Let the ethanol composition of the bottoms from Column 1 be 0.0001 mol% ethanol, because, as will be shown, it is easy to obtain a very high water purity. For Column 2, set the pressures at 300 kPa at the condenser outlet, 305 kPa at the top tray, and 375 kPa at the reboiler. Specify a distillation composition of 88 mol% ethanol compared to the predicted azeotrope of 85.6 mol% ethanol. The bottoms from Column 2 is specified as 99.8 mol% ethanol. A major factor in the design of a pressure-swing distillation system is the recycle-to-feed ratio, D2/F in Fig. 11.23(b), which is related to compositions as follows: Let x be the mole fraction of ethanol in any stream. Referring to Fig. 11.23(b), an overall system material balance for total flows gives: F = B1 + B2 (1) An overall system balance on ethanol gives: x F F = x B1 B1 + x B2 B2 (2) A total material balance around Column 2 gives: D1 = D2 + B 2 An ethanol balance around Column 2 gives: x D1 D1 = x D2 D2 + x B2 B2 Substitute Eq. (1) into (2) to eliminate B1. After rearrangement, x B1 - x B2 F = B2 (5) x B1 - x F (3) (4) Exercise 11.15 (continued) Analysis: (continued) Substitute Eq. (3) into (4) to eliminate D1. After rearrangement, x D1 - x B2 D2 = B2 (6) x D2 - x D1 Combine Eqs. (5) and (6) to eliminate B1. After rearrangement, the recycle ratio is: x B1 - x F x D1 - x B2 D2 = (7) F x B1 - x B2 x D2 - x D1 In the limit, for pure bottoms products and azeotropic distillate products, Eq. (7) reduces to: 1 - xAz1 D2 = xF (8) F xAz1 - xAz 2 Eq. (8) shows that the recycle ratio is sensitive to the feed composition and the azeotropic compositions at the pressures of Columns 1 and 2. The smaller the difference between the two azeotropic compositions, the larger the recycle ratio. For this exercise, x F = 0.30, x B1 = 0.0001, and x B2 = 0.998 x D1 = 0.925 , which is slightly greater than the azeotropic mole fraction of 0.943. x D2 = 0.88 , which is slightly smaller than the azeotropic mole fraction of 0.856. Substituting these 5 mole-fraction values into Eq. (7) gives D2/F = 0.4876. Therefore D2 = 48.76 mol/s. This a modest recycle ratio. Using Eqs. (1) to (6), the following material balance is obtained for the system: F Stream: D2 F1 B1 D1 B2 Flow rate, mol/s: Ethanol 30 42.91 72.91 0.00007 72.91 30.00 Water 70 5.85 75.85 69.94 5.91 0.06 Total: 100 48.76 148.76 69.94 78.82 30.06 mol% Ethanol: 30 88 49.0 0.0001 92.5 99.8 The design of the two columns was made as follows with the Chemcad program, by making rough estimates of tray and reflux requirements, followed by use of the SCDS Column or Tower rigorous models to finalize the designs. Column 1: Using the TPXY plot feature of Chemcad, it was found that the relative volatility, , for ethanol with respect to water, at 13 kPa, varies widely over the composition range, from 1.020 at 92.5 mol% ethanol to 12.0 at 0.0001 mol% ethanol, both mole fractions in the liquid phase. Accordingly, it was believed that the FUG method would not give reliable estimates of stages and reflux. To obtain rough estimates of the stage requirements and feed locations, the Fenske equation (9-11) was applied to each of the three sections of the column, as summarized in the following table, using a pressure of 13 kPa rather than taking into account pressure drop. Exercise 11.15 (continued) Analysis: Column 1 (continued) Mole fraction ethanol in liquid 0.925 0.88 0.30 0.000001 Relative Geometric volatility average relative ethanol to volatility water 1.020 1.05 1.086 1.97 3.587 6.56 12.0 Minimum stages by Fenske equation 10.7 2.2 5.2 Location or Column section Overhead Overhead to Feed 1 Feed 1 Feed 1 to Feed 2 Feed 2 Feed 2 to Bottoms Bottoms The minimum stages were multiplied by about 2 to obtain the following estimates of equilibrium stages for input to the SCDS model: Total number of stages = 2(10.7 + 2.2 + 5.2) = 36.2. This was increased to 41 + total condenser = 42 to be on the conservative side. Feed 1, the recycle from Column 2, entering at 2(10.7) = 21.4, increased to Stage 23. Feed 2, the fresh feed, entering at 2(2.2) = 4.4 stages further down, changed to Stage 35 after it was found that the section between the two feeds was dominated by the lower end of the range of relative volatility. Determination of a reasonable estimate for the reflux ratio was difficult. A rough estimate was made with the Class I Underwood equation (9-20), using Feed 1, with x = ethanol mole fraction 13 . R = 13Rmin = . xD 1 - xD - EtOH,H 2 O x F1 1 - x F1 EtOH,H 2 O - 1 13 . = 0.925 1 - 0.925 - 105 . 0.88 1 - 0.88 1.05 - 1 = 10.3 Since the relative volatility in the section between the overhead and the Feed 1 entry is close to one, the reflux ratio is sensitive to relative volatility. Accordingly, a conservative reflux ratio estimate of 15 was used. The specifications for running Chemcad with the SCDS model were as follows: Feed 1 at the bubble point at 120 kPa. Feed 2 at the bubble point at 120 kPa. Thus, both feeds would become partially vaporized upon entry into Column 1 operating under vacuum. Column P = 15 kPa Top pressure = 13 kPa Condenser P = 2 kPa No. of stages = 42 (includes total condenser and partial reboiler) 1st feed stage = 23 2nd feed stage = 35 Distillate flow rate = 78.82 mol/s (from above material balance) Molar flow rate of ethanol in bottoms = 0.00007 mol/s Estimated distillate rate = 78.82 mol/s (specified value) Estimated reflux rate = 15D = 15(78.82) = 1182 mol/s Estimated temperatures: Stage 1, 35oC Stage 12, 60oC Stage 2, 40oC Stage 11, 55oC Exercise 11.15 (continued) Analysis: Column 1 (continued) A converged result, was obtained in 5 iterations, with a resulting reflux ratio of 13.47. The material balance was in agreement with the above material balance table for the system. Although the feed locations could probably be improved, no pinch region is evident, as shown in the liquid-phase composition profile below. Also included on the next page are the column profiles for temperature, liquid flow, and vapor flow, as well as the stage compositions at the bottom of the column, which show rapid changes, making the high bottoms purity easy to attain. Column 1 was sized with Chemcad for valve trays, with a 30-inch tray spacing. The resulting column inside diameter was determined to be 24.5 feet, which is large of the high reflux ratio and vacuum conditions. The percent of flooding varied from 86% at the top to 38% at the bottom. The condenser duty was computed to be 48.11 MW, with a reboiler duty of 47.42 MW. For the 40 theoretical plates, the total pressure drop was computed to be 32 kPa. For a final design, the computations should be repeated taking into account tray pressure drop and tray efficiency. Exercise 11.15 (continued) Analysis: Column 1 (continued) Exercise 11.15 (continued) Analysis: Column 1 (continued) Exercise 11.15 (continued) Analysis: Column 1 (continued) Analysis: (continued) Column 2: Using the TPXY plot feature of Chemcad, it was found that the relative volatility, , for water with respect to ethanol, at 300 kPa, is close to one, but does not vary widely over the composition range, from 1.027 at 88 mol% ethanol at the distillate to 1.149 at 99.8 mol% ethanol at the bottoms, both mole fractions in the liquid phase. Accordingly, it was believed that the FUG method, the ShortCut Column of Chemcad, could give good estimates. Using the above material balance table, the specifications were 5.85/5.91 = 0.9898 for the recovery of water to the distillate, and 42.91/72.91 = 0.5885 for the recovery to the distillate of ethanol. The feed to Column 2 is the distillate from Column 1 after being pumped to a pressure of 350 kPa for entry into Column 2 and heating to the bubble point at that pressure. The results are: Minimum number of equilibrium stages = 47.2 Minimum reflux ratio = 5.85 Operating reflux ratio = 1.3(5.85) = 7.6 Number of equilibrium stages from Gilliland correlation = 86.9 Feed stage = 9 Using some small modifications to these values, the input data to the Tower model of Chemcad was as follows: Top pressure = 300 kPa Condenser P = 5 kPa Column P = 70 kPa No. of stages = 88 (includes total condenser and partial reboiler) Feed stage = 10 from the top Reflux ratio = 7.6 Bottoms molar flow rate = 30.06 Estimated distillate rate = 48.76 mol/s (specified value) Estimated reflux rate = 15D = 7.6(48.76) = 370.6 mol/s Estimated temperatures: Stage 1, 10835oC Stage 12, 116oC Stage 2, 110oC With these input data, a converged solution was obtained, but an ethanol bottoms product of only 98.8 mol% purity was obtained compared to the desired 99.8 mol%. Also a pinch region of about 10 stages was evident just below the feed. Therefore, a second run was made with the following changes: Feed stage = 20 from the top (moved down 10 stages) Distillate flow rate = 78.82 mol/s, in place of reflux ratio Mole fraction of ethanol in bottoms = 0.998, in place of bottoms flow rate Convergence was achieved in 7 iterations, with a reflux ratio = 7.34. The material balance was in agreement with the above material balance table for the system. Although the feed location could probably be improved, no pinch region was now evident, as shown in the liquid-phase composition profile on the next page. Column 2 was sized with Chemcad for valve trays, with a 24-inch tray spacing. The resulting column inside diameter was determined to be 7.5 feet for a nominal 85% of flooding. The condenser duty was computed to be 14.88 MW, with a reboiler duty of 14.86 MW. For the 86 theoretical plates, the total pressure drop was computed to be 74.6 kPa. For a final design, the computations should be repeated taking into account tray pressure drop and tray efficiency Exercise 11.15 (continued) Exercise 11.15 (continued) Analysis: Column 2 (continued) Exercise 11.16 Subject: Feasibility of the separation of benzene (B) from cyclohexane (C) using homogeneous azeotropic distillation with methanol (M) as the entrainer. Given: Fresh feed of 100 kmol/h of 25 mol% B and 75 mol% C. Reference to an article by Ratliff and Strobel. Assumptions: Feed is at the bubble point at 1 atm. Homogeneous azeotropic distillation at 1 atm using M. Find: Whether M can make possible the separation of B from C. Analysis: The normal boiling points of the three components are: Component Normal b. pt., oC B 80.24 C 80.64 M 64.48 From the "Handbook of Chemistry and Physics" and use of the Wilson equation for activity coefficients with the Chemcad program, the following experimental and predicted binary azeotropes, respectively, are found, with reasonably good agreement. Handbook, expt. Wilson, predic. Binary mixture: T, oC Mol% T, oC Mol% B 77.56 53.7 77.62 54.7 C 46.3 45.3 B M C M 57.5 53.9 39.0 61.0 40.0 60.0 58.59 54.97 40.9 59.1 39.0 61.0 No ternary azeotrope is known. In the reference article by Ratliff and Strobel of Continental Oil, they describe the use of homogeneous azeotropic distillation with methanol to purify a benzene-rich stream. The feed to the tower, including the methanol entrainer, is in vol%, 9.1 M, 86.3 B, 2.4 paraffins-napthenes (including C), and 2.2 olefins. The amount of C in the feed is not known. From the azeotropic tower, a bottoms product of 99.0 vol% benzene is obtained, together with 0.3 vol% paraffinsnaphthenes and 0.7 vol% olefins. The distillate in vol% is 36.0 M, 48.6 B, 8.7 paraffinsnaphthenes, and 6.7 olefins. Thus, compared to the design with acetone in Example 11.6, Ratliff and Strobel report the use of only a small amount of M and a recovery of benzene of only 85%. Analysis: (continued) A residue curve map, produced by the Aspen Plus program with the Wilson equation, for the B-C-M system at 1 atm is shown on the next page, where the three binary azeotropes are marked with dots. Note that Aspen Plus predicts the B-M azeotrope at about 39 mol% B, which is in agreement with experiment. There are two distillation boundaries, one from the C-M azeotrope to the B-M azeotrope with little curvature, and one from the C-M azeotrope to the B-C azeotrope, with much curvature. These two boundaries divide the diagram into three distillation regions. Also included on the diagram are two straight lines, one from the B-C feed of 75 mol% C to pure M. The composition of the combined feed and entrainer lies somewhere on this line. Assume that ideally the bottoms product is pure benzene and that the distillate is the C-M azeotrope. Another straight line connects these two product points. The intersection of the two lines is the mixing point. By the inverse lever-arm rule or by material balance, the following preliminary, ideal material balance is obtained: kmol/h: Component Feed Entrainer Distillate Bottoms B 25 0 0 25 C 75 0 75 0 M 0 192 117 0 Total: 100 192 192 25 Even though the distillate and bottoms compositions appear to lie in a region different from the combined feed and entrainer, the separation is considered possible because of the strong curvature of the boundary from the C-M azeotrope to the B-M azeotrope, as discussed on pages 602-604. However, many stages may be necessary to achieve the products. As a first try, assume a reflux ratio of 5 and 100 equilibrium stages. Mix the feed and entrainer together and send the combined feed, as a bubble-point liquid, to stage 40 from the top. Specify a bottoms flow rate of 25 kmol/h. The calculations were made with the Chemsep program, using the continuation option because of convergence difficulties due to the non-ideality of the system. A good separation was not achieved since the bottoms was only 79.2 mol% B. Several of the parameters, including total stages, methanol entry stage, feed entry stage, reflux ratio, methanol flow rate, reflux ratio, and bottoms flow rate. The best result was obtained with the following specifications: Number of equilibrium stages = 100 + total condenser = 101 Methanol flow rate = 115 kmol/h sent to stage 56 from the top, counting the total condenser. Feed sent to stage 71 from the top, counting the total condenser. Reflux ratio = 10 Bottoms flow rate = 24.0 kmol/h Products were as follows, with 99.58 mol% benzene bottoms, but only a 95.6% recovery of benzene: Component Feed Entrainer Distillate Bottoms B 25 0 1.1 23.9 C 75 0 75.0 0.0 M 0 115 114.9 0.1 Total: 100 115 191.0 24.0 Exercise 11.16 (continued) Analysis: (continued) Exercise 11.16 (continued) Exercise 11.17 Subject: Separation of toluene (T) from 2,5-dimethylhexane (H) by a sequence that includes homogeneous azeotropic distillation with methanol (M) as the entrainer. Given: Feed of 100 mol/s of 50 mol% T and 50 mol% H. Reference to an article by Benedict and Rubin. Find: A separation sequence that includes homogeneous azeotropic distillation with M as the entrainer. A feasible design for the azeotropic distillation column. Analysis: The normal boiling points of the three components are: Component Normal b. pt., oC T H 110.64 109.15 M 64.48 From use of the UNIFAC equation for activity coefficients with the Chemcad program, the following predicted minimum boiling binary azeotropes are found for a pressure of 1 atm. Binary mixture: T H T M UNIFAC, prediction T, oC Mol% 106.9 63.8 47.5 52.5 11.2 88.8 H 60.3 20.7 M 79.3 No ternary azeotrope is known. A residue curve map, produced by the Aspen Plus program with the UNIFAC equation, for the T-H-M system at 1 atm is shown on the next page, where the three binary azeotropes are marked with dots. There are two distillation boundaries, one from the T-M azeotrope to the H-M azeotrope with little curvature, and one from the T-H azeotrope to the H-M azeotrope, with much curvature. These two boundaries divide the diagram into three distillation regions. Also included on the diagram are two straight lines, one from the T-M feed of 50 mol% H to pure M. The composition of the combined feed and entrainer must lie somewhere on this line. If we take advantage of the high degree of curvature of the second distillation boundary, we can assume that ideally the bottoms product can be pure toluene, with a distillate of the H-M azeotrope. Another straight line connects these two product points. The intersection of the two lines is the mixing Exercise 11.17 (continued) Analysis: (continued) Exercise 11.17 (continued) Analysis: (continued) point. By the inverse lever-arm rule or by material balance, the following preliminary, ideal material balance is obtained: mol/s: Component Feed Entrainer Distillate Bottoms T 50 0 0 50 H 50 0 50 0 M 0 192 192 0 Total: 100 192 242 50 An initial run is made with the Chemcad SCDS model using the following specifications: Combined feed and entrainer at 1 atm and the bubble point. Column P = 0 Top pressure = 1 atm Condenser P = 0 No. of stages = 31 (includes total condenser and partial reboiler) Combined feed stage = 16 from the condenser Reflux ratio =3 Bottoms flow rate = 50 mol/s (from above material balance) Estimated (actual) distillate rate = 192 mol/s Estimated (actual) reflux rate = 576 mol/s Estimated temperatures: Stage 1, 60oC Stage 31, 60oC Stage 2, 60oC Stage 30, 60oC This column was difficult to converge, but with the above specifications, plus a 0.50 damping factor, converge was achieved in 25 iterations. The calculated material balance was as follows: mol/s: Component Feed Entrainer Distillate Bottoms T 50 0 0.3 49.7 H 50 0 49.8 0.2 M 0 192 191.9 0.1 Total: 100 192 242.0 50.0 The liquid phase composition profile is on the following page. The condenser and reboiler duties are 12.37 x 104 MJ/h and 12.50 MJ/h respectively. The column profiles are given on a following page. The azeotropic distillation column was sized with Chemcad for sieve trays, with: 24-inch tray spacing 12% downcomer area 10% hole area 1/4-inch holes 80% of flooding The resulting column inside diameter was determined to be 16 feet For a final design, the computations should be repeated taking into account tray pressure drop and tray efficiency. Exercise 11.17 (continued) Analysis: (continued) A feasible separation sequence would include the azeotropic distillation column, considered above, to separate the feed into 99.5 mol% toluene and a near azeotropic mixture of methanol and 2,5 dimethylhexane. The M-H azeotrope would be separated by liquid-liquid extraction with water to give nearly pur H and a methanol-water mixture. The latter would be separated by ordinary distillation to recover the methanol for recycle to the azeotropic tower, and water for recycle to the extraction section. Exercise 11.17 (continued) Analysis: (continued) Exercise 11.18 Subject: Feasibility of the separation of a mixture of methyl acetate (A) and methanol (M) by a sequence that includes homogeneous azeotropic distillation using as the entrainer either n-hexane (H), cyclohexane (C), or toluene (T). Given: Feed rate of 16,500 kg/h of 55 wt% A and 45 wt% M. Desired products are 99.5 wt% A and 99 wt% M. H, C, and T as potential entrainers. Find: Feasibility of a separation sequence that includes homogeneous azeotropic distillation with one the three potential entrainers. If not feasible, find an alternative feasible process. Analysis: The molecular weights and normal boiling points of the five components of interest are: Component Molecular Normal boiling point, o Weight C Methyl acetate, A 74.08 57.1 Methanol, M 32.04 64.7 n-Hexane, H 86.17 69.0 Cyclohexane, C 84.16 80.7 Toluene, T 92.13 110.8 Convert the feed specification from a weight to a mole basis. Component Feed, kg/h Feed, kmol/h Feed, mol% Methyl acetate, A 9,075 122.5 34.58 Methanol, M 7,425 231.7 65.42 Total: 16,500 354.2 100.00 From the Handbook of Chemistry and Physics, the homogeneous binary azeotropes are: Binary pair Normal boiling pt., Mol% o C A-M 54.0 65.3, 34.7 A-H 51.8 64.2, 35.8 A-C 55.5 80.1, 19.9 M-T 63.7 88.3, 11.7 Also, heterogeneous binary azeotropes are formed by the pairs M-H and M-C. No azeotrope is formed with the pair A-T. The following two homogeneous ternary azeotropes are formed. No ternary azeotrope is formed by A-M-T. Ternary system Normal boiling pt., Mol% o C A-M-H 45.0 24.5, 29.4, 46.1 A-M-C 50.8 40.7, 34.5, 24.8 Exercise 11.18 (continued) Analysis: (continued) The calculations can be made with the ChemCad program, using the SCDS model. The results will depend on the method used to compute K-values. Because neither the NRTL nor the UNIQUAC methods in ChemCad include parameters for the methanol-normal hexane pair, use the UNIFAC method. Consider the use of n-hexane as the entrainer. A two-column method used successfully by Laroche, Bekiaris, Andersen, and Morari in Ind. Eng. Chem. Res., 31, 21902209 (1992), sends to the first column the feed and a recycle of near ternary azeotrope from the second column. The bottoms is nearly pure methanol. The distillate is fed to the second column, where the bottoms is nearly pure methyl acetate and the distillate is the near ternary azeotrope, which is recycled to the first column. Care is needed in using n-hexane because as noted above, it forms a two-liquid-phase system over part of the composition space as shown in the triangular diagram on the following page. Operation of both columns must be outside the two-liquid-phase region. Included on the diagram are the composition points for the azeotropes and approximate distillation boundary curves. For a first approximation, assume that the recycle distillate from the second column is 1500 kmol/h of 29 mol% A, 34.5 mol% M, and 36.5 mol% H. With a reflux ratio of 0.5 and 22 stages (including total condenser and partial reboiler) in the first column, with the feed and recycle distillate sent to stage 11 from the condenser, a very pure M is obtained. A 99.97 mol% M bottoms is produced with 14 stages, with stage 7 as the feed stage. With a reflux ratio of 4 and 42 stages (including total condenser and partial reboiler) in the second column, with the feed (distillate from the second column) sent to stage 21 from the condenser, a bottoms of 98.6 mol% A is obtained. This could not be improved by increasing the reflux, increasing the number of stages, or shifting the feed stage. The liquid composition profiles are shown on subsequent pages, where it is seen that no pinch zones are evident. The material balance is as follows: Component Methanol, M Methyl acetate, A n-Hexane, H Total kmol/h: Feed F1 231.7 122.5 0.0 354.2 Entrainer, D2 516.4 436.6 547.0 1500.0 B1 230.0 0.1 0.0 230.1 D1 = F2 518.1 559.0 547.0 1624.1 B2 1.7 122.4 0.0 124.1 The material balance lines are included on the triangular diagram on the next page. Exercise 11.18 (continued) Analysis: (continued) Exercise 11.18 (continued) Analysis: (continued) Exercise 11.19 Subject: Separation of a mixture of ethanol and water into nearly pure ethanol and pure water, using a three-column sequence that includes heterogeneous azeotropic distillation with benzene. Given: 150 mol/s of the ethanol-water azeotrope at 1 atm. Benzene as the entrainer for heterogeneous azeotropic distillation. Find: A suitable three-column design. However, because the feed is at near-azeotropic composition, the first (Preconcentrator) column is not needed. Analysis: The molecular weights and normal boiling points of the three components of interest are: Component Molecular Normal boiling Weight Point, oC Water, W 18.02 100.0 Benzene, B 78.11 80.1 Ethyl alcohol, E 46.07 78.5 From the Handbook of Chemistry and Physics, the binary azeotropes are: Binary pair Normal Type Mol% Mol% Mol% lower boiling Azeotrope overall upper liquid phase pt., oC liquid phase W-B 69.4 Heterogeneous 29.8% W 0.26% W 99.98% W 70.2% B 99.74% B 0.02% B W-E 78.2 Homogeneous 10.5% W 89.5% E B-E 67.8 Homogeneous 55.2% B 44.8% E Also, the following heterogeneous ternary azeotrope is formed: Ternary Normal Type Mol% Mol% Mol% lower System boiling Azeotrope overall upper liquid phase o pt., C liquid phase W-B-E 64.6 Heterogeneous 23.3% W 5.0% W 66.7% W 53.9% B 76.0% B 1.7% B 22.8% E 19.0% E 31.6% E The column calculations were made with the ChemCad program. Important decisions were the choice of activity coefficient method for the liquid and the binary interaction parameters for the selected method because these decisions can have a significant influence on the results of the calculations. The UNIQUAC method was selected, with values of the binary interaction parameters taken from the Appendix of the article by Bekiaris, Meski, and Morari, IEC Res., 35, 207-227 (1996). They give the parameters as values of (uij - ujj)/R in K. For the ChemCad program, the parameters must be given as (uij - ujj) in cal/mol. Therefore, the Bekiaris, Meski, Exercise 11.19 (continued) Analysis: (continued) and Morari values were multiplied by 1.987 cal/mol-K, giving the following values of (uij - ujj) for input to ChemCad: i j Water Benzene Ethanol Water ---719.890 405.036 Benzene 1795.85 ---764.780 Ethanol -65.995 -85.507 ---With these values, the computed azeotropes at 1 atm, which deviate somewhat from the above experimental values, were as follows, where the homogeneous azeotropes were determined with the TPYX plot feature and the heterogeneous azeotropes were determined by running three-phase flash calculations with the LLVF model until the vapor and overall liquid compositions were equal. The most serious deviation is the lower liquid phase composition of the ternary azeotrope. From the built-in vapor pressure correlations in ChemCad, the normal boiling points of water and benzene are the same as the above literature values, but the ethanol value is 78.3oC. Binary pair W-B W-E B-E Normal boiling pt., oC 69.2 78.2 67.8 Type Azeotrope Heterogeneous Homogeneous Homogeneous Mol% overall 29.8% W 70.2% B 9.7% W 90.3% E 55.0% B 45.0% E Mol% overall 19.7% W 53.1% B 27.2% E Mol% upper liquid phase 0.47% W 99.53% B Mol% lower liquid phase 99.94% W 0.06% B Ternary System W-B-E Normal boiling pt., oC 63.8 Type Azeotrope Heterogeneous Mol% upper liquid phase 2.4% W 81.7% B 15.9% E Mol% lower liquid phase 47.7% W 7.0% B 45.3% E The above values also differ somewhat from the values cited in Section 11.6. A binodal plot for the liquid-liquid equilibrium of the ternary system, as computed with ChemCad at a temperature of 64oC, is given on the next page. Many liquid-liquid phase equilibrium tie lines are shown. Added to the diagram are the above computed temperatures and compositions for the azeotropes, and estimates of the distillation boundary lines taken from the above-cited Bekiaris, Meski, and Morari article Exercise 11.19 (continued) Analysis: (continued) The plot on the previous page differs somewhat from the residue curve map in Figure 11.29, but the same three distillation regions are evident. Therefore, as discussed in Section 11.6, it appears possible, using the three column scheme of Fig. 11.31 (without the first column because the feed is already close to the azeotropic composition of the ethanol-water mixture), to obtain nearly pure ethanol from the bottom of the second column and nearly pure water from the bottom of the third column. Runs with ChemCad were made by sending the feed at the bubble point temperature directly to the azeotropic column in Fig. 1131a. Unfortunately, repeated attempts to achieve nearly pure ethanol, using different combinations of entrainer flow rates and compositions were unsuccessful. The best ethanol purity achieved was 95 mol%. For that case, 14 equilibrium stages were used in the azeotropic column plus a decanter-condenser combination and a partial reboiler, with the feed to stage 4 from the top. The entrainer recovery column had a reflux ratio of two and 15 stages, including a total condenser and a partial reboiler, with the feed to stage 5 from the condenser. The distillate from the third column and the benzene-rich liquid phase from the decanter were recycled to the top of the azeotropic column. Convergence of both the column calculations and the recycle streams was difficult. A reasonably converged material balance, keyed to the three-column diagram of Fig. 11.31 is as follows: Stream in Fig. 11.31a Component Water Benzene Ethanol Total: Mol/sec: D1 15 0 135 150 D3 11.4 20.0 46.2 77.6 L2 3.8 88.8 25.7 118.3 V2 23.3 108.8 71.9 204.0 B2 6.9 0.2 135.0 142.1 D2 19.5 20.0 46.3 85.8 B3 8.1 0.0 0.1 8.2 Exercise 11.20 Subject: Separation of a mixture of isopropanol and water into nearly pure isopropanol and pure water, using a three-column sequence that includes heterogeneous azeotropic distillation with benzene. Given: 120 mol/s of the isopropanol-water azeotrope at 1 atm. Benzene as the entrainer for the heterogeneous azeotropic distillation step. Find: A suitable three-column design. The molecular weights and normal boiling points of the three components of interest are: Component Molecular Normal boiling Weight point, oC Water, W 18.02 100.0 Benzene, B 78.11 80.1 Isopropyl alcohol, P 60.09 82.3 From the Handbook of Chemistry and Physics, the experimental binary azeotropes are: Binary pair Normal Type Mol% Mol% upper Mol% lower boiling Azeotrope overall Liquid phase liquid phase o pt., C W-B 69.4 Heterogeneous 29.8% W 0.26% W 99.98% W 70.2% B 99.74% B 0.02% B W-P 80.4 Homogeneous 31.7% W 68.3% P B-P 71.5 Homogeneous 60.6% B 39.4% P Also, the following heterogeneous ternary azeotrope is formed: Ternary Normal Type Mol% Mol% Mol% lower System boiling Azeotrope overall upper liquid phase o pt., C liquid phase W-B-P 65.7 Heterogeneous 26.7% W 8.8% W 95.1% W 54.0% B 68.5% B 0.1% B 19.3% P 22.7% P 4.8% P The column calculations were made with the ChemCad program. Important decisions were the choice of activity coefficient method for the liquid and the binary interaction parameters for the selected method, because these decisions can have a significant influence on the results of the calculations. The UNIQUAC method was selected, with built-in values of the binary interaction parameters for pairs with isopropanol; but for the benzene-water pair, values were taken from the Appendix of the article by Bekiaris, Meski, and Morari, IEC Res., 35, 207-227 (1996). They give the parameters as values of (uij - ujj)/R in K. For the ChemCad program, the parameters Exercise 11.20 (continued) Analysis: (continued) must be given as (uij - ujj) in cal/mol. Therefore, the Bekiaris, Meski, and Morari values were multiplied by 1.987 cal/mol-K. In summary, the following values of (uij - ujj) were used for the ChemCad runs: i j Water Benzene Isopropanol Water ---719.890 109.550 Benzene 1795.85 ---319.618 Isopropanol 300.190 54.883 ---With these values, the computed azeotropes at 1 atm, which deviate somewhat from the above experimental values, were as follows, where the homogeneous azeotropes were determined with the TPYX plot feature of ChemCad and the heterogeneous azeotropes were determined by running three-phase flash calculations with the LLVF model of ChemCad until the assumed feed composition gave vapor and overall liquid compositions that were equal. The most serious deviations are for the ternary heterogeneous azeotrope. From the built-in vapor pressure correlations in ChemCad, the normal boiling points of water and benzene are the same as the above literature values, but the isopropanol value is 82.5oC. Binary pair W-B W-P B-P Normal boiling pt., oC 69.2 80.2 71.6 Type Azeotrope Heterogeneous Homogeneous Homogeneous Mol% overall 29.8% W 70.2% B 32.6% W 67.4% P 58.5% B 41.5% P Mol% overall 23.8% W 52.3% B 23.9% P Mol% upper liquid phase 0.47% W 99.53% B Mol% lower liquid phase 99.94% W 0.06% B Ternary System W-B-P Normal boiling pt., oC 60.6 Type Azeotrope Heterogeneous Mol% upper liquid phase 4.6% W 67.6% B 27.8% P Mol% lower liquid phase 89.3% W 0.3% B 10.4% P A binodal plot for the liquid-liquid equilibrium of the ternary system, as computed with ChemCad at a temperature of 66oC, is given on the next page. Many liquid-liquid phase equilibrium tie lines are shown. Added to the diagram are the above computed temperatures and compositions for the azeotropes, and estimates of the distillation boundary lines, which divide the composition plot into three distillation regions. Analysis: (continued) Exercise 11.20 (continued) As discussed in Section 11.6, it appears possible, using the three column scheme of Fig. 11.31a (without the first column because the feed is already at or near the azeotropic composition of the isopropanol-water mixture), to obtain nearly pure isopropanol from the bottom of the second column and nearly pure water from the bottom of the third column. Runs with ChemCad were made with a feed of 120 mol/s of near-azeotropic composition of 67 mol% P and 33 mol% W at the bubble point temperature, sent directly to the azeotropic column in Fig. 11.31a. An estimate of the total entrainer flow rate and composition (streams L2 and D3 combined) was made by locating a composition point near the distillation boundary between Regions 1 and 2 and near the binodal curve separating the two-liquid-phase region from the one-liquid-phase region. The composition was taken as 8.5 mol% W, 56 mol% B, and 35.5 mol% P. To determine an initial approximation of the total flow rate of the entrainer, the overhead vapor from the azeotropic column was assumed to be near the ternary azeotropic composition. By material balance, the entrainer flow rate was computed to be 240 mol/s. The azeotropic column (column 2 in Fig. 11.31a) was then computed with the SCDS model in ChemCad, using the following specifications, with the pressure = 1 atm everywhere: Total number of equilibrium stages (including partial reboiler) = 20 Feed stage = stage 5 from the top No condenser Bottoms flow rate = 80.4 mol/s (the flow rate of P in the feed) Assumed distillate rate = 239.6 mol/s Assumed reflux rate = 200 mol/s Assumed stage temperatures = 61, 82, 62, and 78OC The converged results showed a large pinch region where little or no separation occurred, and the bottoms purity with respect to isopropanol was not acceptable. Accordingly, a number of additional runs were made with varying conditions for the total number of stages, feed stage location, and total entrainer flow rate. A bottoms product of isopropanol purity of better than 99.5 mol/s at a flow rate of 80.4 mol/s was achieved with 20 equilibrium stages, feed to the top stage, and a total entrainer flow rate of 300 mol/s. The entire separation process was then simulated with Chemcad, according to the flowsheet in Fig. 11.31a, but without the first column (Preconcentration). The flowsheet included the azeotropic column (column 2) with recycle entrainer from the entrainer recovery column (column 3), and a condenser-decanter to condense the overhead vapor from column 2 and split the condensate into two liquid phases, with the organic-rich phase sent as entrainer reflux to the top of column 2 and the water-rich phase sent to column 3. To achieve a nearly pure water bottoms product from column 3, only 4 equilibrium stages (including a total condenser and partial reboiler) were needed with the feed to stage 2 (the top stage of the column) and a reflux ratio of 1. The distillate from column 3 was recycled to the top of column 2. Several loops were needed to converge the material balance for the recycle process. Exercise 11.20 (continued) Analysis: (continued) The resulting material balance was as follows: Stream in Fig. 11.31a Component Water Benzene Isopropanol Total: Mol/sec: D1 39.6 0 80.4 120.0 D3 5.8 0.2 6.1 12.1 L2 19.9 169.3 99.5 288.7 V2 65.2 169.5 105.8 340.5 B2 0.1 0.0 80.2 80.3 D2 45.3 0.2 6.3 51.8 B3 39.5 0.0 0.2 39.7 From the above material balance, the isopropanol product purity is 99.9 mol% and the water product purity is 99.6 mol%. The benzene loss is negligible. The material balance lines are shown on the triangular diagram on the following page. The separation in column 1 is in distillation region 1, but the decanter sends the distillate, D2, into region 2 where the separation in column 3 occurs. The composition profiles for columns 2 and 3 are shown in two subsequent pages, where it is seen that no pinch zones are evident. Exercise 11.20 (continued) Analysis: (continued) Exercise 11.20 (continued) Analysis: (continued) Exercise 11.21 Subject: Separation of a mixture of acetic acid (A) and water (W) into nearly pure acetic acid and pure water, using a two-column sequence that includes heterogeneous azeotropic distillation with n-propyl acetate (NPA). Given: 1,000 kmol/h of 20 mol% acetic acid in water. Find: A suitable two-column design. Analysis: The phase behavior of the acetic acid/water/n-propyl acetate system at 1 atm was explored with the UNIQUAC equation for liquid-phase activity coefficients, noting that the normal boiling points in oC are 118.1 for A, 100 for water, and 101.4 for NPA. Using the ChemCad and Aspen Plus programs, the following data were obtained, including a residue curve map for the vapor-liquid system at 1 atm and a ternary liquid-liquid diagram for the system at 80oC, both shown on the next page. Acetic acid and water form a minimum-boiling azeotrope at 90 mol% water that boils at 99.47oC. Acetic acid and n-propyl acetate do not form an azeotrope. Water and n-propyl acetate form a heterogeneous azeotrope at approximately 40 mol% NPA and 80oC. This azeotrope separates into a mixture of almost pure water and 83 mol% NPA. The two diagrams show that a ternary mixture, when distilled will tend to give a bottoms of almost pure acetic acid and an overhead that when condensed will split into two liquid phases, one of near pure water, and the other of concentrated NPA that can be recycled to the column as entrainer. Consider a separation sequence that omits the preconcentator column shown in Fig. 11.31 because the feed is already concentrated in water. The feed is sent directly into the heterogeneous azeotropic distillation column with a decanter to split the two liquid phases resulting from the condensation of the overhead vapor. An initial geometrical construction for determining the separation in that column is included in the ternary diagram on the next page. The feed, F, of 800 kmol/h of W and 200 kmol/h of A enters the column along with an entrainer, R, of saturated NPA-rich phase containing 18 mol% W, 10 mol% A, and 72 mol% NPA in equilibrium with a saturated water-rich phase containing 97.5 mol% W, 2 mol% A, and 0.5 mol% NPA, which becomes the distillate. The amount of entrainer reflux is fixed so as to give the mixing point, M, such that a line through that point and about 99 mol% A in W will intersect a liquid-liquid tie line that connects the reflux and the nearly pure liquid water distillate in the region close to the binary heterogeneous azeotrope of W and NPA. From this construction, the amount of reflux is determined as 1,000 kmol/h. An initial material balance was then computed, from the feed and mole fraction compositions, with the following result: kmol/h: Component Feed Entrainer Make-up NPA Bottoms Overhead vapor Distillate w/o makeup Water 800 180 0 2 978 798 Acetic acid 200 100 0 185 115 15 NPA 0 720 4 0 720 4 Total: 1000 1000 4 187 1813 817 Analysis: (continued) Exercise 11.21 (continued) Exercise 11.21 (continued) Analysis: (continued) The SCDS distillation model of Chemcad was used with no condenser and a partial reboiler. The overhead vapor was sent a three-phase flash condenser (decanter) operating at the bubble point, modeled with the LLVF unit. The NPA-rich phase would be recycled to the column as entrainer reflux entering the top stage, but in the calculation the guessed amount of 1,000 kmol/h at the above composition was used and then converged by trial and error, accounting for the necessary makeup NPA. An initial guess was made of 20 stages, including the reboiler, with the feed entering stage 6 from the top. Other specifications were 1 atm for the column, feed and entrainer at the bubble-point temperature at 1 atm, 187 kmol/h of bottoms, estimates of 1800 kmol/h for the distillate, 1000 kmol/h for the reflux; 80, 110, 90, and 100 oC for the four temperatures; and a damping factor of 0.3. The calculations converged without difficulty in 14 iterations, but the entrainer flows minus the makeup NPA did not equal the NPArich liquid from the decanter. Also 10 of stages appeared to be pinched. Therefore, the number of stages was reduced to 10, the feed stage was moved to stage 4 from the top, and the component flow rates in the entrainer were revised by trial and error until a reasonable closure of the material balance was obtained. The final material balance was as follows: kmol/h: Component Feed Entrainer Make-up NPA Bottoms Overhead vapor Distillate w/o makeup Water 800 184 0 1.3 982.7 798.4 Acetic acid 200 58 0 183.6 74.4 16.0 NPA 0 760 3.6 0.1 763.5 3.5 Total: 1000 1002 3.6 185.0 1820.6 817.9 Temperatures in the column varied from 85.6oC at the top to 117.5o C at the bottom. The temperature in the decanter was 83.9oC. The following is a plot of the liquid-phase composition profile showing that no pinch existed in the column. The acetic acid product is 99.2 mol% acetic acid, while the bottoms water product is 97.6 mol% water. This could be further purified in a second column, with the small amount of acetic acid-rich product recycled. Exercise 11.21 (continued) Analysis: (continued) Exercise 11.22 Subject: Reactive distillation of isobutene (IB) with methanol (M), in the presence of n-butene (NB) to produce MTBE. Effect of the feed stage location for methanol. Given: Mixed butenes feed of 195.44 mol/s of IB and 353.56 mol/s of NB entering stage 11 as a vapor at 350 K and 11 bar. Methanol feed of 215.5 mol/s at 320 K and 11 bar. Column operates at 11 atm and has 15 equilibrium stages, with a total condenser and partial reboiler. Reflux ratio = 7 and bottoms flow rate is set at 197 mol/s. Catalyst for the reaction is provided in the amount of 204.1 kg/stage only at stages 4 through 11. Kinetic data are given in Example 11.9. Use UNIQUAC method for liquid-phase activity coefficients, with the binary interaction parameters in Example 11.9, and RK EOS for vapor-phase fugacities. Find: Distillate and bottoms products as a function of feed stage location for methanol. Analysis: Chemcad can be used, but here, the calculations were made with Aspen Plus version 10.1, but convergence was not easy to achieve. The initial case was run with the methanol fed to stage 7 from the top, counting the total condenser as the first stage. The input data, which are given on the next two pages, differ slightly from the listing for Example 11.9. With this feed stage, the overall material balance in mol/s are: Component Feed 1 Feed 2 Distillate Bottoms Methanol 215.50 0.00 31.26 0.13 Isobutene 0.00 195.44 8.25 3.08 n-butene 0.00 353.56 343.69 9.87 MTBE 0.00 0.00 0.19 183.92 Total: 215.50 549.00 383.39 197.00 To obtain the results for the other methanol feed-stage locations, the following procedure was used. Starting from Stage 7 and using the previous results to start each new run, results were obtained for methanol feed stages of 8, 9, 10, and 11. Working backward from Stage 7, results were obtained for methanol feed stages of 6 and 5. Then, results were obtained, starting from Stage 2, for methanol feed stages of 2, 3, and 4, where the input on the next two pages were modified under BLOCK B1 RADFRAC to: X-EST 17 MEOH 0.01 / 17 MTBE 0.95 Y-EST 17 MEOH 0.02 / 17 MTBE 0.875 To obtain results for methanol feed stages 12 to 16, the input data on the next two pages were first modified under BLOCK B1 RADFRAC to: X-EST 17 MEOH 0.90 / 17 MTBE 0.00 Y-EST 17 MEOH 0.90 / 17 MTBE 0.00 This gave results for methanol feed stages of 12, 13, 14, 15, and 16; and 12 back to 11. The values for % conversion of isobutene were as follows, and compare well with Fig. 11.39. Stage 2 3 4 5 6 7 8 9 10 11 11 12 13 14 15 16 %Conv 89 91 92 93 94 94 95 95 96 96 0.7 0.7 0.6 0.6 0.6 0.6 Analysis: (continued) Exercise 11.22 (continued) Analysis: (continued) Exercise 11.22 (continued) Exercise 11.23 Subject: Reactive distillation of isobutene (IB) with methanol (M), in the presence of n-butene (NB) to produce MTBE. Use of activities instead of mole fractions in the kinetic equations. Given: Mixed butenes feed of 195.44 mol/s of IB and 353.56 mol/s of NB entering stage 11 as a vapor at 350 K and 11 bar. Methanol feed of 215.5 mol/s at 320 K and 11 bar to stage 10. Column operates at 11 atm and has 15 equilibrium stages, with a total condenser and partial reboiler. Reflux ratio = 7 and bottoms flow rate is set at 197 mol/s. Catalyst for the reaction is provided in the amount of 204.1 kg/stage only at stages 4 through 11. Kinetic data is given in Example 11.9, but substitute activities for mole fractions in the kinetic equations. Use UNIQUAC method for liquid-phase activity coefficients, with the binary interaction parameters in Example 11.9, and RK EOS for vapor-phase fugacities. Find: Distillate and bottoms products, and comparison of results with those of Example 11.9. Analysis: Kinetic equations for the reaction rates are most often written in terms of component concentrations, component mole fractions, or component partial pressures (in the case of gasphase reactions). However for liquid-phase reactions when the components form a non-ideal solution, chemical equilibrium can only be satisfied when the rate equation is written in terms of activities (product of mole fraction times liquid-phase activity coefficient). For the formation of MTBE from isobutene and methanol, the Rehfinger and Hoffmann article was developed in terms of component activities. However, in Example 11.9, component mole fractions are used, as seen in Eqs. (1) and (2). This assumes that all component activity coefficients are equal to one. With Aspen Plus, built-in kinetic rate-law expressions are not available in terms of activities. Therefore, a FORTRAN subroutine for the rate-law must be provided. On the following four pages, the Aspen Plus input data are listed for solving this exercise. Under the Aspen Plus paragraph REACTIONS R-MTBE-U REAC-DIST, reference is made to the usersupplied subroutine RAMTBE. This subroutine is listed on the subsequent five pages, where the kinetic rate equation is written in terms of activities. The Aspen Plus input data are written to explore the variation of methanol feed stage. For methanol fed to stage 10, the conversion of isobutene, when using activities, is 99.5% compared to a 95.6% conversion in Example 11.9, where mole fractions are used. Analysis: (continued) Exercise 11.23 (continued) Analysis: (continued) Exercise 11.23 (continued) Analysis: (continued) Exercise 11.23 (continued) Analysis: (continued) Exercise 11.23 (continued) Exercise 11.23 (continued) Analysis: (continued) Analysis: (continued) Exercise 11.23 (continued) Exercise 11.23 (continued) Analysis: (continued) Analysis: (continued) Exercise 11.23 (continued) Analysis: (continued) Exercise 11.23 (continued) Exercise 11.24 Subject: Reactive distillation of isobutene (IB) with methanol (M), in the presence of n-butene (NB) to produce MTBE. Chemical equilibrium at each stage that has catalyst. Given: Mixed butenes feed of 195.44 mol/s of IB and 353.56 mol/s of NB entering stage 11 as a vapor at 350 K and 11 bar. Methanol feed of 215.5 mol/s at 320 K and 11 bar to stage 10. Column operates at 11 atm and has 15 equilibrium stages, with a total condenser and partial reboiler. Reflux ratio = 7 and bottoms flow rate is set at 197 mol/s. Chemical equilibrium is achieved at stages 4 through 11. RK EOS for vapor-phase fugacities. Assumptions: Use UNIFAC method for liquid-phase activity coefficients. Find: Distillate and bottoms products, and comparison of results with those of Example 11.9. Analysis: The ChemCad program, with the SCDS model and reactive distillation option, was used to make the equilibrium stage calculations. To do this, it is assumed that the required chemical equilibrium constant can be computed by taking the ratio of the forward to the backward rate constant. From Eqs. (1) and (2) in Example 11.9, K = 1.375 x 10-5 exp(42014/(RT)) (1) From Fig. 11.38(a), all of the reaction takes place at a temperature of about 350 K. From Eq. (1), at T = 350 K, with R = 8.314 J/mol-K, K = 25.6. In ChemCad, let lnK = A. Therefore, A = 3.24. Input data for ChemCad, included: No. of stages = 17, with M fed to stage 10, and the butenes fed to stage 11. Estimated temperatures in K were 350, 340, 400, and 420 for stages 1, 2, 16, and 17. A damping factor of 0.25 was used with a maximum of 50 iterations. A single reaction was specified for the equilibrium option and stages 1-3 and 12-17 were disabled. The problem was very slow to converge, taking almost the maximum number of iterations. The overall material balance is as follows: Mol/s: Component Feed 1 Feed 2 Distillate Bottoms Methanol 215.50 0.00 36.28 0.12 Isobutene 0.00 195.44 13.58 2.76 n-butene 0.00 353.56 338.39 15.17 MTBE 0.00 0.00 0.15 178.95 Total: 215.50 549.00 388.40 197.00 The % conversion of isobutene to MTBE = 91.64%. This compares to 95.6% in Example 11.9. The profiles are shown on the next three pages and compare favorably to Figs. 11.38(a to c). Exercise 11.24 (continued) Analysis: (continued) Exercise 11.24 (continued) Analysis: (continued) Exercise 11.24 (continued) Analysis: (continued) Exercise 11.25 Subject: Supercritical extraction of 1 mol/s of 10 wt% ethanol in water by 3 mol/s of carbon dioxide at 305 K and 9.86 MPa in a staged contactor. Given: 10 equilibrium stages. K-values from Example 11.10. Find: Flow rates and compositions of the exiting extract and raffinate. Compare results to those of Example 11.10. Analysis: The ChemCad program with the Tower Plus model was used to make the calculations, as in Example 11.10. Isothermal conditions were employed by allowing heat transfer at each stage. The polynomial K-value option was used to supply the constant K-values given in Example 11.10 (0.115 for ethanol, 0.00575 for water, and 34.5 for CO2) The calculations converged in 7 iterations. The overall material balance was as follows: Component Carbon dioxide Ethanol Water Total: Mol/s: Feed 0.0000 0.0417 0.9583 1.0000 Solvent 3.0000 0.0000 0.0000 3.0000 Raffinate 0.0287 0.0277 0.9422 0.9986 Extract 2.9713 0.0140 0.0161 3.0014 These results are almost identical to those for Example 11.10, which used just 5 equilibrium stages. This should not be surprising! Supercritical extraction is analogous to stripping or liquid-liquid extraction. Thus, an approximate calculation can be made with the Kremser method based on stripping factors for ethanol and water, and an absorption factor for carbon dioxide. For ethanol, the stripping factor is S = KV/L = 0.115(3)/1 = 0.345. From the Kremser plot in Fig. 5.9, the fraction extracted is approximately 35%, independent of the number of stages above 3. For water, S = 0.00575(3)/1 = 0.0173, and the Kremser plot gives approximately 1.7% extracted, again independent of the number of stages above 3. For carbon dioxide, an absorption factor is used, where A = L/KV = 1/[34.5(3)] = 0.00966. From the Kremser plot, approximately 1% of the carbon dioxide is transferred to the raffinate. These values are very close to the following values computed from the above material balance: % extraction of ethanol = 0.0140/0.0417 x 100% = 33.6% % extraction of water = 0.0161/0.9583 x 100% = 1.68% % loss of carbon dioxide to raffinate = 0.0287/3.0 x 100% = 0.957% Exercise 11.26 Subject: Development of a mathematical model for the supercritical extraction of a solute from particles of natural material. An example is the extraction of caffeine from coffee beans by supercritical CO2. Given: Rate of extraction is independent of the flow rate of CO2 past the particles, but dependent on the particle size. Find: A mathematical model for the rate of extraction. The parameter in the model that must be determined by experiment. Analysis: In the extraction of a solute from a natural material by a fluid solvent, the following mechanism may be considered. It is discussed in several places in the literature, including an article by Goto, Roy, and Hirose in The Journal of Supercritical Fluids, 9, 128-133 (1996), entitled "Shrinking-Core Leaching Model for Supercritical-Fluid Extraction". The particle of natural material is soaked in the solvent so that the solvent penetrates into the particle, usually causing it to swell. The material may be thought of as a porous or cellular matrix that fixes or traps the solute in the matrix. However, in the presence of the solvent, the solute is released and can diffuse from the interior of the material to the surface of the particle, and from there into the bulk of the solvent. Thus, the solute mass-transfer process can be viewed as taking place in two steps, one (internal mass-transfer resistance) by diffusion in the interior of the material and the other (external mass-transfer resistance) by convection in the exterior solvent surrounding the material. Assuming that the solvent flows past the particles so as to sweep away the solute as it leaves the surface of the particles, it can be assumed that the external mass-transfer resistance is negligible because the rate of extraction is given as independent of the solvent flow rate past the particle. This is so because it is recalled from Chapter 3 that for forced convection past a spherical particle, Eq. (3-170), the mass-transfer coefficient is proportional to the fluid mass velocity or velocity past the particle to the 0.6 exponent and inversely proportional to the particle diameter to the 0.4 exponent. Since there is no effect of velocity, all of the mass transfer resistance resides in the interior of the particle, where a Fickian diffusion process can be assumed. As suggested by the Goto, Roy, and Hirose (cited above), the interior diffusion process can be viewed as one in which the rate of diffusion in the non-extracted interior core of the particle is much slower than in the outer shell of the particle where most of the solute has been extracted. This results in a sharp boundary or interface between the inner core and outer shell. In the inner core, the concentration of solute is at its initial value. With time, the core region shrinks as the sharp boundary progresses toward the center of the particle. This is the basis for the so-called shrinking-core model widely used in modeling leaching operations. The model was first conceived by Yagi and Kunii in 1955 ["Fifth Symposium (International) on Combustion," Reinhold Publishing Corp., NY (1955), pp. 231-244] for application to gas-solid combustion, and extended to liquid-solid leaching by Roman, Benner, and Becker in 1974 [Trans. Soc. Mining Engineering of AIME, 256, 247-256 (1974)]. In its general form, takes into account both Analysis: (continued) internal and external mass-transfer resistances. Let us develop the model for the case of negligible mass-transfer resistance in the liquid external to the particle. Assume that drc /dt, the rate of movement of the interface at the particle radius, rc , is small with respect to the diffusion velocity of solute A, through the outer shell of the particle. This is referred to as the pseudo-steady-state assumption. The importance of this assumption is that it allows us to neglect the accumulation of solute as a function of time in the outer shell layer as that layer increases in thickness, with the result that the model can be formulated as an ordinary differential equation rather than as a partial differential equation. Thus, the rate of diffusion of solute A through the outer shell is given by Fick's second law, (3-74), ignoring the term on the left-hand side and replacing the molecular diffusivity with an effective diffusivity, De, for the solute through the solvent in the complex matrix of natural material: De d 2 dc A (1) r =0 r 2 dr dr where cA is the concentration of solute in the outer shell of the particle and r is the radial distance from the center of the particle. The boundary conditions are: c A = c As = c Ab at r = rs Exercise 11.26 (continued) c A = c A0 at r = rc where the subcripts are s for the particle surface, b for the bulk, 0 for initial condition, and c for the interface between the outer shell and core of the particle. These boundary conditions hold because the mass-transfer resistance in the liquid film or boundary layer is assumed negligible and the concentration of solute in the core remains at its initial value as the core shrinks. If Eq. (1) is integrated twice and the boundary conditions are applied, the result after simplification is: rc r c A = c A0 - c A0 - c Ab (2) rc 1- rs To obtain a relationship between the location of the interface rc and time t, differentiate Eq. (2) with respect to r and evaluate the differential at r = rc. This gives: - c A0 - c Ab dc A = (3) dr r = rc rc rc 1 - rs The rate of diffusion at r = rc is given by Fick's first law, e.g. Eq. (3-66): dN A dc A - = 4rc2 De (4) dt dr r = rc 1- Exercise 11.26 (continued) Analysis: (continued) Where, N A is the amount of A diffused. Combining Eqs. (3) and (4): d N A 4rc De c A0 - cAb = dt r 1- c rs By material balance for the solute in the core: ( ) (5) dN A A d 4 3 4rc2 A drc =- rc = - dt M A dt 3 M A dt where: (6) A = initial mass of solute per unit volume of solid particle M A = molecular weight of solute Combining Eqs. (5) and (6): - A 1 1 2 - rc drc = De c A0 - c Ab dt M A rc rs Integrating Eq. (7) and applying the boundary condition: (7) rc = rs at t = 0 gives: A rs2 t= 6 De M A c A0 - c Ab For complete extraction, rs = 0, and Eq. (8) becomes: r 1- 3 c rs 2 r +2 c rs 3 (8) t= A rs2 6 De M A c A0 - c Ab (9) Equations (5) and (9) indicate that both the rate of extraction and time for extraction depend on the effective diffusivity of the solute in the particle, which must be measured experimentally. Exercise 11.27 Subject: Supercritical extraction of a slurry of 0.002 kmol/h of solid -carotene in 0.20 kmol/h of water with a solvent of recycle carbon dioxide in a single-stage contactor operating in the supercritical region to extract 99% of the carotene. Given: Design of a supercritical extractor using the Group-Contribution EOS of SkjoldJorgensen, at 353 K and 1013 bar, as shown in Fig. 1 of the cited article by Cygnarowicz and Seider. Find: Design of a similar extractor using the Peng-Robinson EOS with the Wang-Sandler mixing rules. Analysis: In the Cygnarowicz and Seider (C-S) article, the solubility of the carotene in water is neglected and a fit of experiment data for the solubility of carotene in CO2 is given in Fig. 2 of the C-S article. Thus, it was only necessary here to model a single-stage flash of CO2 and water. An initial run was made with the Aspen Plus program with the FLASH2 model under the following conditions, based on the PR EOS with the Wang-Sandler mixing rules (PRWS): Feed of 0.2 kmol/h of water. Solvent of 0.8004 kmol of CO2 and 0.0081 kmol/h of water. Single-stage extraction at 353 K and 1013 bar. The result of the flash was only a single phase containing both the feed and the solvent. Thus, use of the PR EOS with Wang-Sandler mixing rule did not give the required two equilibrium phases shown in the C-S article. When the extraction pressure was reduced to 101.3 bar (100 atm), the following results, which are close to the C-S results, were obtained: kmol/h: Component Feed Solvent CO2 Water Total: 0.0 0.2 0.2 0.8004 0.0081 0.8085 Raffinat e 0.0275 0.1950 0.2225 Extract 0.7729 0.0131 0.7860 However, at this much lower pressure, Fig. 2 of the C-S article shows that the mole fraction of carotene in the extract is an unacceptable 4 x 10-8 instead of the desired value of 0.0025. This is a huge difference and illustrates the importance of using reliable methods for estimating K-values in the supercritical region. In this case, the PR EOS with the Wang-Sandler mixing rules does not appear to be applicable. Exercise 11.28 Subject: Supercritical extraction of 100 mol/min of 10 mol% acetone in water by carbon dioxide in a staged contactor operating in the supercritical region to recover 98% of the acetone Given: Design of a supercritical extractor using the Group-Contribution EOS of SkjoldJorgensen, which is in good agreement with the experimental VLE data for acetone-CO2 at 313 K, as shown in the cited article by Cygnarowicz and Seider. Find: Design of a similar extractor using the Peng-Robinson EOS with the Wang-Sandler mixing rules. Analysis: In the Cygnarowicz and Seider (C-S) article, the experimental phase equilibria data for the acetone-CO2 system at 313 K is shown in their Fig. 1. The Aspen Plus program was used to prepare a similar plot based on the PR EOS with the Wang-Sandler mixing rules (PRWS). The result is shown on the next page. Comparison of that plot with Fig. 1 in the C-S article shows that the liquid composition is not predicted well. For example, when the pressure is 70 atm, Fig. 1 in the C-S article gives a carbon dioxide mole fraction of 0.84 in the liquid phase, while the PRWS prediction in the figure on the next page gives a mole fraction of 0.94. Therefore, the design given in Fig. 3 of the C-S article may not give the desired results for PRWS. Consequently, a new design was sought based on the following specifications in the C-S article: Extractor pressure > 70 atm Extractor temperature > 310 K CO2 feed rate > 200 mol/min Acetone recovery of at least 98% Water recovery of at least 6%. The Aspen Plus program was applied with the RADFRAC model, using the non-ideal algorithm with severe damping because of the difficulty in converging the calculations. However, initially, a single-stage isothermal flash was used to establish preliminary conditions of extractor pressure and temperature, and CO2 feed rate that would insure the presence of both vapor and liquid phases. One set of conditions that accomplished this was 350 K, 70 atm, and 300 mol/min of CO2. However, when these feed conditions were used with a 10-stage column, an acetone recovery of only 68.8% was achieved. Consequently, other combinations of feed conditions were tried, with some of the results as follows, where a solvent feed rate of 300 mol/min of CO2 was used in all cases: Pressure, Feed % Extraction % Extraction atm temperatures, K of acetone of water 70 350 68.8 4.53 70 360 79.1 5.64 70 370 89.7 6.97 80 370 97.9 7.37 The last result is the best, because it satisfies the specifications, except that not quite 98% of the acetone is extracted. The material balance for this case is given on the next page. Exercise 11.28 (continued) Analysis: (continued) Mol/min: Component Feed Solvent Raffinate Carbon dioxide 0.0 300.0 8.62 Acetone 10.0 0.0 0.21 Water 90.0 0.1 83.47 Total: 100.0 300.1 92.30 Extract 291.38 9.79 6.63 307.80 Compared to the C-S results, a higher temperature, lower pressure, smaller CO2 flow rate, and higher number of equilibrium stages are required. Unlike the C-S result, the above is not the optimal result, but is feasible depending on the reliability of the PR EOS with the WS mixing rules. Energy balances were made around each stage, resulting in stage temperatures varying from 355 to 365 K. Exercise 12.1 Subject: Revision of rate-based equations to account for entrainment and occlusion. Given:. Rate-based model Eqs. (12-4) to (12-18). Find: Modified equations Analysis: Entrainment: Let j = ratio of entrained liquid (in the exiting vapor) that leaves Stage j to the liquid leaving Stage j, yi ,n and yiI,n . Then, the entrained component liquid flow rate leaving Stage j = jxi, (1 + rjL ) L j . Correspondingly, the entrained component liquid flow rate entering Stage j = j+1xi,j+1 (1 + rjL 1 ) L j +1 . + Occlusion: Let j = ratio of occluded vapor (in the exiting liquid) that leaves Stage j to the vapor leaving Stage j, (1+ rjV )V j . Then the occluded component vapor flow rate leaving Stage j = j yi,j (1+ rjV )V j . Correspondingly, the occluded component vapor flow rate entering Stage j = j-1 yi,jV 1 (1 + r j -1 )V j -1 . The liquid-phase component material balance, Eq. (12-4), and vapor-phase component material balance, Eq. (12-5), become, respectively, M iLj (1 + rjL + j ) L j xi , j - L j -1 xi , j -1 - j +1 (1 + rjL 1 ) L j +1 xi , j +1 - f i ,Lj - N iLj = 0, i = 1, 2, ..., C , + , M iV, j (1 + rjV + j )V j yi , j - V j +1 yi , j +1 - j -1 (1 + rjV-1 )V j -1 yi , j -1 - f iVj + N iV, j = 0, , i =1, 2, ...., C The liquid-phase energy balance, Eq. (12-6), and vapor-phase energy balance, Eq. (12-7), become , respectively, E jL (1 + rjL + j ) L j H jL - L j -1 H jL-1 - j +1 (1 + rjL 1 ) L j +1 H jL+1 - H jLF + E V (1 + rjV + j )V j H V - V j +1 H V+1 - j -1 (1 + rjV-1 )V j -1H V-1 - H VF j j j j j C i =1 C C C i =1 fi ,Lj + Q L - e L = 0 j j i =1 f iVj + QV + eV = 0 , j j The total phase material balances, Eqs. (12-16) and (12-17), become, respectively, M TL, j (1 + rjL + j ) L j - L j -1 - j +1 (1 + rjL 1 ) L j +1 - + V M T , j (1 + rjV + j )V j - V j +1 - j -1 (1 + rjV-1 )V j -1 - f i ,Lj - N T , j = 0 f iVj + N T , j = 0 , i =1 Exercise 12.2 Subject: Revision of rate-based equations to account for a chemical reaction in the liquid phase. Given:. Rate-based model Eqs. (12-4) to (12-18). Assumption: Perfect mixing in the liquid on a stage. Find: Modified equations for: (a) chemical equilibrium (b) kinetic rate law Analysis: Let the chemical reaction be: A A + B B R R + SS where: = stoichiometric coefficient of component i where it is (+) for products R and S, and (-) for reactants A and B. Let the change in flow rate of component i in the liquid on stage j due to reaction be: i rk , j (1) k where: Mj = volumetric holdup of liquid on stage j. rk,j = chemical reaction rate, dck,j /dt , for the limiting reactant, k, on stage j. Let the rate law be a power-law expression, in terms of the reference limiting reactant, k, ni , j = - M j - Ef RT j dck , j dt = kf ,j c i =1 C ni i, j - c i =1 C mi i, j KC , j where: k f , j = A f e - ci , j = xi , j cT , j cT , j = total concentration of all components in the liquid on stage j K C , j = chemical equilibrium constant in terms of concentrations on stage j superscripts n, m = reaction orders for forward and backward reactions, respectively, which are related to the stoichiometric coefficients by i = mi - ni Exercise 12-2 (continued) Analysis (a): (continued) (a) Chemical equilibrium Eqs. (12-5), (12-7) to (12-15), and (12-17) to (12-18) apply . In Eq. (12-4), liquid mole fractions xi,j are in chemical equilibrium as governed by the chemical equilibrium constant and the stoichiometry, such that: M iLj (1 + rjL ) L j xi , j - L j -1 xi , j -1 - ni , j - f i ,Lj - N iLj = 0, i = 1, 2, ..., C , , In energy balance, Eq. (12-6), liquid enthalpies include heats of formation. Eq. (12-16) becomes: C C i M TL, j (1 + rjL ) L j - L j -1 - nk , j - f i ,Lj - N T , j = 0 i =1 k i =1 (b) Kinetic rate law Eqs. (12-5), (12-7) to (12-15), and (12-17) to (12-18) apply. Eq. (12-4) becomes: M iLj (1 + rjL ) L j xi , j - L j -1 xi , j -1 - ni , j - f i ,Lj - N iLj = 0, i = 1, 2, ..., C , , where ni , j is computed from the reaction rate equation, Eq. (1), above. In Eq. (12-6), enthalpies must include heat of formation. Eq. (12-16) becomes: C C i M TL, j (1 + rjL ) L j - L j -1 - nk , j - f i ,Lj - N T , j = 0 i =1 k i =1 Exercise 12.3 Subject: Reduction of the number of equations in the rate-based model. Given:. Rate-based model based on Eqs. (12-4) to (12-18). Find: Method to reduce the number of equations. Analysis: either: In Chapter 10, equilibrium-stage models are written for each equilibrium stage in terms of Case 1: Case 2: xi , j , yi , j , L j ,V j , and Tj li , j , i , j , and Tj (2C + 3) variables (2C + 1) variables In Section 12.1, the rate-based model is written in terms of : xi , j , yi , j , xiI, j , yiI, j , N i , j , TjL , TjV , TjI , L j , and V j The equations could be rewritten to replace: (5C + 5) variables xi , j , yi , j , L j , and V j by li , j and i , j to give (5C + 3) variables. For example, Eq. (12-4) would become: M iLj (1 + rjL )li , j - li , j -1 - f i ,Lj - N iLj = 0, i = 1, 2, ..., C , , This would eliminate Eqs. (12-11) and (12-12). In Ref. 16, Taylor and his colleagues use component flow rates. This has the advantage that the component material balances are linear and fewer equations are needed because there are two fewer variables per stage. In Ref. 17, Taylor and his colleagues use component mole fractions and total flow rates. This choice is preferable when total flow rates are useful, such as when taking into account entrainment, occlusion, backmixing, weeping, and tray hydraulics. The percentage reduction in the number of variables and equations for the rate-based model is not nearly as great as for the equilibrium model because, at least not for many components, that number is dominated by the 5C term. Exercise 12.4 Subject: Mass-transfer rates and tray efficiencies from data for a perfectly mixed tray. Given: Vapor-liquid traffic, component mole fractions, and mass-transfer coefficients for a tray in a column separating acetone (1), methanol (2), and water (3) at 14.7 psia Find: (a) Component molar diffusion rates. (b) Component mass-transfer rates. (c) Murphree vapor tray efficiencies. Analysis: Because mass-transfer coefficients are given for the gas phase, work with the values of yi ,n and yiI,n . Because rates instead of fluxes are given, the equations developed in this section are used with rates rather than fluxes. (a) Compute the reciprocal rate functions, R, from Eqs. (12-31) and (12-32), assuming linear mole fraction gradients such that zi can be replaced by ( yi + yiI ) / 2 . Thus, z1 = (0.4913 + 0.5291) / 2 = 0.5102 z2 = (0.4203 + 0.4070) / 2 = 0.4136 z3 = (0.0884 + 0.0639) / 2 = 0.0762 z z z 0.5102 0.4136 0.0762 V R11 = 1 + 2 + 3 = + + = 0.000509 h / lbmol k13 k12 k13 2154 1750 2154 z z z 0.4136 0.5102 0.0762 V R22 = 2 + 1 + 3 = + + = 0.000487 h / lbmol k23 k 21 k 23 2503 1750 2503 V R12 = - z1 1 1 1 1 = -0.5102 - = -0.000055 h / lbmol - 1750 2154 k12 k13 1 1 1 1 = -0.4136 - = -0.000071 h / lbmol - 1750 2503 k 21 k 23 V R21 = - z2 In matrix form: RV = 0.00050