Ind. Eng. Chem. Process Des. Dev. 1985, 2 4 , 707-773
x = liquid phase mole fraction z = lattice coordination number
Subscripts A = non-alcohol B = alcohol c = chemical contribution of a property calcd = calculated exptl = experimental N , = number of carbon atoms in an alcohol p = physical contribution Superscripts m = property of mixing * = reference solutions containingmolecules of only one type Greek Letters Aho = heat of hydrogen bond formation, = -7 kcal/g-mol 8, = area fraction of group m
&, = volume fraction of the alcohol = volume fraction of alcohol monomer $ J *=~volume ~ fraction of alcohol monomer in the pure alcohol state
L i t e r a t u r e Cited Anand, S. G.; Groiier, J. P. E.; Klyohara, 0.; Halpin, G. J.; Benson, G. C. J . Chem. f n g . Data 1975, 20, 184. Andersen, B. D.; Ryttlng, J. H.; Lindenbaum, S.; Hlguchl, T. J . Phys. Chem. 1975, 79, 2340. Benson, G. C.; Murakami, S.; Jones, D. E. G. J . Chem. Thermodyn. 1971, 3 , 719. Bondi, A. “Physical Properties of Crystals, Llqulds and Glasses”; Wlley: New York, 1968. Brusset, H.; Desgranges, M.; Lecoq, J. C. Can. J . Chem. 1972, 50, 3207. Coomber, B. A.; Wormald, C. W. J . Chem. 7hernm$yn. 1978. 8 , 793. Doan-Nguyen, T. H.; Vera, J. H.; Ratcliff, G. A. J . Chem. f n g . Data 1978, 2 3 , 218. Findlay, T. J. V. Aust. J . Chem. 1981, 14, 520. Flory, P. J. J . Chem. Phys. 1944, 12, 425. Fredenslund, Aa.; Gmehling, J.; Rasmussen, P. ”Vapor-Llquld Equilibrium Using UNIFAC”, Elsevler: Amsterdam, 1977.
707
Gmehllng, J.; Rasmussen, P.; Fredenslund, Aa. Ind. f n g . Chem. Process 0es:Dev. 1982,21, 118. Grolier, J. P. E. Inti. Data Series, Thermodynamic Research Center, Texas A&M University, College Station, TX, 1978. Klyoharo, 0.; Anand, S. C.; Benson, G. C. J . Chem. Thermodyn. 1974, 6 , 355. __.
Lei, T. T.; Doan-Nguyen, T. H.; Vera, J. H.; Ratcllff, G. A. Can. J . Chem. f n g . 1978. 56, 358. Murakami, S.; Benson, G. C. Bull. Chem. Soc. Jpn. 1973, 46, 74. Murakami, S.; Fujlshiro, R. Bull. Chem. Soc. Jpn. 1866, 39, 720. Nagata, I.2.Phys. Chem. (Le/pz@)1973, 252, 305. Nagata, I.; Kazama. K. J . Chem. fng. Data 1977, 22, 79. Nagata, I.; Ohta, T. Chem. f n g . Sci. 1978, 33, 177. Nagata, I.; Ohta. T.; Nakagawa, S. J . Chem. f n g . Jpn. 1976, 9 , 276. Nagata, I.; Yamada, T.; Nakagawa, S. J. Chem. f n g . Data 1975, 20, 268. Nakanishl, K.; Touhara, H.; Watanabe, N. Bull. Chem. Soc. Jpn. 1970, 43, 2671. Nguyen, T. H.; Ratcliff, 0. A. Can. J . Chem. f n g . 1974. 52, 641. Posa, C. G.; Nunez, L.; Viiiar, E. J . Chem. Thermodyn. 1972, 4 , 275. Prausnk, J. M. ‘Molecular Thermodynamics of Fluld-Phase Equilibria”; Prentlce-Hail, Engiewod Cliffs, NJ, 1969. Redlich, 0.; Klster, A. J. J. Chem. Phys. 1947, 15, 849. Renon, H.; Prausnltz. J. M. Chem. f n g . Sci. 1967. 22, 299. Rupp, W.; Hetzei, S.; Ozlui, I.; Tasslos, D. Ind. f n g . Chem. Process D e s . Dev. 1984, 23, 391. Ryttlng, J. H.; Anderson, B. D.; Higuchl, T. J . Phys. Chem. 1978, 82. 2240. Sakai, Y.; Sadaoka, Y.; Yamamoto, T. Bull. Chem. Soc. Jpn. 1973, 46, 3575. Sarlni, C. G.; Winterhalter, D. R.; Van Ness, H. C. J . Chem. f n g . Data 1965, 10, 188. Siman, J. E.; Vera, J. H. Can. J . Chem. f n g . 1979, 57, 355. Singh, S.; Rao, C. N. R. J . h y s . Chem. 1987, 71. 1074. SkjoWorgensen, S.; Rasmussen, P.; Fredenslund, Aa. Chem. fng Sci. 1980, 35, 2389. Stathls, P. J. M.S. Thesis, N.J. Institute of Technology, Newark, NJ 1983. Tucker, E. E.; Christian, S. D. J . Phys. C b m . 1977, 8 1 , 1295. Van Ness, H. C.; Abbott, M. M. Inti. Data Series, Texas A&M Unlversity, College Station, TX, 1978. Villamanan, M. A.; Casanova, C.; Roux. A.; Grollier, J-P. J . Chem. f n g . Data 1982, 27, 89. Welmer, R. F.; Prausnltz, J. M. F!vdrocerbon Process. 1965, 44, 237.
.
Received for review February 24, 1984 Revised manuscript received July 13, 1984 Accepted August 6, 1984
Heat- Integrated Distillation Columns for Ternary Separations Hao-Chleh Cheng and Wllllam L. Luyben’ Depadment of Chemlcal Engineering, Lehbh University, Bethlehem, Pennsylvanla 180 15
The steady-state designs of several heat-integrated distillation systems were studied for a ternary separation. These heat-integration configurations were compared with configurations without heat integration on the basis of energy consumption. A direct comparison of the best configuration found in this work to the optimum configuration reported by Taka& et al. was also made in tqms of energy consumption and equipment cost. The specific chemical system studied was the benzene-toluene-m-xylene separation with high purities. Feed compositions were 25/50/25, the same as used by Takama et ai. The best scheme was found to be the prefractionator/sidestream column configuration with reverse heat integration with low operating pressure. Reductions in energy consumption of 35 % and 45 % were obtained ’for simple and complex configurations, respectively, by Incorporation of heat integration. The best scheme found in this work consumes 20-30% less energy than the scheme proposed by Takama et ai. without additional capital investment.
Introduction Energy Conservation in Distillation. Distillation is
an indispensable part of the petroleum, natural gas, and chemical industries and is often the major energy consumer. Mix et al. (1978) reported that distillation consumes about 3% of the energy used in the United States and estimated that a 10% energy saving in distillation would amount to saving nearly 100000 bbl of oil/day. 0196-4305/85/1724-0707$07.50/0
Frequently, over 95% of the energy added to the reboiler of a distillation column is removed by the cooling water at the top of the column. Many energy-saving configurations have been known for a long time, but only recently, because of the drastic increase in energy cost, have they become attractive. The approaches to effect energy conservation in distillation generally fall into two categories: (1) heat integration or multiple effect methods and (2) 0 1985 American Chemical Society
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Ind. Eng. Chem. Process Des. Dev., Vol. 24, No. 3, 1985
heat-pump or vapor recompression methods. The concept of the heat pump is to use the overhead vapor as a heat source for the reboiler of the same column. The adverse temperature difference is overcome by adding compression work to the vapor to raise its pressure and condensing temperature. Thus a compression system analogous to a refrigeration cycle is needed to implement a heat pump system. Either an external medium or the process stream may serve as the working fluid. Null (1976) presented guidelines for conditions under which heat pumps might be economical substitutes for. conventional distillation process design. The most prodsing condition was whenever direct refrigeration or chilled water is required for condensation in conventional columns. Shaner (1978) applied heat pumps to air separation plants and a C3 splitter. In both cases, refrigeration was used for conventional distillation design. Distillation schemes with heat pumps require considerable compression work when the temperature difference to be overcome is large. For this reason, they are usually not used in multicomponent distillation or for high relative volatility systems. Heat integration involves the use of the heat load of an overhead condenser of a column within a configuration as the heat source for the reboiler of another column within the same configuration. Robinson and Gilliland (1950) illustrated three ways in which multieffect principles can be used to make energy supply and removal to or from a distillation process more efficient. Tyreus and Luyben (1975) presented a case study of double-effect column design for a C3 splitter and methanol-water, both using steam and cooling water. They concluded that a one-tower distillation system with a steam consumption over about 30 000 lb/h may be suitable for heat integration. Reversibility is another factor which has been explored by many workers. In multicomponent systems, complex configurations such as prefractionators and sidestream columns have been shown to require less energy than conventional single-feed, two-product columns. Petlyuk et al. (1965) were among the first to study reversibility for multicomponent distillation and complex distillation configurations for ternary separations. One of their schemes was discussed by Stupin and Lockhart (1972). Doukas and Luyben (1978) studied both simple and complex confiiations for the BTX (benzene, toluene, and xylene) separation. For some feed compositions,they found that either a single sidestream column or a prefractionator/sidestream column was more economical than conventional configurations. They also reported that the dominant part of the total annual cost is the cost of utilities. In connection with a multi-objective optimization study, Takama et al. (1982) proposed a number of alternatives for BTX separation and found a best compromise solution which involved a heat integrated prefractionatorlsidestream column configuration. When more than three components are to be separated by distillation, the number of alternatives becomes enormous. Several studies have been made to systematically identify the best or one of the best configurations out of many possibilities. Studies were generally confined to sequencing simple columns for best heat source and heat sink matching. King (1971) reported four general heuristics to screen out some of the large number of possible schemes. Westerberg and Stephanopoulos (1975) proposed an algorithmic technique in which optimization methods were used to determine the most economical configuration. Elaahi and Luyben (1983) explored several complex configurations for a four-component distillation.
t
I
A
38
44
110.8.C
25
17
-
07!
105 mm
mm
r
I
I1I.l*C
I
1 1
m-Xgls"* Preducl
1136.C
Figure 1. Case 1. (Numbers above streamlines without units designate flow rates in kg-mol/h. Numbers inside heat exchangers designate heat duty in loe kcal/h. Numbers inside columns designate tray numbers.)
Scope of This Work. This work is basically an extension of that of Doukas and Luyben (1978) to include heat integration in complex configurations of distillation columns. The first part of this study was concerned with an energy consumption study for the separation of a given mixture of benzene, toluene, and m-xylene to desired specifications. Heat integration was incorporated in both conventional configurations and complex configurations. Operating prmure, which is an important factor in energy consumption, was also studied. The second part of this work was a direct comparison of the best compromise solution reported by Takama et al. (1982) with other alternatives. For this reason, the same feed and product specifications were used. Detailed information on the process conditions was obtained from Umeda (1982). The assumptions made in this study were (a) 100% tray efficiency; (b) ideal gas and ideal liquid behavior for vapor-liquid equilibrium; (c) saturated liquid reflux; (d) minimum temperature difference at the pinch end of heat exchanger: 10 "C; (e) temperature difference in reboiler/condenser: 15 OC; (f) pressure drop per tray: 2.5 mmHg in vacuum columns and 5 mmHg in atmospheric or pressure columns; (g) pressure drop across condenser: 50 mmHg, under vacuum, and 125 mmHg, under atmospheric or high pressure. Configurations Studied There were 11 cases studied in this work. The general design specificationsfor these cases were (a) feed rate: 300 kg-mol/h; (b) feed temperature: 15.5 "C; (c) feed composition: benzene, 25 mol %; toluene, 50 mol %; m-xylene, 25 mol %; (d) product specification: benzene, 99.9 mol %; toluene, 99.9 mol 90;m-xylene, 99.9 mol % ,except for cases 10 and 11, in which higher product purities were used. Configurations without Heat Integration. Case 1. This is a configuration of two conventional columns in sequence (see Figure 1)where the bottom product of the first column is the feed to the second column. This type of Configuration is called Light Out First (LOF). Pressures in the reflux drums of both columns were fixed at atmospheric pressure. Feed to the first column is preheated by overhead vapor of the second column and then by the bottom product stream of the second column. From the first column, benzene was removed as the top product. Toluene was the overhead product of the second column,
Ind. Eng. Chem.
13
5 4 4.C
709
I 1 ,,ox
Figure 2. Case 2. I CI1
Process Des. Dev., Vol. 24, No. 3, 1985
Figure 4. Case 4.
35; 976.C
1 '-
I
r
i!$ 190
224996
7.991
I
112 4 %
t
Figure 5. Case 5.
Figure 3. Case 3.
and m-xylene was the bottom product. Case 2. This case,also an LOF configuration (seeFigure 2), was the same as case 1 except that lower column pressures were used. The reflux drum pressures of both columns were set by the bubble point pressures of the overhead products at 50 "C which allowed cooling water to be used as the cooling medium. The first column operated at 264.5 mmHg and the second at 90 mmHg. Case 3. This case (see Figure 3) consisted of a prefractionator and a sidestream column (PF). The first column made a split between the heaviest and the lightest components. Bottom and top products of the first column were fed to different feed trays in the second column which separated them into the three products. The fvst column was operated with a partial condenser, producing a vapor product which was fed into the second column. The pressure in the reflux drum of the second column was fixed at atmospheric pressure. The pressure of the first column was specified so that a pressure drop of 1psi was obtained to drive the vapor product from the first column to the second column. Feed to the first column was preheated by the overhead vapor of the first column and by the bottom product of the second column. Case 4. This case (see Figure 4) was the same as case 3 (a P F confiiation) except that the reflux drum pressure of the second column was set by the bubble point pressure (264.5 mmHg) of the overhead product at 50 OC,instead of at atmospheric pressure. The pressure drop between the reflux drum of the first column and the upper section
1
224 994
Figure 6. Case 6.
feed plate of the second column was assumed to be 1psi. Configurations with Heat Integration. Case 5. In this case (see Figure 5) two conventional columns were heat integrated in the same direction as process flows (an LOF confiiation with forward heat integration, LOF/F). The vapor from the first column was used as a heat source in the reboilei of the second column. The pressure of the reflux drum of the second column was set by the bubble point pressure of the overhead distillate at a temperature of 50 "C. The pressure in the reflux drum of the first column was set by the bubble pressure of the overhead
710
I d . Eng. Chem. Process Des. Dev., Vol. 24, No. 3, 1985 Food
1
300 15.5OC
Product
149.997139-C
Toluenr Product
31 *233.S°C
ATMST
Figure 7. Case 7. 25B.3.C
I 19B.C
Figure 9. Case 9. I32'C 9OOC
II
ATM
182.8.C
-..a. 1014.C
j
Figure 8. Case 8.
distillate at a temperature that was greater than the bottom temperature of the second column by 15 "C. The feed to the first column was preheated by the bottom product stream of the second column and by the overhead product of the first column. The heat supply from the overhead vapor of the first column did not provide enough heat to the second column and therefore an auxiliary steam reboiler was needed. Case 6. This case (see Figure 6) differed from case 5 in the direction of heat integration (anLOF configuration with reverse heat integration, LOF/R). The overhead vapor of the second column was used as a heat source to the reboiler of the first column. The reflux drum pressure of the fiist column was set by the bubble point pressure of the overhead distillate a t 50 "C. The pressure of the second column was set by the bubble point pressure of the overhead distillate at a temperature that was greater than the bottom temperature of the first column by 15 "C. The feed to the fiist column was not preheated because the vapor from the second column could provide more heat than was needed in the first column. Thus a feed preheater would only increase the heat that would have to be removed in the auxiliary condenser on the second column and would increase capital cost. Case 7. This was a configuration with a prefractionator and a sidestream column that were heat integrated in the direction of process flow (PF/F) (see Figure 7). Pressures in the reflux drums were set by the same reasoning as in case 5. The heat from the overhead vapor of the first column was more than that required in the second column. Therefore part of it was used to preheat the feed stream to the first column. The bidestream product and the bottom product of the second column were also used to preheat the feed to the first column.
-
m -Xylrne Product
150
Figure 10. Case 10.
Case 8. This case (see Figure 8) differed from case 7 in that the direction of heat integration was opposite to process flow (PF/R). The pressures in the reflux drums of both columns were set by the same reasoning as in case 6. The feed to the first column was preheated in a series of heat exchangers by three streams: (1)the sidestream of the second column, (2) the bottoms from the second colurrin and (3) the overhead product of the second column.. Additional heat had to be added in an auxiliary steam reboiler in the first column. Case 9. This case, also a PF/R configuration, used all the same design assumptions and tray numbers as in case 10 which will be described later. The only difference was that product purities were reduced to 0.999 by changing reflux ratios and product rates. This permitted a direct comparison among case 9 and cases 1-8 with the same purity products. The change in product purities reduced the amount of heat available in the overhead vapor from the second column. Therefore an auxiliary steam reboiler was needed for the first column. Vapor from the second column could not be used to preheat the feed to the first column as in case 10. However, the bottom product of the second column was used as another source to preheat the feed to the first column. See Figure 9. Case 10. This was a PF/R configuration (see Figure 10) presented by Takama et al. (1982). Heat integration direction was the same as in case 8. In this case the design basis was slightly different from those in cases 1-8: (a)
Ind. Eng. Chem. Process Des. Dev., Vol. 24, No. 3, 1985 711 Table I. Steadv-State Design SDecifications
feed, kg-mol/h
DZST, kg-mol/h BOT, kg-mol/h NT NF RR PD. mmHe
feed, kg-mol/h
DZST, kg-mol/h
LS,kg-mol/h
BOT, kg-mol/h NT NF NS RR
feed, kg-mol/h
DZST, kg-mol/h BOT, kg-mol/h NT NF RR PD, mmHa
feed, kg-mol/h
DZST, kg-mol/h LS, kg-mol/h BOT, kg-mol/h NT NF NS RR
feed, kg-mol/h
DZST, kg-mol/h LS, kg-mol/h BOT, kg-mol/h NT NS NF RR
case 1 col. 1 col. 2 300 224.994 75.006 150.0 224.994 74.994 38 44 25 17 2.92 1.47 760 760 case 3 col. 1 col. 2 300 150/150 74.994 150 --150.003 75.003 150 29 62 11 55/27 40 0.95 5.77 case 5 col. 1 col. 2 300 224.994 149.997 75.006 74.997 224.994 35 39 23 14 1.27 2.61 1960 90 case 7 col. 1 col. 2 300 150/150 150 75.006 _-150 150 74.994 31 60 13 52/21 40 1.18 3.17 case 9 col. 1 col. 2 300 150/150 150 75.024 149.997 150 74.070 38 88 53 12 74/31 0.83 3.97
___
___ ___
___
feed, kg-mol/ h DZST, kg-mol/h LS, kg-mol/h BOT, kg-mol/h
NT NF NS RR
col. 1 300 150
___
150 29 11
__0.76
case 2 col. 1 col. 2 300 224.994 75.006 150.0 224.994 74.994 36 33 22 13 2.00 1.12 264.5 90 case 4 col. 1 col. 2 150/150 300 74.988 150 150.006 75.006 150 57 28 50/24 10 _-38 4.55 0.72 case 6 col. 1 col. 2 300 224.994 75.006 150.0 224.994 74.994 37 44 22 16 1.66 1.50 264.5 795 case 8 col. 1 col. 2 300 150/150 150 75.00 _-150.00 150 75.00 29 82 10 69/29 52 0.61 2.54 case 10 col. 1 col. 2 300 150/150 150 74.97 150.02 150 75.01 38 88 53 14 74/31 1.2 7.3 case 11 col. 2 150/150 74.955 150.036 75.009 82 67/30 53 4.01
___
___
___
___
PE, mmHg TD, OC TE, OC QE, lo6 kcal/h QD, lo6 kcal/h QT, lo6 kcal/h
PD, mmHg PE, mmHg TD, "C TE, O C QE, lo6 kcal/h QD, lo6 kcal/h QT, lo6 kcal/h
PE, mmHg TD, "C TE, O C QE, lo6 kcal/h QD, lo6 kcal/h QE, lo6 kcal/h QT, lo6 kcal/h
PD, mmHg PE, mmHg TD, O C TE, "C QB, lo6 kcal/h QD, lo6 kcal/h QE, lo6 kcal/h QT, lo6 kcal/h
PD,mmHg
PB, mmHg TD, "C TE, O C QE, lo6 kcal/h QD, lo6 kcal/h QE, lo6 kcal/h QT, lo6 kcal/h
PD, mmHg PB, mmHg TD, "C TE, O C QE, lo6 kcal/h QD, lo8 kcal/h QE, lo6 kcal/h QT, lo6 kcal/h
case col. 1 1075 80.3 131.1 2.29 2.19 5.25
1 col. 2 1105 110.8 153.7 2.96 3.01 5.25
case 2 col. 1 col. 2 404.5 222.5 49.5 49.6 96.8 98.7 2.20 2.60 1.77 2.90 4.80 4.80
case 3 col. 1 col. 2 975 760 1240 1195 107.6 80.2 141.2 159.9 2.44 2.57 1.16 3.78 5.01 5.01
case 4 col. 1 col. 2 385 264.5 505 457 77.6 49.4 108.0 121.2 2.33 2.05 0.93 3.27 4.38 4.38
case 5 col. 1 col. 2 2280 227.5 49.6 114.5 162.9 99.3 2.96 2.03 3.10 1.87 2.96 0.16 3.12 3.12 case 7 col. 1 col. 2 2425 264.5 2705 464.5 137.2 49.5 175.8 121.8 2.84 1.45 2.37 2.46 2.84 0.0 2.84 2.84 case 9 col. 1 col. 2 988 7372 1653 8305 101.4 178.0 153.2 258.2 2.45 2.31 2.22 2.19 0.36 2.31 3.00" 3.OOa case col. 1 176 298.5 49.3 91.5 2.53 2.28 0.0 2.82
case 6 col. 1 col. 2 407 1140 49.5 112.4 97.0 155.0 2.47 3.19 1.57 3.03 0.0 3.19 3.19 3.19 case 8 col. 1 col. 2 176 1620 298.5 2155 49.4 107.0 91.4 182.8 2.17 2.05 2.04 1.86 0.31 2.05 2.36 2.36 case 10 col. 1 col. 2 988 7372 1653 8305 101.4 178.0 153.3 258.3 2.63 3.76 2.67 3.64 0.0 3.76 4.Oga 4.09 11 col. 2 1620 2155 107.0 182.9 2.82 2.64 2.82 2.82
Includes steam used in preheater
pressures in the reflux drums of the first and second columns were set at 1.3 and 9.7 atm, respectively; (b) minimum temperature differences at the pinch ends of heat exchangers were 26 OC;(c) temperature difference in the reboiler/condenser was 25 OC;(d) pressure drops per tray were 7.5 mmHg/tray in the first column and 6.3 mmHg/tray in the second column; (e) pressure drops across condensers were 0.5 atm. The product purities were very high: 0.999913 (benzene), 0.999733 (toluene), and 0.999821 (m-xylene).
Case 11. This case was the same as case 8 but with the very high purities featured in case 10. Because of the increased reflux ratios required, the heat available in the vapor from the second column was now more than that required in the first column. Feed preheating in the first column was also modified since increasing the feed preheat did not reduce energy consumption. See Figure 11. Column Design Procedures. Feed flash calculations were performed after preliminary estimates of feed preheating had been made. For conventional columns,
712
Ind. Eng. Chem. Process Des. Dev., Vol. 24, No. 3, 1985 Table 11. Energy Reauirement for Each Configuration energy consump, case description Duritv lo6 kcal/h Without Heat Integration 1 LOF (1 atm) 0.999 5.25 2 LOF (vacuum) 0.999 4.80 3 P F (1 atm) 0.999 5.01 4 P F (vacuum) 0.999 4.38 With Heat LOF/F (vacuum) LOF/R (vacuum) P F / F (vacuum) PF/R (vacuum) PF/R (1.3 atm) P F / R (1.3 atm) PF/R (vacuum)
5 6 7 8 9 10' 11
Figure 11. Case 11.
short-cut design methods were used to find the total number of theoretical trays and the approximate feed location. The reflux ratio was set at 1.1times the minimum reflux as calculated from the Underwood (1948) equations. For complex columns, each section was preliminarily treated as a pseudo-conventional column to which shortcut methods could be applied to find the number of trays and feed location. The number of trays in each section was set equal to twice the minimum number of trays as calculated from the Fenske (1932) equation. After an approximate design was obtained, refinements were made by changing the total number of trays and feed and sidestream locations in a rigorous column design program. A rigorous computer program, based on the WangHenke (1966) algorithm, was used to obtain specified product purities by changing reflux ratio and product flow rates once the total number of trays and the feed locations were fixed. The superficial vapor velocity, based on the Fair (1961) correlation, was used to calculate column diameters. Hot streams were used to preheat feed streams to the first and second columns if feed preheating could reduce external heat duty. Since pressures and temperatures of potential heating sources for feed preheating depend on the final column design, an evolutionary development of the feed preheating systems was made. Design details for all cases are given in Table I. QE is the reboiler heat suplied by an external source in each column. QT is the total external heat supplied. QB is the energy consumed in each reboiler from either external sources or from heat integration. Results and Discussion Table I1 summarizes the energy consumptions of all cases. Without heat integration, case 4 (prefractionator/sidestream column at low pressures) showed the minimum energy consumption. The lower energy consumption of case 2 compared to case 3 shows that a conventional configuration with lower operating pressure may be better than a complex configuration with higher operating pressure. With heat integration, case 8 was found to be the best, consuming 45% less energy than the best conventional c a s (2.36 vs. 4.38 X lo6kcal/h). Complex configurations with heat integration were more energy-efficient than conventional configurations with heat integration. Reverse heat integration appears to be better than forward heat integration for this system. This is consistent with the results of Chiang and Luyben (1983), who studied only binary systems. Reverse heat integration gives the lowest pressure in the high-pressure column. If relative volatilities are
Integration 0.999 0.999 0.999 0.999 0.999 0.9998 0.9998
3.12 3.19 2.84 2.36 3.00 4.09 2.82
'Chiyoda Study (Umeda, 1982). Note: comparison of energy consumption: case 8 vs. case 9: (3.00 - 2.36)/3.00 = 21% at lower purity; case 11 vs. case 10: (4.09 - 2.82)/4.09 = 31% at higher purity. Table 111. Equipment Sizing and Capital Estimation for Case 11 exch. no.
Q, lo6 kcal/h 2.28 2.53 0.11 0.17 0.52 2.82 0.35 0.28 0.17
total
col. 1 2
total
I. Heat Exchangers LMTD, (h m2 OC) O C
U,,, kcal/ 500 400 500
400 400 800 400 400 500
19.6 16.9 70.3 26.2 16.0 20.0 22.9 28.9 22.5
A, m2 232.6 374.3 3.1 16.2 81.3 176.2 38.2 24.2 15.1
cap. cost, $ 21000 30500 2300 4700 10800 18000 7200 5800 4500
961.3
105400
11. Column flooding DIA, price/tray, cap. cost, factor NT m $ $ 0.691 29 2.19 1030 29 870 0.617 82 1.96 800 65 600 95 470
total investment: $200 870
strong functions of pressure, the lower pressure can reduce total energy consumption. Cases 8 and 11 consumed less energy than the cases proposed by Takama et al. (1982) (cases 9 and 10) at both lower and higher purity levels by 20 and 30%, respectively. This reduction is due to (1) lower column operating pressures leading to reduced reflux ratios, (2) lower temperature differences in reboiler/condenser, and (3) lower temperature differences in exchangers. In addition to saving in the total amount of energy consumed, cases 8 and 11 had significantly lower base temperatures. The base temperature of cases 8 and 11 was only 183 OC in the column with the reboiler. In cases 9 and 10, the base temperature was 258 "C. This means that higher pressure steam or other high-temperature heating medium, like Dowtherm, would be required. This made the energy conservation of cases 8 and 11 even greater. The steam consumption for achieving vacuum was assumed to be insignificant (McKetta, 1983). Tables 111and IV compare the equipment sizes and costs for case 11and case 10. Heat exchangers were assumed to be fixed-tube-sheet type made of carbon steel. Distillation columns were assumed to be carbon steel with sieve trays. Unit costs of heat exchangers and column trays were
Ind. Eng. Chem. Process Des.
Table IV. Equipment Sizing and Capital Estimation for Case 10
exch. no. 1 2 3 4
Q, IO6
kcal/h 0.94 1.73 2.63 0.60 0.41 0.67 0.38 3.76 0.33 0.41 0.64 0.55
5 6 7 8 9 10 11 12
total
col. 1 2
total
I. Heat Exchangers U,,kcall LMTD, ( h m 2 “C) 400 500 400 400 500 400 400 800 400 500 500 500
“C 53.2 74.2 26.1 66.4 143.0 31.2 42.2 20 61 60.7 44.7 56.8
A , m2 44.2 46.6 251.9 22.6 5.7 53.7 22.5 235.0 13.5 13.5 28.4 19.4
cap. cost,$ 7600 8000 23500 5700 2700 8800 5660 21000 4350 4350 6100 5150
757.0
102910
11. Column flooding DIA, price/tray, cap. cost, factor- N T m i $ 0.691 38 1.77 700.0 26 600 0.617 88 2.08 910.0 80080 106680
total investment: $209 590
obtained from Peters and Timmerhaus (1980). Only the cost of trays was considered since the costa of the column shells were almost the same. This was due to their having approximately the same total weight. Total heat transfer area for case 11 is 27% more than that used in case 10. However, the total cost for heat exchangers was about the same because the heat transfer area of case 11 was mostly concentrated in three big heat exchangers. This reduced the average unit price. By comparison of the total cost of all columns and heat exchangers, both cases 11 and 10 need about the same investment. I t is necessary, of course, to compare the difficulty of operation. There are heuristic rules for separation sequencing which recommend avoiding vacuum distillation (Nadgir and Liu, 1983). To make a final decision on the selection of energysaving systems, it is important to recognize that the reduction in energy consumption achieved by the proposed method may be offset by somewhat poorer plant operability due to the low operating pressure. The controllability of these alternative schemes will be the subject of a future paper.
Dev., Vol. 24, No. 3, 1985
713
Conclusions The prefractionator/sidestream configuration with reverse heat integration at low operating pressure was the most energy-saving scheme. The most energy-saving scheme found in this work consumes about 30% less energy than the scheme proposed as the best compromise solution by Takama et al. (1982)with about the same investment. Nomenclature BOT = bottom flow rate, kg-mol/h DZA = column diameter, m DZST = distillate flow, kg-mol/h NT = total number of trays NF = feed tray NS = sidestream tray PD = reflux drum pressure, mmHg PB = column bottom pressure, mmHg QD = condenser heat duty, lo6 kcal/h QB = reboiler heat duty, lo6 kcal/h QE = heat duty from external source for reboiler, lo6 kcal/h QT = total heat duty from external source, lo6 kcal/h RR = reflux ratio TD = distillate temperature, “C TB = reboiler temperature, “C Registry No. Benzene, 71-43-2;toluene, 108-88-3; m-xylene, 108-38-3.
Literature Cited Chiang, T. P.; Luyben, W. L. I d . Eng. Chem. frocess Des. Dev. 1983,22, 175. Doukas. N.; Luyben, W. L. Ind. Eng. Chem. frocess D e s . Dev. 1978, 17, 272. Elaahi. A.; Luyben, W. L. Ind. Eng. Process Des. D e v . 1983, 22, 80. Fair, J. R. Petro.lChem. Eng. 1961, 33(9),211. Fenske, M. R. Ind. Eng. Chem. 1932, 2 4 , 482. King, C. J. “Separatlon Processes”; McQraw-HIII: New York, 1971. McKetta, J. J., Ed. “Encycbpedla of Chemical Processing and Design”; 1983; Voi. 17, p 177. Mix, T. J.; Deweck, J. S.; Weinberg, M.; Armstrong, R. C. Chem. Eng. frog. 1978, 74(4), 49. Nadglr, V. M.; Lin, Y. A. A I C M J . 1089, 29(6), 926. Null, H. R. Chem. Eng. frog. 1978, 72(7), 58. Peters, M. S.; Timmerhaus, K. D. “Plant Design and Economics for Chemical Engineers”, 3rd ed.; McGraw-Hill: New York, 1980. Petlyuk, F. V.; Plantonov. V. M.; Slavlnskl, D. M. Int. Chem. Eng. 1965, 5 , 555. Roblnson, C. S.; Glliiland, E. R. “Elements of Fractional Dlstlllatlon”,4th ed., McQraw-nil: New York, 1950. Shaner, R. L. Chem. Eng. frog. 1978, 74(5), 47. Stupln, W. J.; Lockhart, F. J.. Chem. Eng. f r o g . 1972, 88(10), 71. Takama, N.; Kwlyama, K.; NiMa, K.; Kinoshita, A.; Shlroko, K.; Umeda, T. chem. Eng. frog. 1982, 78(9), 83. Tyreus, 9. D.; Luyben, W. L. Hydrocerbon Process. 1975, 54(7), 93. Underwood, A. J. V. Chem. Eng. frog. 1948, 44, 603. Umeda. T.. Chlyoda Chemlcal Eng. and Construction Co., private communication, 1982. Wang, J. C.; Henke, 0.E. Hydrocarbon Process 1966, 45(8), 155. Westerberg, A. W.; Stephanopoulos, G. Chem. Eng. Sci. 1975 30, 963.
Received for review March 12, 1984 Accepted July 27, 1984
