Article pubs.acs.org/IECR
Methanol/Trimethoxysilane Azeotrope Separation Using PressureSwing Distillation William L. Luyben* Department of Chemical Engineering, Lehigh University, Bethlehem, Pennsylvania 18015, United States ABSTRACT: The process to produce trimethoxysilane involving the reaction of methanol with silicon produces a mixture of unreacted methanol and trimethoxysilane that requires separation so that the methanol can be recycled back to the reaction section. This binary mixture forms a maximum-boiling homogeneous azeotrope of 28.65 mol % methanol at 1 bar and 87.94 °C. The composition of the azeotrope changes fairly significantly with pressure, so separation into high-purity component streams is viable using pressure-swing distillation. The purpose of this paper is to design a pressure -swing distillation process for this separation that uses heat integration of the two columns, which operate at different pressures and different temperatures. One of the unique features of this system is the use of a vacuum column in which tray pressure drop has a significant effect on the optimum selection of design parameters such as the number of trays. The specifications for the bottoms compositions must be adjusted as pressures are changed so as to not get too close to or too far away from the azeotropic compositions. The economic optimum design has two columns operating at 7 and 0.25 bar with an auxiliary reboiler used on the low-pressure column. Feed composition strongly impacts the economics but not the basic process topology and operating conditions.
1. INTRODUCTION The separation of azeotropic mixtures is one of the most challenging and intriguing problems in chemical engineering. Many papers and books have discussed various techniques to overcome the natural distillation boundaries imposed by the presence of azeotropes. The book by Stichlmair and Fair1 summarizes a variety of approaches to achieve separations. Doherty and Malone2 discuss many of the design issues and emphasize the use of ternary residue curve maps for gaining physical insight into how multicolumn systems can be developed that achieve separations in these systems. Luyben and Chien3 explore both the design and the control of several types of azeotropic separation systems. Azeotropes occur in several different flavors. They can be classified in terms of the number of components (binary and ternary), the number of phases (homogeneous and heterogeneous), and the boiling temperature (minimum and maximum boiling). Since there are several types of azeotropes, there are a variety of separation methods. The four basic methods are azeotropic (using a light entrainer), extractive (using a heavy solvent), pressure-swing (using two columns as different pressures), and decanter/strippers (for binary heterogeneous azeotropes). The first two methods involve addition of a third component, which can raise issues of contamination of the products. The last two methods do not require adding different components but depend on the inherent physical properties of the chemical components and their temperature dependence. Pressure-swing distillation can be very effectively used if the composition of the azeotrope changes significantly with pressure. It is the technique of choice when it can be used because it has two distinct advantages. First, no other component needs to be added. Second, the columns can be heat integrated (use overhead vapor from the high-pressure, high-temperature column as a heat source in the reboiler of the low-pressure, low-temperature column). This feature can significantly reduce the energy requirements of the separation. © 2014 American Chemical Society
Pressure-swing distillation systems have been widely applied in industry and studied in academia for many years. One of the early studies,4 which considered issues of both design and control, was presented almost three decades ago. Examples of other applications include acetone/methanol,5 acetone/chloroform,6 and methylal/methanol.7 The use of pressure-swing distillation for separating minimum-boiling azeotropes was compared with separating maximum-boiling azeotropes using hypothetical components to access the inherent differences.8 It appears that the use of pressure-swing distillation for the separation of the methanol−trimethoxysilane azeotrope has not been reported in the literature. A process for the production of trimethoxysilane (C3H10O3Si, CAS No. 2487-90-30) was been described in a 1989 patent by Childress and Ritscher.9 It involves three reactors in series with counter-current flow of a solvent/silicon slurry with a vapor methanol stream. An excess of methanol is fed, so it must be recovered from the reactor effluent and recycled. The binary system of methanol and trimethoxysilane forms a maximum-boiling homogeneous azeotrope of 28.65 mol % methanol at 1 bar and 87.94 °C. Fortunately the composition changes significantly with pressure. Trimethoxysilane is used as a coupling agent, which is a compound that provides a chemical bond between two dissimilar materials, usually an inorganic and an organic. Organo-silanes are well-suited in these applications because of the ability to incorporate an organic-compatible functionality and an inorganic-compatible functionality within the same molecule. Tian et al.10 discuss a typical application. Received: Revised: Accepted: Published: 5590
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Figure 1. Txy diagram for methanol/trimethoxysilane at 1 bar.
Figure 2. Pressure dependence of methanol/trimethoxysilane azeotrope.
2. PHASE EQUILIBRIUM CONSIDERATIONS
maximum-boiling homogeneous azeotrope is predicted at 1 bar to have a composition of 28.65 mol % methanol and a temperature of 87.94 °C. Note that the temperature of the azeotrope is higher than the boiling point temperatures of either of the two pure components, indicating a maximum-
The normal boiling points of methanol and trimethoxysilane, as given in the Aspen APV 80.PUR Databank, are 64.7 and 84.45 °C, respectively. Using UNIFAC physical properties, a 5591
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Figure 3. Vapor pressure of trimethoxysilane.
Figure 4. Heat-integrated pressure-swing flowsheet for methanol/trimethoxysilane.
boiling azeotrope. Figure 1 gives a Txy diagram for the system at 1 bar. Figure 2 shows how the composition of the azeotrope changes with pressure. The composition can be shifted by about 10 mol % by adjusting the pressure over a range from vacuum conditions to 8 bar. Pressures higher than about 7 bar have little effect on the azeotropic composition. Higher pressures will result in higher reboiler temperatures, which would require more expensive heat sources in the reboiler of
the high-pressure column. So there are high-pressure considerations. As pressure is lowered, the required temperature in the reflux drum of the low-pressure column decreases. In order to be able to use cooling water in its condenser, the temperature cannot be lower than about 45 °C. Figure 3 shows that the vapor pressure of trimethoxysilane, which goes overhead in the lowpressure column, is about 0.25 bar at 45 °C. Therefore the operating pressure in the low-pressure column is set at 0.25 bar. 5592
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Figure 5. xy diagram for methanol/trimethoxysilane at 0.25 and 7 bar.
A word of caution is in order at this point. As mentioned above, Aspen UNIFAC physical properties have been used in this simulation study. No experimental VLE data has been found in the literature to confirm the UNIFAC predictions of the pressure dependence of the maximum-boiling azeotrope. The accuracy of the UNIFAC group-contribution method with a silicon atom in the molecule is unknown.
The low-pressure column is operated at as low a pressure as possible and still uses inexpensive cooling water in its condenser. For a 99 mol % trimethoxysilane distillate product, a reflux-drum temperature of 46 °C occurs at a pressure 0.25 bar. At this pressure, the composition of the azeotrope is 33.22 mol % methanol. This distillation boundary prevents producing bottoms from the low-pressure column that is any higher than a composition slightly lower than the azeotropic composition. Therefore the composition of the bottoms product from the low-pressure column must be slightly lower than the azeotropic composition at the appropriate pressure. It is important to note that this column operates under vacuum conditions. Tray pressure drop can be very significant in vacuum columns. The absolute pressure in the bottoms where the azeotrope occurs is typically significantly larger than the pressure in the reflux drum. Therefore the composition of the azeotrope should be considered at the base pressure, not the reflux-drum pressure. In the flowsheet shown in Figure 4, the reflux-drum pressure is 0.25 bar (azeotrope at 33.22 mol % methanol). But the base pressure of the 31-stage column (using a tray pressure drop of 0.04 psi per tray) is 0.333 bar, which corresponds to an azeotropic composition of 32.35 mol % methanol). The distillation boundary is lower at 0.333 bar than at 0.25 bar, so the specified bottoms composition must consider this difference. The bottoms composition is specified to be 1 mol % lower than the azeotropic composition at 0.333 bar. As
3. DESIGN OF LOW-PRESSURE COLUMN With a homogeneous maximum-boiling azeotrope, the pressure-swing distillation system produces high-purity products in the overhead distillate streams from the two columns. This is the inverse of the configuration to separate minimumboiling azeotropes (the bottoms streams are the high-purity products).The bottoms stream from each column is fed to the other column. Figure 4 shows the flowsheet, which includes heat integration (overhead vapor from the high-pressure column provides heat to a reboiler in the low-pressure column). High-purity methanol is the distillate product from the high-pressure column. High-purity trimethoxysilane is the distillate product from the low-pressure column. The feed flow rate is 100 kmol/ h of a binary mixture with a composition of 50 mol % methanol and 50 mol % trimethoxysilane. Product purities are specified to be 99 mol % for both distillate product streams. 5593
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Figure 6. Effect of pressure in HPC.
high-pressure column means a larger differential temperature driving force in the condenser/reboiler, which reduces the required size and capital cost. Another important consideration is how the required heat duty in the reboiler of the low-pressure column matches the required heat duty in the condenser of the high-pressure column. If these duties do not match perfectly (“neat” operation), an auxiliary reboiler in the low-pressure column or an auxiliary condenser in the high-pressure column will be required. For the case shown in Figure 4, the total required reboiler duty in the low-pressure column is 2.067 MW. The condenser duty in the high-pressure column is only 1.579 MW. So an auxiliary reboiler is required that adds an additional 0.488 MW. The heat source in this auxiliary reboiler is low-pressure steam (160 °C) since the temperature at the base of the low-pressure column is quite low (56.6 °C). The duty of the reboiler in the high-pressure column is 2.797 MW, and medium-pressure steam (184 °C) must be used since the base temperature is 166 °C at the base pressure of 7.08 bar. Therefore, the steam energy cost for the heat integrated systems has two components. The 2.797 MW in the highpressure column is assumed to cost $8.22 per GJ (Turton et al.11). The 0.488 MW in the auxiliary reboiler in the lowpressure column is assumed to cost $7.78 per GJ. As the pressure in the high-pressure column is varied to find the optimum, the heat duties in the four heat exchangers vary. Figure 6 gives quantitative results with the number of stages in both columns fixed at 31 and the pressure in the reflux drum of the low-pressure column fixed at 0.25 bar. As pressure increases, all of the heat duties decrease: reboiler duty in the high-pressure column QRHP, reboiler duty in the low-pressure column QRLP, and condenser duty in the high-pressure column QCHP. The difference between QCHP and QRLP is the heat duty of the auxiliary reboiler Qaux that reaches a minimum at 6 bar. As the pressure in the high-pressure column increases, the azeotropic composition decreases. To maintain the same 1 mol
shown in Figure 4, the bottoms composition is 31.35 mol % methanol. The heuristic of using a 1 mol % differential in composition between the limiting azeotropic composition and the specified bottoms composition is justified in a later section of this paper. The important implication of this strong effect on base pressure in a vacuum column is that changing the number of trays will affect the attainable bottoms composition. Adding more trays in the low-pressure column increases base pressure and lowers bottoms composition. This reduces the difference between the two bottoms compositions in the two columns and results in more recycle between the two columns (both bottoms flow rates increase as the difference between the two bottoms compositions decrease). Theoretical stages are used in the Aspen simulations, but packed columns might be used in practice if there are no fouling problems. The pressure drop used in the stage model is fairly low (0.04 psi per stage) and is typical of packing pressure drop per height of a theoretical tray.
4. DESIGN OF HIGH-PRESSURE COLUMN The selection of the optimum pressure for the high-pressure column involves an engineering trade-off between ease of separation and temperature level (and cost) of the energy source in its reboiler. Increasing pressure makes the separation in the high-pressure column easier, as shown in Figure 5. The equilibrium curve moves further away from the 45° line as pressure increases. The fatter curve indicates an easier separation. However, higher pressures mean higher base temperatures and more expensive energy sources. A second important issue is how pressure affects heat integration between the two columns. The high-temperature vapor from the top of the high-pressure column is condensed in a heat exchanger that serves as a reboiler for the low-pressure column. The base temperature in the low-pressure column is set by its composition and pressure. A higher pressure in the 5594
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% difference between the bottoms specification and the azeotropic composition, the bottoms specification is adjusted at each pressure. Table 1 gives numerical values of the design
Table 2. Cases with Various Numbers of Stages NTHP/NTLP
Table 1. Effect of Pressure in HPC P (bar)
yaz (mf MeOH)
xB,HP (mf MeOH)
QRHP (MW)
QCHP (MW)
QRLP (MW)
4 5 6 7 8
0.2332 0.2245 0.2177 0.2124 0.2082
0.2432 0.2345 0.2277 0.2224 0.2182
3.434 3.138 2.942 2.797 2.689
2.309 1.988 1.760 1.580 1.433
2.719 2.414 2.121 2.067 1.961
HPC ID (m) QR (MW) QC (MW) LPC ID (m) QR (MW) QC (MW) Qaux (MW) Recycle (kmol) Total Capital (106 $) Total Enegry (106 $/y) TAC (106 $/y)
parameters, the resulting heat duties, and the total cost of energy (medium-pressure steam in the high-pressure column reboiler and low-pressure steam in the auxiliary reboiler in the low-pressure column). As shown in Figure 6, the minimum energy cost occurs with the high-pressure column operating at 7 bar. It should be noted that high-pressure steam (254 °C at $9.88 per GJ) has to be used for pressures above 7 bar because the base temperature increases and the differential temperature driving force drops below the heuristic value of 30 °C. For example, at 8 bar the base temperature is 173 °C while the temperature of medium-pressure steam is only 184 °C.
a
21/31
31/31
41/31
1.594
1.316
1.242
1.304
1.253
4.064 2.789 2.352
2.797 1.579 2.353
2.494 1.274 2.354
2.753 1.544 2.768
2.534 1.306 2.298
2.067 2.728 0.488 116.6
2.067 3.175 0.793 116.6
3.233 4.328 1.689 113.2
1.916 3.032 0.610 119.7
1.582
1.508
1.577
1.577
1.731
1.615
0.8448
0.9411
1.128
1.080
2.142
1.348
1.367
1.654
1.658
2.067 3.170 0.772a 116.6
31/21
41/41
Auxiliary condenser on HPC since QRLP is less than QCHP.
the LPC. Total capital cost increases despite the shorter LPC because its diameter increases and the areas of its heat exchangers increase. 4. Adding more stages in both columns slightly reduces the heat duties in the HPC reboiler and condenser and in the LPC condenser. However, the lower bottoms specification in the LPC results in a larger recycle flow rate and higher duty in the auxiliary reboiler. Capital cost increases because of the taller columns. These results demonstrate that the base case design with 31 stages in each column is close to the economic optimum in terms of minimizing TAC.
5. NUMBER OF STAGES All of the cases considered above used columns with 31-stages. In this section we justify this selection by looking at the total annual cost (TAC) of columns with various numbers of stages. The pressures in the reflux drums of the high- and low-pressure columns are fixed at 7 and 0.25 bar, respectively. Of course the pressure in the base of the low-pressure column changes as the number of trays is changed, so the specification of its bottoms is adjusted appropriately. Increasing the number of stages improves the separating capacity of the column, so the reflux ratio and the energy consumption are reduced. But capital cost increases as the height of the column increases. There is also interaction between the two columns, so the TAC of the entire system must be considered. Table 2 summarizes equipment sizes and economic results (capital investment and energy costs) for a number of cases in which the numbers of stages are varied in both columns. Several trends can be seen: 1. Decreasing the number of stages in the HPC from 31 to 21 increases its reboiler duty as expected. However, its condenser duty increases and is larger than the required reboiler duty in the LPC. Therefore an auxiliary condenser is needed in the HPC instead of an auxiliary reboiler in the LPC. Energy cost increases because of the large increase in QRHP. 2. Increasing the number of stages in the HPC from 31 to 41 decreases its reboiler and condenser duties. The heat duty in the auxiliary reboiler in the LPC increases by almost a factor of 2 from the base case (31/31). The net effect is a slight increase in total energy costs. 3. Reducing the number of stages in the LPC permits a higher bottoms specification because of the lower total tray pressure drop. This reduces the recycle flow rate, which reduces the load in the HPC (lower reboiler and condenser duties). However, there are large increases in required duties in the condenser and auxiliary reboiler in
6. COMPOSITION DIFFERENCE BETWEEN xB AND yaz In all the cases considered in the previous sections, the specified bottoms compositions were assumed to be 1 mol % away from the corresponding azeotropic composition at the given pressure. For example, in the base case (Figure 4), the bottoms composition in the high-pressure column (22.24 mol % methanol) was set 1 mol % higher than the azeotrope at 7 bar (21.24 mol %). The bottoms composition in the lowpressure column (31.35 mol % methanol) was set 1 mol % lower than the azeotrope at 0.333 bar (32.35 mol %). We test this heuristic assumption in this section by evaluating the economics using other composition differences (Δxspec). The numbers of stages and pressures are those used in the base case. Table 3 gives the results of this study. The composition difference is increased to 1.5 mol % and then decreased to 0.8 mol %. The big effect is the impact of Δxspec on the recycle flow rate. In the base case, it is 116.6 kmol/h. When Δxspec is reduced to 0.8 mol %, the recycle flow rate decreases to 110.6 kmol/h. However, the load in the HPC increases because of the more difficult separation (reflux ratio increases from 2.693 to 2.99). Reboiler and condenser duties in the HPC increase. The low-pressure column appears to be more affected by recycle flow rate than by the difficulty of separation. The reflux ratio is slightly lower (5.551 versus 5.648 in the base case), and the duty in the auxiliary reboiler drops from 0.488 to 0.356 MW. The net result is a slight increase in total capital 5595
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7. EFFECT OF FEED COMPOSITION All of the cases discussed up to this point have used a feed composition of 50 mol % methanol and a feed flow rate of 100 kmol/h. With this feed composition, the feed is introduced into the high-pressure column because, as shown in Figure 5, the feed composition is above the azeotropic compositions at both the high pressure (21.24 mol % at 7 bar) and the low pressure (32.35 mol % at 0.25 bar). The concentration of methanol in the feed depends on the methanol conversion back in the reaction section of the process. Higher conversions mean lower methanol concentrations in the feed and lower feed flow rates to the pressureswing distillation system. To explore this situation, a feed composition corresponding to a fairly high reactor conversion is studied in this section. A feed composition of 15 mol % is assumed. The molar flow rate of trimethoxysilane leaving the reactor is kept constant at 50 kmol/h, so the new feed flow rate to the distillation system is 58.82 kmol/h. Since the 15 mol % methanol concentration is below the azeotropic compositions, the feed is introduced into the low-pressure column instead of into the high-pressure column. Figure 7 gives the revised flowsheet with a feed composition of 15 mol % fed on Stage 8 of the low-pressure column. The number of stages in both columns is kept at 31, and the pressure in the low-pressure column is kept at 0.25 bar so that cooling water can be used. A range of pressures in the highpressure column is explored to see if the optimum pressure remains at 7 bar. The feed tray locations in the new flowsheet are adjusted to give the minimum reboiler duties. A comparison of Figure 4 with Figure 7 shows a number of interesting changes as the feed composition and total feed flow rate are reduced. 1. The load on the HPC is much smaller because the distillate flow rate drops from 50 to 8.402 kmol/h. The reboiler duty is less than half despite a higher reflux ratio.
Table 3. Effect of Composition Difference between Azeotrope and Bottoms Specification (Δxspec) Δxspec xBHP xBLP Recycle HPC QR QC ID RR Energy Cost LPC Qaux QC ID RR Energy Cost Total Capital Total Energy TAC
(mol % MeOH) (mol % MeOH) (mol % MeOH) (kmol/h)
0.8 22.04 31.55 110.6
1.0 22.24 31.35 116.6
1.5 22.74 30.85 134.6
(MW) (MW) (m)
2.882 1.702 1.337 2.979 0.7471
2.797 1.579 1.316 2.693 0.7251
2.726 1.400 1.302 2.273 0.7067
(106 $/y)
0.356 3.126 2.335 5.551 0.0873
0.488 2.728 2.353 5.648 0.1157
0.716 3.330 2.413 5.980 0.1757
(106 $) (106 $/y) (106 $/y)
1.543 0.8344 1.3487
1.508 0.8448 1.3476
1.564 0.8823 1.404
(106 $/y) (MW) (MW) (m)
investment from $1,508,000 to $1,543,000 and a small decrease in total energy costs ($844,800 per year to $834,400 per year). The total annual cost of the base case (Δxspec = 1 mol %) is slightly lower than the Δxspec = 0.8 mol % case. When Δxspec is increased to 1.5 mol %, recycle flow rate increases to 134.6 kmol/h. This increases the load in the LPC with the auxiliary reboiler duty increasing from 0.488 to 0.716 MW. The reboiler duty in the HPC decreases because of the easier separation. The net effects are increases in both capital and energy costs. These results demonstrate that the assumed Δxspec = 1 mol % criterion is valid.
Figure 7. Flowsheet with low feed composition. 5596
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(8) Luyben, W. L. Pressure-Swing Distillation for Minimum and Maximum-Boiling Homogeneous Azeotropes. Ind. Eng. Chem. Res. 2012, 51, 10881−10886. (9) Childress, T. D.; Ritscher, J. S. Trimethoxysilane preparation via the methanol-silicon reaction using a continuous process and multiple reactors. US 5084590 A, 1991 (10) Tian, D.; Chu, X. H.; Yu, D. H.; Yue, Y. Z.; Zhoa, P.; Sun, X. L.; Liu, W. L. Immobilization of Polymethyl Methacrylate Brushes on Hydroxyapatite under Molecular Weight Control. Ind. Eng. Chem. Res. 2011, 40, 6109−6114. (11) Turton, R.; Bailie, R. C.; Whiting, W. B.; Shaelwitz, J. A. Analysis, Synthesis and Design of Chemical Processes, 2nd ed.; Prentice Hall: 2003.
2. Since the HPC condenser duty is smaller, the duty of the auxiliary reboiler in the LPC increases from 0.488 to 0.729 MW. 3. The recycle flow rate (LPC bottoms) decreases from 116.6 to 70.8 kmol/h because of the smaller feed flow rate to the process. The recycle flow rate depends directly on the feed flow rate and the difference between the bottoms compositions in the two columns. 4. Both columns have smaller diameters and smaller heatexchanger areas so capital investment is reduced. A range of pressures in the HPC were explored. The economic results showed that 7 bar gave the minimum TAC ($770,300 per year compared to the previous case of $1,348,000 per year). Total capital dropped from $1,508,000 to $991,100. Total energy cost dropped from $1,508,000 per year to $439,900 per year. These results show, as we would expect, that a higher conversion in the reaction section of the process reduces the load (energy costs in both the HPC reboiler and the auxiliary reboiler) and reduces the capital investment in the distillation methanol-recovery section of the process.
8. CONCLUSION An optimum economic design of a pressure-swing distillation process has been developed for separating the homogeneous maximum-boiling azeotrope in the binary methanol/trimethoxysilane system. The low-pressure column operates under vacuum conditions in which base pressure in strongly affected by the number of trays. The bottoms compositions of the two columns are near the azeotropic compositions at the appropriate pressure. Therefore the specified bottoms compositions must be adjusted as base pressures change. The heuristic of setting the bottoms specified composition at 1 mol % away from the corresponding azeotropic composition is demonstrated to be valid. Feed composition does not alter the design of the system except for the location where the feed is introduced.
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AUTHOR INFORMATION
Corresponding Author
*E-mail:
[email protected]; phone: 610-758-4256; fax: 610-7585057. Notes
The authors declare no competing financial interest.
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REFERENCES
(1) Stichlmair, J. G.; Fair, J. R. Distillation Principles and Practices; Wiley-VCH: 1998. (2) Doherty, M. F.; Malone, M. F. Conceptual Design of Distillation Systems; McGraw-Hill: 2001. (3) Luyben, W. L.; Chien, I. Design and Control of Distillation Systems for Separating Azeotropes; Wiley: 2010. (4) Abu-Eishah, S. I.; Luyben, W. L. Design and control of a twocolumn azeotropic system. Ind. Eng. Chem. Process Des. Dev. 1985, 24, 132−140. (5) Luyben, W. L. Distillation Design and Control using Aspen Simulation, 2nd ed.; Wiley: 2013. (6) Luyben, W. L. Comparison of Extractive Distillation and Pressure-Swing Distillation of the Acetone/Chloroform Separation. Comput. Chem. Eng. 2013, 50, 1−7. (7) Yu, B.; Wang, Q.; Xu, C. Design and Control of Distillation System for Methylal/Methanol Separation. Part 2: Pressure Swing Distillation with Full Heat Integration. Ind. Eng. Chem. Res. 2012, 51, 1293−1310. 5597
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