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Ind. Eng. Chem. Res. 2000, 39, 1614-1623
Extractive Distillation Solvent Characterization and Shortcut Design Procedure for Methylcyclohexane-Toluene Mixtures Peter G. Tiverios† and Vincent Van Brunt* Department of Chemical Engineering, University of South Carolina, Columbia, South Carolina 29208, and O’Neal Engineering, 10 Falcon Crest Drive, Greenville, South Carolina 29607
A preliminary solvent screening tool for two-column extractive distillation flowsheets is described whose results can be used to initialize process simulators. The tool uses shortcut approximations to extractive distillation that accurately predict the relative column performance of alternative solvents with rigorous simulations using the Aspen Plus process simulator. The activity coefficient ratio of each feed component in the presence of the solvent, the boiling point difference between the high boiler and the solvent, and an a priori heat integration criterion characterize potential solvents. It was found that staging requirements for the flowsheets could be approximated using 4 times the number of stages predicted using the Fenske equation when it was applied over each column section. Using these approximations resulted in a shortcut method for evaluating and comparing alternative solvents and solvent families, predicting the staging requirements and quickly determining the feasibility of completely heat integrating a flowsheet. Background Doherty and co-workers4,6,7,8,10 have extended our understanding of azeotropic and extractive distillation. Most of their work is based on azeotropic distillation using geometric methods of nonlinear analysis incorporating residue curves. As described by Hoffman,1 the path of a residue curve is essentially the distillation path at total reflux, and accurate residue curves are generally constructed from laboratory vapor liquid equilibrium (VLE) data generated using a batch still. Doherty constructs theoretical residue curves by combining a VLE model such as UNIFAC with a differential material balance describing the dynamics of a still pot. Generally, if a separation is possible at total reflux, then the separation is physically possible and can be optimized by changing column parameters. Wahnschafft et al.2 used residue curves to graphically construct pinch point trajectories (e.g., Stichlmair et al.3) as separation limits for establishing the feasibility of ternary separations with various entrainers based on achievable bottoms and distillates. And most recently, Rooks et al.4 developed an equation-based approach rather than the graphical approach for determining possible distillation regions for homogeneous mixtures. Stichlmair and Fair5 summarize this work. Once the feasibility of a separation is determined, potential mass separating agents (MSA) can be evaluated further by considering energy requirements. Both minimum solvent flow rates and minimum reflux ratios can be used as indicators of a flowsheet’s energy requirements. Levy et al.6,7 used tangent pinch points to calculate minimum reflux ratios. Knapp et al.8 determined minimum entrainer flows using a geometric method involving bifurcations of the finite difference equations describing the middle section of the column. Koehler et al.9 provided an excellent overview of evaluation techniques for determining minimum energy * University of South Carolina. E-mail: 29208Vanbrunt@ engr.sc.edu. † O’Neal Engineering. E-mail:
[email protected].
requirements of distillation flowsheets from approximation techniques to rigorous solutions. The approximation techniques considered ranged from using Underwood’s method for both simple columns and extended for complex columns to pinch point criteria for nonideal and azeotropic mixtures (Koehler et al.,9 Levy et al.6,7). The rigorous solutions range from tangent pinch point criteria (Levy et al.,6 Fidkowski et al.10) to column partitioning (Poellmann et al.11). Introduction By recognizing that narrowing the range of suitable extractive distillation solvents using a shortcut procedure prior to rigorous simulation would save significant time and energy, the goal of this work was to develop a way to quickly screen alternative solvents for use in extractive distillation flowsheets prior to their evaluation using residue curve maps, rigorous simulation, or laboratory analysis. An additional goal was to explore ways of determining staging requirements and whether a heat-integrated flowsheet was possible. An aliphatic-aromatic split was chosen to compliment the earlier Gerster et al.12 work that addressed C3-C5 splits. Although a considerable amount of research has been done on the aliphatic-aromatic splits, thermodynamic modeling packages such as UNIFAC have not been adequately explored in the literature, and it was felt that a re-evaluation of the early work was warranted. In this work, the separation of toluene from methylcyclohexane (MCH), a frequently used pedagogical example of extractive distillation, is analyzed with the intent of characterizing potential extractive distillation solvents and developing a methodology for their systematic evaluation. A representative group of solvents (listed in Table 1) was selected to evaluate in a two column extractive distillation flowsheet (Figure 1). The solvent screening method and preliminary design procedure developed here is divided into the four steps described in the following sections. First, an initial screening of potential solvents begins by estimating their effectiveness by using UNIFAC to approximate the
10.1021/ie990654k CCC: $19.00 © 2000 American Chemical Society Published on Web 05/09/2000
Ind. Eng. Chem. Res., Vol. 39, No. 6, 2000 1615 Table 1 no.
solvent
classification
γI
γJ
1 2 3 4 5 6 7 8 9
phenol acetophenone methyl n-amyl ketone cyclohexanone o-cresol DMF propylene glycol furfural aniline
aromatic, alcohol aromatic ketone cyclic heterocycle, ketone aromatic, alcohol amide, ketone glycol heterocycle, aldehyde aromatic, amine
2.15 1.96 1.47 1.43 1.90 5.01 4.94 2.12 2.35
1.33 1.02 0.88 0.88 1.27 1.82 2.74 1.26 1.23
Figure 1.
nonideality of the liquid phase and by calculating the aliphatic-aromatic activity coefficient ratio in the presence of the solvent. Next, after a solvent’s separation effectiveness and its impact on the activity coefficient ratio of the solutes has been determined, the potential for solvent recovery and its separability from the higher boiling solute are evaluated by using the boiling point difference between the solvent and the higher boiling aromatic. Third, an approximation to the staging requirements is made, and last, the potential for heat integrating the flowsheet is determined. Preliminary Screening The Shortcut Method. Steps 1 and 2. Initial Solvent Screening. Characterizing solvents begins with analysis of the mixture. The phase equilibrium is described using eq 1, which assumes an ideal vapor phase and accounts for only the nonideality of the liquid phase.
yi γiPisat Ki ) ) xi P
(1)
The relative volatility, defined as the ratio of the equilibrium constants, describes the relative ease of separation of component i from component j and is approximated by eq 2.
Ki γiPisat ) Rij ) Kj γ P sat
(2)
j j
A mixture with a relative volatility close to 1.0 is extremely difficult to fractionate without the use of a mass separating agent; MCH and toluene have a relative volatility of 1.004 at 1.0 atm. A good extractive
γMCH/γtol ∆T 1.62 1.93 1.67 1.63 1.49 2.75 1.80 3.00 1.91
128 164 301 81 376 76 138 92 132
cascade I cascade II + III cascade IV + V RMT RTS RMS azeotrope 3.304 4.999 2.626 2.104 3.838 1.282 1.342 3.002 3.867
1.813 1.99 1.719 1.737 1.638 2.842 2.072 4.11 2.37
6.732 7.235 3.228 1.962 8.727 4.447 30.93 4.084 8.04
no no no no no yes yes yes no
solvent will increase the relative volatility by disproportionately modifying the respective liquid activity coefficients to increase the activity coefficient ratio (γi/γj). This can be achieved by increasing γi versus γj or decreasing γj versus γi. In the first case, γi > γj, the solvent preferentially rejects component i, increasing its effective vapor pressure and chemically enhancing its vaporization. In the second case, γj < γi, the solvent preferentially reduces the volatility of component j usually by enhanced secondary bonding. The effect of the solvent on the γMCH/γtol ratio is usually considered to be a measure of the enhanced interactions between the solvent and toluene predominately resulting from secondary or hydrogen bonding of the solvent with the p orbitals of the toluene and are a function of its active ligands. Solvent characterization specific to the MCH-toluene system was first investigated by Dunn et al.13 using laboratory vapor-liquid equilibrium (VLE) data. For initial screening, a mixture similar to the one used by Dunn was considered, i.e., 20 mol % toluene, 20 mol % MCH, and 60 mol % solvent at 234 °F. Figure 2 shows the effect of each solvent on the MCH and toluene activity coefficients. In region 1, species i (MCH) is preferentially rejected from the solvent whereas species j (toluene) is retained and its solubility in the solvent is enhanced. In region 3, the exact opposite effect takes place: species j is rejected and species i’s solubility is enhanced. In region 4, the solubility of both species is enhanced in the solvent (both activity coefficients are less than 1.0). Region 2 represents the region where the presence of the solvent rejects both species and enhances their volatility. It may be further divided into two subregions above and below the 45° line. The region above the line represents a preferential rejection of component i relative to j; however, the solvent is enhancing the volatility of both species. Below the line, component j is rejected relative to component j. Note that all of the solvents evaluated were either in region 1 or 2A, above the line; i.e., they increased the activity coefficient ratio. Of particular interest is that only two of the solvents lowered the activity coefficient of the toluene. This graph clearly shows that all the others increase the volatility of both species. The success of each extractive solvent is not only measured by its ability to increase the relative volatility of a close boiling mixture but also by its ability to be readily recovered from the bottoms stream. Solvents are also characterized using the boiling point difference between the solvent and the high boiler (toluene) as a measure of stripping column performance (the ability to separate the solvent from the high boiler) and the activity coefficient ratio of the low and high boilers (MCH and toluene) in the presence of the solvent as a measure of extractive distillation column performance
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Figure 2. Solvent characterization.
Figure 3. Methylcyclohexane-toluene separation γ ratio vs stripping factor.
(the ability to separate the low from high boiler in the presence of the solvent). The hydrocarbon chain length of these solvents has as great as or a greater impact on their potential as solvents as the active ligand. The active ligand affects the γ ratio, and the chain length affects the ∆T. In Figure 3, solvents are compared by plotting the γ ratio versus ∆T for each solvent in a flowsheet with the ordinate inversely proportional to the staging of the extraction column and the abscissa inversely proportional to he staging of the stripping column. The better solvents have the largest values of both coordinates. At the end of step 1, solvents that have both increased the relative volatility between the solutes and have a moderate boiling point differences between the high boiler and themselves have been identified. In step 2, these are further screened to identify potential processing problems by eliminating those solvents that form
azeotropes between the solvent and either solute. Azeotropic data references (Horsley,15 Gmehling et al.16) can be consulted, and azeotropes can be predicted by use of a thermodynamic model or determined in the laboratory. Although activity generation is a prime consideration when selecting a solvent so that separation of the aliphatic from the aromatic is achieved, separation of the solvent from the aromatic is equally important, and Figure 3 indicates these relative affects. Furfural has the highest impact on the γ ratio; however, it form azeotropes with MCH and toluene, respectively, rendering recovery and purification of the feed components impossible without increasing flowsheet complexity. Flowsheets for organics that form heterogeneous azeotropes using liquid-liquid extraction are an alternative, and membranes may be used as well (Seader and Henley17).
Ind. Eng. Chem. Res., Vol. 39, No. 6, 2000 1617 Table 2. Cascades cascade
description
function
I II
top section of extractive column middle section of extractive column
III
bottom section of extractive column
IV V
rectifying section of stripping column stripping section of stripping column
solvent recovery and aliphatic purity separation of aromatic and aliphatic enhanced by the solvent; toluene recovery separation of aromatic and aliphatic enhanced by the solvent; MCH recovery solvent recovery, aromatic purity aromatic recovery, solvent purity
Table 3. Extractive Distillation Column solute solv. distil. solv. staging recovery (mol/h) total feed feed rate rate stages stage stage casc. I casc. II casc. III (mol/h) (mol/h) toluene MCH solv. phenol acetophenone methyl n-amyl ketone cyclohexanone o-cresol DMF propylene glycol furfural aniline
reboiler duty (BTU/h)
conden. duty (BTU/h)
reflux ratio
100 74 98
70 42 61
31 12 38
31 12 38
39 30 23
30 32 37
51 50.25 50.5
330 330 330
49.99 49.65 49.96
49.61 328.6 6.12×106 6.57×106 49.71 329.8 4.75×106 5.12×106 49.79 329.3 6.88×106 6.90×106
86 99 39 43 88 60
50 79 21 33 58 42
32 8 3 4 18 26
32 8 3 4 18 26
18 71 18 29 40 16
36 20 18 10 30 18
50.5 50.1 68.75 58.25 615 52.5
330 330 330 330 330 330
49.55 49.75 49.99 49.99 49.69 49.99
49.67 49.57 49.32 49.66 49.48 49.64
329.6 329.7 310.6 321.4 318.3 327.2
6.47×106 6.32×106 2.85×106 3.63×106 3.64×106 4.44×106
6.44×106 6.98×106 3.60×106 5.29×106 3.92×106 5.02×106
12.0 9.4 13.0 12.0 13.2 3.9 7.0 5.0 8.5
Table 4. Stripping Column recovery staging
total feed reflux stages stage casc. IV casc. V ratio mol/h phenol acetophenone methyl n-amyl ketone cyclohexanone o-cresol DMF propylene glycol furfural aniline
toluene %
solvent
mol frac.
mol/h
conden. duty (BTU/h)
reboiler duty (BTU/h)
15 18 28
8 7 16
7 11 12
8 7 16
8.8 7.4 13.6
49.99 99.98 49.58 99.16 49.32 98.64
0.990 0.993 0.990
325.31 7.12×106 3.41×106 327.49 6.09×106 1.44×106 326.01 1.22×107 6.68×106
28 17 80 36 30 13
9 9 75 3 22 9
19 8 5 33 9 4
9 9 75 3 22 9
15 7.4 7 4.6 7 6.2
49.50 49.55 47.15 48.50 49.48 49.08
0.990 0.991 0.970 0.960 0.989 0.990
326.30 326.73 301.28 308.54 314.8 323.93
Staging Requirements and Heat Integration Shortcut Flowsheet Evaluation and Comparison with Simulation. In this section, a shortcut approximation for staging requirements and heat integration of extractive distillation flowsheets are developed. For the shortcut approximation, this flowsheet was broken into five distinct cascades using pseudobinary mixtures to approximate the behavior in each cascade. These cascades are shown in Table 2. The model feed, F, was taken to be a 50 mol/h MCH and a 50 mol/h toluene stream that was introduced as a saturated liquid into the extractive distillation column at an appropriate feed location and a 330 mol/h solvent, S, stream that was introduced at another arbitrary location above the feed stage. Both feed locations were later optimized. The more volatile MCH was removed as a liquid from the total overhead condenser of the extractive distillation column, and the toluene-solvent mixture was recovered from its bottoms. The bottoms was then fed to the stripping column where the toluene was removed as a liquid in a total overhead condenser and the solvent was recovered from the bottoms. The stripping column bottoms was recycled back to the extractive distillation column and used as the solvent feed with sufficient fresh makeup solvent introduced to replace the solvent loss in the two product streams. The simulation specifications were allowed to float whenever possible. Both columns were specified to
99.00 99.10 94.30 97.00 98.96 98.18
1.15×107 6.05×106 5.63×106 4.05×106 5.76×106 5.14×106
7.12×106 1.29×106 2.82×106 2.02×106 2.61×106 1.23×106
recov. col. overhead strip. col. conden. reboiler temp. (°F) temp. (°F) 412.3 412.6 412.5
369.9 405.4 314.2
412.7 412.6 378.8 374.4 403.8 411.7
323.7 405.9 312.6 370.7 332 372.2
operate at pressures sufficient to generate stream temperatures that would allow heat integration of the extractive distillation column overhead condenser with the stripping column reboiler. The stripping column was specified to operate at 14.7 psia (no vacuum systems required), and the extractive distillation column was specified to operate at 150 psia (standard materials of construction-low pressure). The flowsheet could have been configured by swapping column pressures and integrating the stripping column condenser with the extractive distillation column reboiler. However, the respective heat duty of the stripping column condenser was about half of the extractive distillation column reboiler duty, and this resulted in a flowsheet that could only be partially heat integrated. The configuration chosen reduced the temperature of the stripping column to minimize the possibility of solvent degradation. A minimum-approach temperature difference of 10 °F between the extractive distillation column overhead condenser and the stripping column reboiler was used to create designs that did not have excessive heat-transfer area. The staging, feed locations, reflux ratio and distillate rate were optimized until 99% of the MCH and toluene fed was recovered, and each product purity was 99 mol % as well. The simulation results are tabulated in Tables 3 and 4. Step 3A. Pseudobinary Approximation. The shortcut procedure analyzed the flowsheet cascade by cascade
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Figure 4. Recovery column section I simulated staging vs Fenske correlation.
and the relative volatility used in the Fenske equation was developed using approximations to the mixture at both ends of each respective cascade. Cascade I was considered separately from the lower two cascades in the extractive distillation tower. It separates the low boiler from the solvent and was analyzed using this pseudobinary mixture. The top of cascade I was considered an essentially pure low boiler (50 mol MCH) with the solvent present in infinite dilution. In terms of concentrations, the distillate was considered to have respective mole fractions of 0.99 and 0.01. The bottom of cascade I was assumed to be a binary mixture of the solvent and the low boiler and approximated using the relative flows fed to the column (50 mol MCH and 330 mol of solvent). The relative volatility used was the geometric average of the top and bottom streams of cascade I. Cascades II and III separate the low from the high boiler with the solvent being removed with the high boiler in the bottom’s stream. A pseudobinary was assumed to be MCH and toluene with the respective mole fraction ratios of 0.99/ 0.01 at the top and 0.01/0.99 at the bottom. The relative volatility was estimated using the mixed stream feed composition (XMCH ) 0.12, Xtol ) 0.12, Xsol ) 0.76) at its bubble point at the column pressure (150 psig). The stripping column, cascades IV and V, separates the high boiler from the solvent and was analyzed as a typical binary separation with mole fraction ratios of 0.99/0.01 at the top and 0.01/0.99 at the bottom. The relative volatility used was the geometric average of the feed and bottom streams. Similar to cascade I, the bottom stream relative volatility was calculated assuming infinite dilution of the high boiler in solvent. The feed stream properties for the stripping column were estimated at the extractive distillation column (cascade II and III) bottoms stream condition.
Step 3B. Flowsheet Staging Requirements. The staging requirements for each cascade from the simulation were predicted using the Fenske equation evaluated with the pseudobinarys:
N)
ln[(XLB/XHB)TOP/(XLB/XHB)BOT] ln R
(3)
The staging requirements used in the simulations and those predicted by the pseudobinary Fenske equation are plotted with respect to the R ratio in Figures 4-6. Individual solvents are referred to by number (see Table 1) in these graphs. The pseudobinary Fenske equation trends with the staging requirements predicted from the simulation reasonably well. However, to accurately predict staging and relative feed location, it had to be applied to each column section in the extractive distillation column as described above rather than by a single application over the entire column. It has been the standard practice to use twice the staging predicted by the Fenske equation when using shortcut methods. Although the Fenske equation is shown to trend with the simulation staging prediction, this work shows the staging to be 3-4 times the staging predicted by the Fenske equation rather than the accepted rule of thumb. An accurate forecast of staging requirements is essential for evaluating capital costs for competing solvents. As a consequence of this analysis, step 3B is to conservatively predict the staging requirements for the flowsheet column section by column section using 4 times the Fenske predictions, even though the trend was from 3 to 4 times the minimum. The specific concentrations for each cascade and the staging for the flowsheet is shown in Table 5.
Ind. Eng. Chem. Res., Vol. 39, No. 6, 2000 1619
Figure 5. Recovery column cascade II and III simulated staging vs Fenske correlation.
Figure 6. Stripping column: simulated staging vs Fenske correlation. Table 5. Compositions and Relative Volatilities Used in the Fenske Equation and the Relationship between Prediction vs Simulation pseudobinary cascade
XLB-TOP
XHB-TOP
XLB-BOT
XHB-BOT
R
LB
HB
Fenske correlation
I II + III IV + V
0.99 0.99 0.01
0.01 0.01 0.99
0.13 0.01 0.99
0.87 0.99 0.01
RM-S RM-T RT-S
MCH MCH toluene
solvent toluene solvent
Nsim ) 4NFenske Nsim ) 3.75NFenske Nsim ) 3NFenske
Step 3C. Flowsheet Heat Integration. The ability to heat integrate the two columns can result in a 50% reduction in operating costs. For the integration considered here between the extractive distillation condenser and the stripping column reboiler, an approxi-
mation to their duties is shown to result in a heat integration criterion to quickly evaluate whether this type of energy reduction is possible. For other configurations, a similar criterion can be developed and follows by analogy to what is presented (see Appendix A). The
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Figure 7. Heat integration criterion.
Figure 8. Heat integration simulated.
criterion can be written for the MCH-toluene system as
(
(
)
Rtol-sol - 1 RMCH-sol - 1
)(
fMCH fsol ∆HMCH >1 ftol ∆Hsol 1 + Rtol-sol fsol
1 + RMCH-sol
)
(4)
This criterion only requires estimation of the relative volatilities between the two solutes and the solvent, the latent heats of vaporization for the components present in the heat exchangers, and the relative feed flows of MCH, toluene, and solvent. Although it was derived for
the MCH-toluene system, the criterion is general for this heat integration scheme. In Figure 7, eq 4 is plotted on the ordinate, the solvent number is plotted on the abscissa, and the shortcut heat integration criterion shows a one to one correspondence with the rigorous simulation results. Thus, potential solvents can be further screened by rejecting those that do not lend themselves readily to heat integration, simply by evaluating eq 4. The extractive distillation column condenser duties and stripping column reboiler duties (generated from the Aspen Plus simulations) are shown in Figure 8 with respect to the temperature difference between these two streams for each of the nine solvents that underwent detailed design. Heat integration is not viable for every
Ind. Eng. Chem. Res., Vol. 39, No. 6, 2000 1621
Figure 9. Recovery column: solvent flow rate.
solvent due to either insufficient available energy from the condenser or an insufficient temperature difference between the condenser and the reboiler. Cyclohexanone has its stripping column reboiler duty greater than its recovery column condenser duty. Its negative energy difference indicates that at these conditions it could not be heat-integrated. Although propylene glycol, acetophenone, and o-cresol have sufficient energy to permit heat integration, the stream temperature difference is insufficient and would result in excessive stripping column reboiler surface area. This could be remedied by increasing the extractive distillation column operating pressure and subsequently elevating its overhead condenser temperature. However, excessively high temperatures could cause thermal degradation of the components and may be impractical for other reasons as well. Other Considerations. To evaluate whether the optimization of solvent flow would influence the fixed solvent flow results presented, the Aspen Plus simulator was used to determine the optimum flow for each of the nine representative solvents. The simulation specifications were relaxed to allow a wide range of solvent flows to be evaluated in the previously configured columns without further column optimization. The toluene product recovery and purity remained 99% in the stripping column; however, the MCH purity (solvent recovery) and the solvent recycle purity were permitted to be sloppy. Figure 9 shows the dependence of the simulated optimum solvent flows on the solute activity coefficient ratio for the extractive distillation columns that were previously optimized for staging, feed location, and reflux using a fixed solvent flow rate. These simulations reflect the strong dependence of optimized solvent flow on solute activity coefficient ratio. However, it was found, in general, that sloppy solvent splits increased overall energy costs due to the additional separations required to recover the solvent and should be discouraged. The solvent flow can be optimized without sacrificing the separation specifications, thereby reducing solvent inventory, operating and capital costs. Summary of the Screening Method. The screening method, outlined in Figure 10, is as follows:
Figure 10. Solvent selection.
Step 1. Preliminary evaluation of the solvent by plotting the γ ratio of solutes for each solvent (12%, 12% mixture of solutes with 76% solvent can be used) versus ∆T between the high boiler and the solvent.
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Step 2. Those solvents that form azeotropes with the solutes are rejected. Steps 3A and 3B. Staging is evaluated using pseudobinary mixtures with compositions shown in Table 5 and the staging requirements of 4 times those predicted by the Fenske equation individually applied to each cascade. Step 3C. The potential for heat integration of each flowsheet is evaluated using the criterion in eq 4. Those that do not meet the criteria are rejected. Step 4. Proceed to rigorous evaluation. The results of the preliminary screening can be used to initialize process simulators for more detailed analysis and optimization of the remaining solvents. Conclusions A preliminary solvent screening tool for extractive distillation has been developed whose results can be used to initialize process simulators. UNIFAC provided a reliable tool to determine the relative behavior of solutes and predict their activity coefficient ratio in the presence of a solvent. This along with the temperature difference between the high boiler and the solvent enabled a graphical prescreening of potential solvents. Both the hydrocarbon chain length of the solvent and its active ligand end groups have significant effects on a solvent’s influence on the solutes activity coefficient ratio. Although prescreening using the solute γ ratio is an indicator of a solvent’s effectiveness, shortcut approximations to extractive distillation have been shown to accurately predict the relative column performance when compared to the rigorous simulations using Aspen Plus. The staging requirements in extractive distillation flowsheets are approximately four times that obtained using the Fenske equation applied over each appropriate column section. The pseudobinary Fenske equation was valid when applied over cascade I, cascades II and III together, and cascades IV and V together because each of these cascade groups essentially behaved as binary separations. Heat integration of the extractive distillation overhead condenser with the stripping column reboiler was investigated, and an a priori heat integration criterion using predicted relative volatilities was found to trend and predict relatively rigorous simulation results. Heat integration of the stripping column overhead condenser with the extractive distillation column reboiler was investigated but found insufficient for total heat integration for the MCH-toluene system. In addition, solvent requirements were shown to be highly dependent on the activity coefficient ratio of the solutes in the presence of the solvent. In general, using shortcut approximations to prescreen extractive distillation solvents using physical properties, predicted γ ratios using UNIFAC, and predicted relative volatilities was found to be expedient for both solvent screening and as an initiation tool prior to rigorous simulation. On the basis of the methodology described above, the nine solvents were evaluated. Azeotrope formation removed furfural, DMF, and propylene glycol from consideration, and the inability to heat integrate the flowsheet removed cyclohexanone from consideration. Of the remaining solvents, acetophenone and then phenol appear to be the most favorable solvents when considering solvent flow require-
ments, overall staging requirements, and recovery of all components. However, the acetophenone flowsheet would both require raising the operating pressure of the recovery column to elevate the overhead condenser temperature and increase the differential temperature between the two columns to optimize equipment size. Appendix A Developing a Shortcut Heat Integration Criterion. Cascade I in the extractive distillation column is essentially a binary separation between the aliphatic low boiler and the solvent and simultaneous solution of the solvent feed stage q-line, the equilibrium expression for a binary, and the cascade I operating line results in an expression for the minimum reflux. The equation for the cascade I operating line equation is yi ) R/(R + 1)xi + xdi/(R + 1), and the q-line equation for the solvent feed is yi ) q/(q - 1)xi + zf/(q - 1).
RmzF + qxD
) Rm(1 - zF) + q(1 - xD) R[xD(q - 1) + zF(Rm + 1)] (Rm + 1)(1 - zF) + (q - 1)(1 - xD)
(A1)
Here R is the relative volatility of the low boiler (MCH) with respect to the solvent, zF is the feed mole fraction of MCH, and q is the traditional liquid quality of the feed stream and is defined as
q)
HG - HF HG - HL
(A2)
Equation A1 can be rigorously solved for Rm if the feed quality and distillate fraction are known. In the case of a saturated liquid feed at the bubble point (q ) 1), eq A1 can be solved for the minimum reflux ratio in terms of the separation factor, the feed composition, and the distillate composition.
Rm )
[
]
1 - xD 1 xD -R R - 1 xF 1 - xF
(A3)
Further approximation of the distillate by a pure lowboiling stream (xD ) 1) results in the simplified relationship between the minimum reflux ratio, the effective separation factor, and the feed composition.
Rm )
1 xF(R - 1)
(A4)
Energy and material balances for the top cascade relates the condenser duty to the reflux ratio:
HVV ) hdD + hdL + Qc
(A5)
V)L+D
(A6)
Qc ) (R + 1)D∆HMCH
(A7)
Substitution of the reflux approximation, effective feed flows for the feed, and relative volatility from eq A4 into eq A7 results in an approximation for the condenser duty in terms of predictable and known quantities.
Ind. Eng. Chem. Res., Vol. 39, No. 6, 2000 1623
Qc )
[
1 + (RMCH-sol - 1)[mMCH/(mMCH + msol)]
]
1 mMCH∆HMCH (A8) Here, m refers to the molar flow of a component in the feed. Similar to the analysis of the condenser of the extractive distillation column, the stripping column reboiler duty was related to toluene and solvent flow. Assuming a saturated liquid feed and minimal heat effects in the stripper (constant molar overflow) results in its rectifying section vapor flow rate equaling the stripping section rate. Similar to the cascade I analysis above, a mass balance around the rectifying section of the stripping column results in an expression for the vapor flow in terms of the reflux ratio and distillate flow rate. The reboiler duty can be approximated by
Qr )
[
]
1 + 1 mtol∆Hsol (Rtol-sol - 1)[mtol/(mtol + msol)] (A9)
Heat integration for this scheme is possible only if the extractive distillation column condenser duty is greater than the stripping column reboiler duty.
Qc-r > Qr-strip or Qc-r/Qr-strip > 1
(A10)
Reducing eq A10 results in a compact criterion requirement for heat integration between the extractive distillation column condenser and the stripping column reboiler.
( )(
fL 1 + RL sol RH sol - 1 fsol ∆HL >1 RL sol - 1 fH ∆Hsol 1 + RH sol fsol
(
Notation
)
)
B ) column bottoms flow (mol/h) D ) column distillate flow (mol/h) fMCH ) feed rate of methylcyclohexane (mol/h) ftol ) feed rate of toluene (mol/h) fsol ) feed rate of solvent (mol/h) m ) flow rate (mol/h) F ) feed HG ) enthalpy of the saturated vapor feed stream HL ) enthalpy of the saturated liquid feed stream HF ) enthalpy of the actual feed stream HV ) enthalpy of vapor stream V hd ) enthalpy of liquid distillate stream L or D hL ) enthalpy of bottoms stream L or B Ki ) equilibrium coefficient L ) column liquid flow (mol/h) Pisat ) component vapor pressure P ) total pressure ∆Hvap ) heat of vaporization (HV - hd) ∆Hvap ) heat of vaporization (HV - hL) ∆HL ) heat of vaporization (HV - hL) of low boiler ∆Hsol ) heat of vaporization (HV - hL) of solvent Qc ) condenser duty Qr ) reboiler duty
(A11)
R ) reflux ratio (L/D) S ) solvent V ) column vapor flow (mol/h) XLB ) low boiler mole fraction in pseudobinary XHB ) high boiler mole fraction in pseudobinary yi, xi ) vapor and liquid mole fractions, respectively R ) relative volatility of LB/HB RTS ) relative volatility of the toluene-solvent in the stripping column RMS ) relative volatility of the MCH-solvent in section I of the extractive distillation column. γi ) liquid activity coefficient φi ) vapor fugacity in mixture; assume ideal vapor phase: φi ) 1
Literature Cited (1) Hoffman, E. J. Azeotropic and Extractive Distillation; WileyInterscience: New York, 1964. (2) Wahnschafft, O. M.; Koehler, J. W.; Blass, E.; Westerber, A. W. The Product Composition Regions of Single-Feed Azeotropic Distillation. Ind. Eng. Chem. Res. 1992, 31, 2345-2362. (3) Stichlmair, J.; Fair, J. R.; Bravo, J. L. Separation of Azeotropic Mixtures via Enhanced Distillation. Chem. Eng. Prog. 1989, 85, 63-69. (4) Rooks, R. E.; Julka, V.; Doherty, M.; Malone, M. F. Structure of Distillation Regions for Multicomponent Azeotropic Mixtures. AIChE J. 1998, 44. (5) Stichlmair, J.; Fair, J. R. Distillation: Principles and Practices; Wiley-VCH: New York, 1999. (6) Levy, S. G.; Doherty, M. F. A Simple Method for Calculating Tangent Pinch Points in Multicomponent Nonideal Mixtures by Bifuracation Theory. Chem. Eng. Sci. 1986, 41, 3155-3160. (7) Levy, S. G.; Van Dongen, D. B.; Doherty, M. F. Design and Synthesis of Homogeneous Azeotropic Distillations: 2. Minimum Reflux Calculations for Nonideal and Azeotropic Columns. Ind. Eng. Chem. Fundam. 1985, 24, 463. (8) Knapp, J. P.; Doherty, M. F. Minimum Entrainer Flows for Extractive Distillation: a Bifurcation Theoretical Approach. AIChE J. 1994, 40, 243-268. (9) Koehler, J.; Poellmann, P.; Blass, E. A Review on Minimum Calculations for Ideal and Nonideal Distillations. Ind. Eng. Chem. Res. 1995, 34, 1003-1020. (10) Fidkowski, Z. T.; Malone, M. F.; Doherty, M. F. Nonideal Multicomponent Distillation: Use of Bifurcation Theory for Design. AIChE J. 1991, 37, 1761-1779. (11) Poellmann, P.; Glanz, S.; Blass, E. Calculating Minimum Reflux of Nonideal Multicomponent Distillation Using Eigenvalue Theory. Comput. Chem. Eng. 1994, 18, S49-S53. (12) Gerster, J. A. Azeotropic and Extractive Distillation. Chem. Eng. Prog. 1969, 65, 43-46. (13) Dunn, C. L.; Millar, R. W.; Pierotti, G. J.; Shiras, R. N. Toluene Recovery by Extractive Distillation. Trans. A.I.Ch.E. 1945, 41, 631-644. (14) Horsley, L. H. Azeotropic Data-III, 1st ed.; American Chemical Society: Washington, DC, 1973. (15) Gmehling, J.; Menke, J.; Krafczyk, J.; Fisher, K. Azeotropic Data; VCH: Germany, 1994. (16) Seader, J. D.; Henley, E. J. Separation Process Principles; Wiley: New York, 1998. (17) Flick, E. W. Industrial Solvents Handbook, 3rd ed.; Noyes Data Corporation: New Jersey, 1985. (18) Henley, E. J.; Seader, J. D. Equilibrium-Stage Separation Operations in Chemical Engineering, 1st ed.; Wiley & Sons: New York, 1981.
Received for review September 1, 1999 Revised manuscript received January 27, 2000 Accepted January 29, 2000 IE990654K
