938
Ind. Eng.
Chem. Process Des. Dev. 1986, 25, 936-941
poration for the us Department of Energy under Contract DE-AC04-76DP00053. Literature Cited Furfarl, S. Hydropyrolysis of Coal; IEA Coal Research: London, England, 1982; ICTISITRPO. Kansa. E. J.: Perlee. H. E. Combust. Flame. 1980. 38. 17. Mchrthy, J.; Ferral, J.; Charng, T.; House&n, J. Assessment of Advanced Coal OesMcatlbn RC4xW.98~NASAlJet ; F‘ropuklon Laboratory: Pasadena, CA, 1981; DOE/ET 13032-2. Nettleton, M. A.; Stirling. R. Combust. Flame, 1974 22, 407.
Park, W. K.; Mayer, R. L. T-ature-Tim-Reaction ~eovumcontrdled Hydfopyfo@k?; Monsanto Research Miamisburg OH, 1981; MLM-2849. Park, W. K.; Mayer, R. L. ulha Fast Rate h)&opymlysis of Coal: Final Report; Monsanto Research: Mlamlsburg, OH 1985; MLM-3120. Taiwalkar, A. T. A Toplcal Report on Coal Hydmpym/ys/s; Instltute of Gas Techno-: Chicago, IL 1983; DOE/MC/19316-1408 (DE83006592). Woodburn, E. T.; Everson, R . C.; Kirk, A. R. M. Fuel 1974, 53, 38.
Received for review September 17, 1984 Revised manuscript received September 18, 1985 Accepted March 28, 1986
Mutaceffect Extractive DfstiMertion for Separating Aqueous Azeotropes Scott Lynn* and Donald N. Hanson Department of Chemical Enginwring, University of California, Berkeley, Callfmnla 94 720
Operating two or more distillation columns in parallel, with the reboiler of one serving as the condenser of the next, is known to reduce energy consumption. Similarly, introducing the feed to an extractive distillation column as a vapor, rather than as a liquid, saves energy when the major feed component is light. Combining these two concepts offers an attractive method for separating light, azeotrope-farming Organic compounds from water. As examples, the energy consumption for separating 99.9 wt % ethanol from feeds of 6 or 10 wt % was estimated for distillation trains of four or five columns. The resub indicate that the process achieves steam consumptions of only 0.94-1.47 kg per kg of ethanol product, which corresponds to 2100 to 3300 kJ/kg of ethanol (6000-9400 Btu/gal).
Numerous organic compounds form aqueous azeotropes, and their separation from water is a commonly faced industrial problem. Many of these compounds, such as the lighter alcohols and ketones, have volatilities close to that of water. In dilute aqueous solution they have relatively large liquid-phase activity coefficients and are readily stripped from solution in the lower section of a distillation column. However, in more concentrated solutions the activity coefficient of water becomes large and the relative volatility of the two components consequently approaches unity. The overheat product of a simple distillation is hence limited to the composition of the azeotrope, and in many systems achieving even that degree of concentration requires a large number of theoretical stages and a high reflux because of the “tangent pinch” phenomenon. Subsequent treatment of the overhead product is needed to obtain the organic compound as a nearly anhydrous product, and that treatment is frequently expensive, both in capital and operating costs. An example of this type of problem is the separation of ethanol from water, which is currently of significant industrial importance because of ethanol’s potential as an octane enhancer in gasoline. The literature on this subject is correspondingly voluminous. The references cited here are, therefore, intended to give only an indication of the scope and variety of interest in the problem. Black (1980) reported the energy required to separate absolute alcohol from a feed containing 6.4 wt % ethanol by six conventional distillation methods. He found values ranging from 0.346 to 1.15 times the heating value of ethanol, 26600 kJ/kg (80000 Btu/gal). Douglas and Feinberg (1983) surveyed nondistillation ethanol separation processes. To put their survey into perspective in terms of economics and energy consumption, they reported that the most competitive conventional technology appeared to be an energyintegrated azeotropic distillation system using diethyl ether 0196-430518611125-0936$01.5010
that required 6000 kJ/kg (17000 Btu/gal) of ethanol product. Holland et al. (1981) discussed modern computing techniques for modeling azeotropic distillation. Leeper and Wankat (1982) report on the use of gasoline as a solvent for extracting ethanol from aqueous solutions near azeotropic concentration. Ladish et al. (1984) have used dry corn meal to dry wet ethanol vapors. Zudkevitch et al. (1984) patented the use of various hydrophobic phenols in an extractive distillation process for this separation. Lynd and Grethlein (1984) used vapor recompression to reduce energy consumption and potassium acetate as the extractive agent. Lee and Pahl (1985) screened various hydrophilic glycols as extractive agents. Example One. Process Evolution Figure 1 is an illustration of the use of conventional distillation followed by extractive distillation to separate a nearly pure, light organic compound from water. The example chosen is the separation of nearly pure alcohol from a feed containing 2.6 mol % (6.4 wt %) ethanol. Substantially all of the ethanol is stripped from the water in a conventional distillation column (column 2), which produces an overhead product of 80 mol % (91 wt %) ethanol. This intermediate product is then fed to a reboiled extractive distillation column (column 3), which produces an overhead product of 99.9+ % ethanol. Substantially all of the water contained in the feed to column 3 leaves in the bott~msstream, together with the extractive agent. In the regeneration column, column 4, this water is the overhead product and the extractive agent from the bottom is recycled to column 3. I t is instructive to review the calculations of the vapor flows in these columns in order to place the relative energy demands in perspective. The operation of the columns is approximated as equimolar overflow, with saturated vapor or liquid feeds, to simplify the calculations below. Less 0 1986 American Chemical Society
Ind. Eng. Chem. Process Des. Dev., Vol. 25, No. 4, 1986 937
3;
(xF)(1 - Yd)
sition, 10.4 for 2.6 mol % ethanol. With an overhead product composition of 80 mol %, (Lps,l/Fl),in = (Lm2!FZ)= 0.0990. The corresponding mmimum vapor flow in the column is eq 2.
Column 2
1
H20
Figure 1. Three-column sequence for separating ethanol from water.
i t & 0.4
0.2
1
0.4
1
1
0.6
I
1
0.8
I
1
I .o
x I (Liquid Mole Fraction Ethanol)
Figure 2. Operating diagram for the conventional distillation of ethanol and water.
than 10% error is introduced by these approximations, and more accurate calculations may readily be made by using computer techniques. The advantage of these simplified calculations is that they allow one to identify the most important heat loads in the process rather easily. Furthermore, they facilitate the evaluation of process modifications. (a) Conventional Distillation Columns. Figure 2 shows the vapor-liquid equilibrium curve for the ethanol-water system (data from Lewis and Carey (1932)), together with the operating lines for a column operating with a saturated liquid feed and the minimum vapor flow for this separation. The column is fed with dilute aqueous ethanol and operated to produce a product of intermediate concentration, 80 mol % ethanol. It is assumed that all of the light component in the saturated liquid feed to the column is taken overhead. The operating lines are the same regardless of whether the overhead product is a vapor, as in column 2 of Figure 1 (and of Figures 4 and 51, or a liquid, as in column 1 of Figures 4 and 5 (to be discussed below). With the assumption of complete recovery of the light component, the general expression for the minimum flow of liquid in the rectifying section becomes eq 1where aF is the separation factor at the feed compo-
For the present example, (Vl/Fl)- = (V2/F2)- = 0.1315. Equations 1 and 2 show how ( L m / n m hand (V/F),, vary with the purity of the overhead vapor. However, the two terms on the right-hand side of eq 2 tend to offset each other so that (V/F)- decreases by only 9% as Yd decreases from 0.80 to 0.22. At this point LRs is zero, Yd is in equilibrium with the feed, and the column is operating as a beer still. This means that for control purposes, the steam requirement of the column depends primarily on the feed rate. It is quite insensitive to the overhead composition (as long as the tangent pinch is avoided) and is even relatively insensitive to feed composition. The minimum vapor flow would, by definition, require a column with an infinite number of theoretical stages to accomplish the separation. In many cases it is found that the optimum column design corresponds to a vapor load about 10% greater than Vmin.For this example (V/OOpt = 0.1446 for columns with either a vapor or liquid overhead product. The vapor flow corresponds to the reboiler duty of the column. Taking the heat of vaporization for water to be 2250 kJ/kg (970 Btu/lb), this optimal vapor flow corresponds to a reboiler duty in the column of about 4930 kJ/kg (14000 Btu/gal) of ethanol in the net overhead product. (b) Extractive Distillation Column. For systems containing two components, A and B, at relatively low pressures, the separation factor (or relative volatility), a, is related to the activity coefficients, y, and the vapor pressures, polof the components by eq 3. am =
YAPAO -
YBPBO
(3)
To be effective in this proceas, an extractive agent must increase the relative volatility of either ethanol or water. Since the vapor pressure of pure ethanol is about 2.5 times that of pure cvater, there is an inherent advantage in using a hydrophilic extractive agent. This will reduce the activity coefficient of water and in some cases increase the activity coefficient of ethanol. The extractive agent should have a low volatility in order to facilitate its separation from ethanol and water but should not have an excessively high boiling point in order to avoid thermal degradation in the reboiler of the regeneration column. Other desirable properties of the extractive agent include low viscosity, toxicity, and cost. Ethylene and propylene glycols satisfy these criteria to varying extents. Trimble and Potts (1935) showed that solutions of water in ethylene glycol are nearly ideal. h a n u j a m and Laddha (1960) found that as little as 50 w t 7'0 glycol breaks the water-ethanol azeotrope. The relative volatilities of both water and ethanol are very large when compared to ethylene glycol, the extractive agent chosen from this example. The vapor pressure of ethanol at any given temperature is about 100 times that of ethylene glycol, whereas that of water ranges from 25 to 50 times greater. For this reason it is a good first approximation to neglect the glycol content of a vapor in equilibrium with a liquid containing ethylene glycol and
938
Ind. Eng. Chem. Process Des. Dev., Vol. 25, No. 4, 1986
X,
(Liquid Mole Fraction Ethanol, Extractant-Free Basis)
Figure 3. Operating diagram for the extractive distillation of ethanol and water.
a significant fraction of any combination of water and ethanol. Such a system can be treated computationally as pseudobinary by handling the compositions of both phases on a glycol-free basis. The pseudobinary equilibrium curve for a liquid phase containing 61.4 mol % glycol and varying ratios of water and ethanol (data of Hix (1982)) is shown in Figure 3. The feed to the extractive distillation column, column 3 in Figure 1, contains about 80 mol % ethanol. The solid operating lines on the operating diagram in Figure 3 correspond to a saturated vapor feed, whereas the dotted lines are for a saturated liquid feed. I t is seen that providing a vapor feed greatly reduces the vapor flow that is needed in the stripping section of the extraction column. This is the natural result of the fact that most of the feed must report in the overhead product of the extraction column. When the assumption of constant molal overflow is used, the mole fraction of glycol will be the same at all points in the column above the reboiler and may be set arbitrarily at 50 mol % . The equilibrium curve for this composition will not be greatly different from that in Figure 3. The separation factor, aF,of ethanol relative to water in the solution in equilibrium with this feed vapor is about 3. Assuming that the column effects a perfect separation between the ethanol and water in the feed, the minimum vapor flow in the rectifying section of the extractive distillation column, expressed in terms of vapor flow per mole of saturated vapor feed to the column, is eq 4. (4) By the Same arguments as before, the optimal vapor flow in the rectifying section would be about 10% greater than the minimum. For the case above (VRS,B/F3)opt
= 1*65
Below the feed tray the vapor flow is (VSS,3/F3)opt = 0.65 Since the vapor flow to the extractive distillation column is 0.0325 mol/mol of total liquid feed to the process in this example, the reboiler duty in this column corresponds to about 700 kJ/kg (2000 Btu/gal) of ethanol product from the process. If the feed were a saturated liquid, the vapor
flow required would be calculated from eq 1 and 2 above and would correspond to a reboiler duty of 1400 kJ/kg (4000 Btu/gal) of ethanol, which is (coincidentally) just twice the duty for a vapor feed. It should be noted that the ethanol reflux enters the extractive distillation column a short distance above the feed point for the glycol. This “knock-down” section ensures that the product ethanol will be free of glycol. ( c ) Regeneration Column. The liquid leaving the bottom of the extractive distillation column contains ethylene glycol, water, and a negligible amount of ethanol. The molar ratio of glycol/water in this stream is 0.85/0.20 or 4.25 for this example. The atmospheric boiling point of this stream is about 150 O C , and its flow is 0.0341 mol/mol of feed to the process. The glycol regeneration column (column 4 in Figure 3) may be operated at a reduced pressure to lower the reboiler temperature, which need not be much higher than the reboiler temperature of the extractive distillation column. Because of the high relative volatility between water and ethylene glycol, the minimum liquid flow needed in the rectifying section of this column is very small. The vapor flow needed for an optimum design will thus be about 10% greater than the water flow contained in the feed to this column, i.e., (V4/F4)opt = l.lXF,4 This vapor flow is about 0.007 mol/mol of process feed or about 250 kJ/kg (700 Btu/gal) of ethanol product. (a) Two-Effect Operation. The simplest extractive distillation process for separating a light azeotrope-forming organic compound from water would consist simply of the three columns in Figure 1. The energy requirement for driving the three reboilers would be about 5880 kJ/kg (16700 Btu/gal) of ethanol. Since the energy required for column 2 is about 84% of the total,it is clear that reducing this quantity must be accomplished if a significant overall reduction is to be achieved. One method of reducing energy consumption would be to use the heat released in condensation to produce boiling. This could be done by operating two conventional distillation columns in parallel, with one operated at a pressure of about one-third that of the other. This process configuration is shown in Figure 4, where column 1 would be the low-pressure column and the condensers of columns 2 and 3 would serve as reboilers. That is, reboiler R1would consist of condensers C2and C3. Note that there is only a fraction of a degree difference in the condensation temperatures of pure ethanol and 80% ethanol at the same pressure, but of course the two condensing streams must be kept separated. The dilute alcohol feed to the system, F , would be broken into two streams, z,F and zzF,to form the feeds to columns 1and 2, respectively. It is desirable to condense the overhead product of column 1 rather than to take it off as a vapor as is done with column 2. However, since it has essentially the same composition as the overhead product of column 2, it may be used as reflux in column 2. In this way the vapor stream leaving column 2 constitutes the net overhead product of both columns. This vapor stream is the feed to column 3 and is unchanged from the previous example. The energy consumption for this process configuration can be estimated as follows: One again makes the simplifying assumptions of constant molar overflows and negligible effect of temperature on the average heat of vaporization. The vapor flow in column 1will be the sum of the vapor flows generated in the reboilers of columns 2 and 3. As noted above, the vapor flow in column 3 will be unchanged, and the ratio of vapor to feed in each of
Ind. Eng. Chem. Process Des. Dev., Vol. 25, No. 4, 1986 939
columns 1 and 2 will be that calculated above, namely, 0.1446. Expressing all of this mathematically one may write v, = v, + v, Vi = 0.1446zIF V2 = O.144622F V3 = 0.65(~F/yd)F 21
+ 2, = 1
Solving these equations yields a value of z1of 0.573, the fraction of the total feed that flows to column 1. The resultant duty in column 2 and its reboiler is reduced to 43% of its former value, so that V, = 0.06171.: The total heat requirement for the process is now about 3050 kJ/kg (8700 Btu/gal) of ethanol. (e) Three-Effect Operation. The energy requirement for the process was cut almost in half by using two columns to do the conventional distillation step instead of only one. This savings is achieved because the extractive distillation and regeneration steps do not require much energy when the feed to the former is a vapor. It will be apparent that the addition of a third conventional distillation column, operated a t a pressure of about 3 times that of column 2 and with ita condenser used to reboil column 2, would reduce the energy requirement still further. This would be at the expense of a greater investment and a more complex operation. The flow configuration for this process modification is shown in Figure 5. The third conventional distillation column, column 5, is also operated to produce a liquid overhead- product that becomes part of the reflux for column 2. The vapor product leaving column 2 is thus the combined output of all three of the conventional columns. The vapor flow in column 2 is equal to that in column 5 , and the vapor flow in column 1 is the sum of those in columns 2 and 3. The corresponding mathematical equations are
v, = v, + v3 v, = v5 Vi = O.1446ZiF Vz = O.1446ZzF V5 = 0.1446~$
va = 0.65(&/ Yd)F (=0.0211F) 21 + 2 2 + 2 5 = 1 Solving these equations yields the result that 43% of the feed is sent to column 1 and 28.5% is sent to columns 2 and 5. The total duty for the three steam-driven reboilers is 2350 kJ/kg (6700 Btu/gal) of ethahol product. An alternative process configuration would be to take the vapor feed to the extraction column from column 5. This would reduce the overall energy consumption slightly but would require a significantly higher temperature in the reboiler of the extraction column, column 3. Not shown in Figure 5, in order to reduce the complexity of the drawing, is the method of recovering the sensible heat from the aqueous streams leaving columns 1, 2, and 5. This heat is a significant fraction of the total and cannot be neglected. One method of recovery would be to feed the water leaving column 5 into the stream leaving the reboiler in column 2. The hotter stream would flash, releasing its excess sensible heat by forming additional vapor to flow through column 2. The combined liquid streams
would similarly be fed to the reboiler of column 1. The total aqueous effluent would then be heat-exchanged against the total feed to the process, with the size of the heat exchanger being set by an economic balance against the value of the sensible heat of the effluent stream. Similarly not shown in Figures 1,4, or 5 is any cooling for the glycol stream leaving the regenerator, column 4, before it reenters the extraction column, column 3. In each of these configurations one would avoid feeding a hot stream to column 3 by suitable heat exchange with the feed to the highest pressure conventional column, which would be the feed to column 5 in Figure 5 . Example Two. Process Design A comparison of the column sizes and energy requirements for the process configurations of Figures 4 and 5 has been made for a specific application. The application is the production of 38 000 metric tons (84 000 000 lbs) per year of nearly pure ethanol. The assumed feedstocks are 6.0 and 10.0 wt % solutions of ethanol in water. The operating rate was arbitrarily taken to be 8000 h/ year. Provision for removing fusel oils or acetaldehyde was not included but could be accomplished as in conventional distillation processing. Ethylene glycol was again chosen as the extractive agent, although for making beveragegrade ethanol propylene glycol might be preferred for its low toxicity. Two process configurations were studied for each feedstock: a four-column sequence (Figure 4) and a fivecolumn sequence (Figure 5 ) . The pressure of plant steam was assumed to be 1.13 MPa (150 psig, 366 "F or 185 "C). Cooling water enters at 30 "C and exits at 45-50 "C. Feed enters at 40 "C and exchanges heat with the effluent water from the distillation train, which exits at 45 "C. The product purity was assumed to be 99.9 wt % ethanol, and the recovery of ethanol overhead was also taken to be 99.9%. The number of theoretical stages in each column was calculated by Underwood's method. The columns were assumed to have sieve trays operating at 65% efficiency at a pressure drop of 0.3 kPa (2.5 torr) per tray for vacuum columns and 0.6 kPa (5.0 torr) per tray for atmospheric and higher pressure columns. Packed columns could also be used. The heat exchange area was estimated by assuming a value of 2840 W/(m2 K) (500 Btu/(h ft2 O F ) ) for the average overall heat-transfer coefficient. The resulting equipment sizes and energy requirments, which represent a feasible although not an optimal design, are given in Table I. Column pressures are not listed because of space limitations but correspond to the temperatures in the respective reboilers and condensers. The operation of the columns was approximated as equimolar overflow to simplify the calculations. More accurate calculations can readily be made by using computer techniques. However, the conclusions that can be drawn from Table I are not expected to be affected by performing the calculations more accurately. For this example, the sensible heat of the effluent water and the heat exchange area required to heat the incoming feed with the effluent water are both significant items: the sensible heat loss is equivalent to 335 kJ/kg (950 Btu/gal) of ethanol product for the 6% feed and 200 kJ/kg (570 Btu/gal) for the 10% feed, while the heat exchange areas are 254 and 146 m2 (2730 and 1570 ft2),respectively. The use of plate-and-frame heat exchangers might allow reduction of both the heat losses and the exchanger areas. Additional Energy-Saving Modifications The heat content of the vapor leaving column 4 is not utilized in the flow sheets shown in Figures 4 and 5. This
940
Ind. Eng. Chem. Process Des. Dev., Vol. 25, No. 4, 1986
F
4
Figure 4. Four-column sequence for separating ethanol from water. Table I. Equipment Sizes and Energy Requirements for 38 000 metric tons (84 000 OOO lbs) of Annual Ethanol Production feedstock ( % ethanol) no. of columns column 1sizes trays diameter, m (ft) C1 area, m2 (ft2) C1 temp, "C Rl area R1temp, " C column 2 sizes trays diameter, m (ft) C2 area, m2 (ft2) C2 temp, O C R2 area, m2 (ft2) R2temp, O C column 3 trays diameter, m (ft) C3area, m2 (ft2) C3 temp, "C R3 area, m2 (ft2) R3 temp, "C column 4 trays diameter, m (ft) C4 area, m2 (ft2) C4 temp, "C R, area, m2 (ft2) R4temp, "C column 5 trays diameter, m (ft) C5 area, m2 (ft2) C6temp, "C R5area, m2 (ft2) R5 temp, "C total energy required kJ/kg of EtOH Btu/gal of ethanol (at 970 Btu/lb of steam)
Note that R, = C2 + C3.*Note that
6 4
6 5
10 4
10 5
50 2.3 (7.5) 74 (800) 60
50 2.1 (7.0) 92 (990) 50
34 2.1 (7.0) 63 (680) 60
34 2.0 (6.0) 79 (850) 50
a
a 85
a
a
88
85
50 1.2 (4.0) 14 (150) 94
34 1.2 (4.0) 22 (240) 101 16 (170) 130
34 1.1(3.5) 5 (50) 94
88 50 1.4 (4.5) 44 (480) 101 20 (220) 130
a
123
48 1.2 (4.0) 80 (860) 98 11 (120) 161
48 1.2 (4.0) 80 (860) 91
18 0.5 (1.6) 2 (20) 60 1 (10) 158
123 48 1.2 (4.0) 80 (860) 91
154
48 1.2 (4.0) 80 (860) 98 11 (120) 161
18 0.5 (1.6) 2 (20) 60 100) 158
18 0.5 (1.6) 2 (20) 60 1 (10) 158
18 0.5 (1.6) 2 (20) 60 1 (10) 158
b
50 0.9 (3.0) 71 (760) 132 71 (760) 158 3300 (9350)
a
2540 (7180)
b 154
34 1.1 (3.5) 49 (530) 132 26 (280) 163 2730 (7740)
2110 (5980)
= Cs.
vapor could be used to preheat the feed to columns 2 and 5 or, if the operating pressures were right, it could be sparged directly into the bottom of column 1 since it consists of essentially pure water. I t would then be nec-
essary to use an outside source of pure water to provide the reflux for column 4. Since the heat content of this vapor stream is small, its utilization may not be economical.
Ind. Eng. Chem. Process Des. Dev., Vol. 25, No. 4, 1986
Q4l
colurm I
Figure 5. Five-column sequence for separating ethanol from water.
It is unlikely that a fourth conventional column, operated in parallel with the other three, would be either economical or technically practical. In addition to the added process complexity, it would be necessary to take overhead vapor from two of the columns in order to obtain the feed needed for the extraction column. This requirement would reduce substantially the additional energy saving that could be realized.
Conclusions A major improvement in the energy efficiency of separating aqueous azeotropes results from the combinations of extractive distillation and multieffect evaporation that are shown in Figures 4 and 5. The advantage results from the combination of three not-uncommon techniques. First, the extractive distillation column is fed with a vapor instead of a liquid, thereby avoiding the revaporization of the major component of its feed. Second, the condensers of both a conventional distillation column and an extractive distillation column are used to drive the boil-up in a second conventional distillation column that operates at a substantially lower pressure. Finally, the overhead liquid product of one or two conventional columns is used to provide part of the reflux for a separate conventional column so that all of the net overhead product of the conventional distillation is available as a vapor at the desired pressure. Further advantages also derive from this system: First, because the intermediate product of the conventional distillation columns is not very close to the azeotropic composition, the problems of a tangent pinch are avoided and the number of plates in the conventional columns is low. Second, the energy of operation is essentially the same as that for making a product of near-azeotropic composition for compounds such as ethanol. The premium brought by the anhydrous product is thus obtained for little additional energy consumption. Third, this process may frequently be retrofitted to existing distillation planta, since the number and size of the columns required is no more than is required by more conventional systems for making an anhydrous product. Finally, the process is applicable to the separation of many organic compounds from water. Examples include the propanols, the butanols,
the light ketones, etc. Nonaqueous azeotropic systems, such as methanol-acetone, can also be separated by this technique if a suitable extractive agent can be found. From the molecular similarity of methanol to ethanol and water, ethylene glycol would be expected to be satisfactory.
Nomenclature F = feed to column, mol (liquid or vapor) L = liquid flow in column, mol po = vapor pressure V = vapor flow in column, mol x = liquid-phase mole fraction of ethanol y = vapor-phase mole fraction of ethanol ZN = fraction of total feed sent to column N Greek Letters a = separation factor or relative volatility y = activity coefficient Subscripts A, B = generalized binary components d = overhead product stream F = feed RS = rectifying (upper) section of column SS = stripping (lower) section of column 1, 2, ...,N = column designations Registry No. Ethanol, 64-17-5;ethylene glycol, 107-21-1. Literature Cited Black, C. Chem. Eng. Prog. 1980, 76(9), 78. Douglas,L.; Felnberg. D. EveluaMon of Nondlstllyetkw, Eth~nolSeparation Pro-
cesses; Solar Energy Research Institute: Golden, CO, July 1983; TR231-1887. Hlx, R. M. M.S. Thesis, Department of Chemical Engineering, Unhrerslty of California, Berkeley, 1982. Holland. C. D.; Gallen, S. E.; Lockett, M. J. Chem. Eng. 1981 (March 23), 185. Ladlsch. M. R.; Veloch, M.; Hong, J.; Blenkowskl, P.; Tsao, G. T. Ind. Eng. Chem. Process Des. D e v . 1984, 2 3 , 437. Lee. F.-M.; Pahl, R. H. I n d . Eng. Chem. Prcd. D e s . D e v . 1985, 24, 168. Leeper. S. A.; Wankat, P. C. Ind. Eng. Chem. Process Des. D e v . 1982, 21, 331. Lewis, W. K.: Carey, J. S. Ind. Eng. Chem. 1932, 24, 882. Lynd. L. R.; Grethleln. H. E. Chem. Eng. Prog. 1984, 80 (Nov), 59. Ramanujam. M.; Laddha, G. S. Chem. Eng. Scl. 1960, 12, 65. Trlmble, H. M.; Pons, W. Ind. Eng. Chem. 1935, 2 7 , 66. Zudkevitch. D.;Belsky, S. E.; Krauthelm, D. US Patent 4 428 798. 1984.
Receiued for review September 13,1985 Accepted April 29,1986
