Decanter Flowsheet for Methanol Recovery in the TAME

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Ind. Eng. Chem. Res. 2009, 48, 10532–10540

Two-Stripper/Decanter Flowsheet for Methanol Recovery in the TAME Reactive-Distillation Process Yi Chang Wu and I-Lung Chien Department of Chemical Engineering, National Taiwan UniVersity of Science and Technology, Taipei 106, Taiwan

William L. Luyben* Department of Chemical Engineering, Lehigh UniVersity, Bethlehem, PennsylVania 18015

The process to produce TAME via reactive distillation requires a methanol-recovery section because the presence of C5/methanol azeotropes means that a significant amount of methanol is present in the distillate from the reactive column. The use of pressure-swing azeotropic distillation and extractive distillation were studied in a previous paper in which both the steady-state design and the plantwide control of the entire process were developed. This paper considers a third alternative flowsheet for the separation of the C5/methanol azeotropes that takes advantage of the heterogeneity of the azeotrope. Two stripping columns and a decanter are used. The total annual cost of this flowsheet is a factor of 4 less than that of the pressureswing system. The system is demonstrated to be easy to control by controlling a tray temperature in each stripping column. 1. Introduction The production of TAME (tert-amyl methyl ether) has grown in importance in recent years because of its use as an additive in gasoline. The predominant additive for many years was MTBE (methyl tert-butyl ether), but it is being phased out because of groundwater pollution concerns. Therefore, the production of TAME has become more widespread, and the importance of TAME as a gasoline blending component has increased in many petroleum refineries around the world. Studies of the design and control of this important system are industrially relevant. Subawalla and Fair1 studied a reactive distillation system for the production of TAME. The steady-state designs of a prereactor and a reactive distillation column were presented. Al-Arfaj and Luyben2 extended this work to design a complete plant with methanol recovery using a pressure-swing azeotropic separation method. They also developed a plantwide control structure for the three-column process with two recycles and two fresh feed streams. Luyben3 provided a comparison between pressure swing and extractive distillation for the recovery of the methanol. The C5 feed stream to the TAME process contains about 24 mol % reactive isoamylenes: 2M1B (2-methyl-1-butene) and 2M2B (2-methyl-2-butene). The remaining components are pentanes and pentenes (largely isopentane, iC5), which are inert in the TAME reaction. TAME is the highest boiling component, so it leaves in the bottoms stream from the reactive distillation column. The inert lighter C5 components leave in the distillate stream along with a significant amount of methanol. Methanol forms minimum boiling azeotropes with many of the C5’s. The reactive column operates at 4 bar (the optimum pressure1 that balances the temperature requirements for reaction with those for vapor-liquid separation). At this pressure, isopentane and methanol form an azeotrope at 339 K that contains 26 mol % methanol. Therefore, the distillate from the * To whom correspondence should be addressed. Tel.: 610-758-4256. E-mail: [email protected].

reactive column contains a significant amount of methanol, which must be recovered. Since the iC5/methanol azeotrope is pressure sensitive (79 mol % iC5 at 10 bar and 67 mol % iC5 at 4 bar), it is possible to use two distillation columns, operating at two different pressures, to separate methanol from the other C5 components. This pressure-swing process is the one studied in an earlier paper.2 An alternative separation process for this system is extractive distillation, which is mentioned in Stichlmair and Fair.4 These authors provide no details of either the steady-state design or the dynamic control. Luyben3 provided a comparison between pressure swing and extractive distillation for recovery of the methanol. The purpose of this paper is to present a third alternative flowsheet for methanol recovery from the C5/methanol stream using a decanter and two stripping columns. 2. Phase Equilibrium The liquid-phase reversible reactions that occur in the reactive distillation column are as follows 2M1B + MeOH S TAME 2M2B + MeOH S TAME 2M1B S 2M2B The bottoms is high-purity TAME. The distillate contains essentially all of the C5 components that are inert in terms of chemical reaction. Since these C5 components form minimumboiling azeotropes with methanol, the distillate also contains a significant amount of methanol. The phase equilibrium of this system is complex because of nonideality. There are a total of 8 components in the system requiring a total of 28 binary parameter pairs if NRTL or UNIQUAC is selected as the physical property method. However, in these 28 binary pairs only 12 have experimental data to obtain the binary parameters. Because the UNIFAC predictions for these 12 pairs match closely with the NRTL or

10.1021/ie900670f CCC: $40.75  2009 American Chemical Society Published on Web 09/15/2009

Ind. Eng. Chem. Res., Vol. 48, No. 23, 2009

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Figure 1. xy plots of the iC5/methanol binary system.

UNIQUAC predictions, the decision was made to use the UNIFAC physical property package in Aspen Plus for all 28 pairs to model the vapor-liquid-liquid equilibrium (VLLE). The major C5 component in the feed to the reactive column is isopentane (iC5), and we illustrate the VLLE behavior using the iC5/methanol binary system. Figure 1 shows the phase behavior at different operating pressures in the form of an xy plot (liquid-phase composition versus vapor-phase composition). Notice that two liquid phases are present (at either end of the horizontal portions of the curves). Thus, a binary mixture of isopentane and methanol will split into two liquid phases in a decanter. Figure 2 gives Txy diagrams for isopentane/methanol at two different pressures: 1 and 2.4 atm. Notice that the temperature of the azeotrope at 1 atm is about 298 K, which makes the use of cooling water in the condenser problematic. However, at 2.4 atm the temperature is about 323 K, so cooling water can be economically used in the condenser. A similar binary heterogeneous system was studied by Luyben5 in the n-butane/water system. A flowsheet suggested by Doherty and Malone6 was used for this study and found to be economically attractive and easy to control. The flowsheet consisted of two stripping columns and a decanter. In the present paper we study the use of this flowsheet for the C5/methanol separation. 3. Process Studied The distillate from the reactive distillation column is the feed stream to the methanol recovery section of the process. It has a composition of methanol (28 mol % methanol) that is near the azeotrope at 4 bar. It is fed at a rate of 1126 kmol/h into the decanter, which operates at 320 K and 2.4 atm, so that cooling water can be used in the condensers. Figure 3 shows the process flowsheet. The light phase (lower density) from the decanter is rich in C5 components and has a density of 617 kg/m3. The heavy phase (higher density) from the decanter is rich in methanol and has a density of 722 kg/m3. It is important to notice that the density difference between the two liquid phases is quite small, so the

separation may be difficult. A holdup time of 30 min in the decanter is assumed in equipment sizing and dynamic studies. The heavy (methanol rich) phase is fed to a stripping column C1, which has a reboiler but no condenser. This column operates at 2.6 atm to provide a pressure drop through a control valve in the overhead vapor line and through the condenser. The bottoms has a composition of 99.9 mol % methanol and is recycled back to the TAME reactive distillation column. The overhead vapor is condensed in a heat exchanger, and the condensate flows into the decanter. The light (C5 rich) phase is fed to a second stripping column C2, which also has a reboiler but no condenser. This column operates at 10 atm. The C5 components are removed in the bottoms stream. The loss of methanol in the bottoms is specified to be 1% of the methanol in the feed. The resulting bottoms composition of methanol impurity is 0.4 mol %. The overhead vapor is condensed in a second heat exchanger, and the condensate flows into the decanter. The reason for selecting the operating pressures in the two strippers is the effect of pressure on the ease of separation. Figure 1 shows that in the region on the left side of the plot the separation becomes more difficult as pressure increases (the xy curve is closer to the 45° line). The left side is the methanolrich region, which occurs in stripper C1 that is fed the heavy methanol-rich stream from the decanter. Therefore, the pressure in C1 should be minimized. A pressure of 2.6 atm is selected so that cooling water can be used in the condenser. Figure 1 also shows that in the region on the right side of the plot the separation becomes more difficult as pressure decreases (the xy curve is closer to the 45° line). The right side is the isopentane-rich region, which occurs in stripper C2 that is fed the light C5-rich stream from the decanter. Therefore, the pressure in C2 should be maximized. A pressure of 10 atm is selected so that the temperature in the base still permits the use of low-pressure steam.

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Figure 2. (A) Txy plots for iC5/methanol at 1 atm. (B) Txy plots for iC5/methanol at 2.4 atm.

4. Economic Optimization The flowsheet was optimized by finding the equipment sizes and operating conditions that minimize total annual cost. There are only two design optimization variables in this system: the number of stages in stripper C1 and the number of stages in stripper C2; thus, the exhaust optimization approach was taken to explicitly enumerate these two variables. The number of stages in each stripper was varied, and for each case the energy requirements (the reboiler duties in both reboilers) and the capital costs (columns, heat exchangers, and decanter) were calculated. The sizing relationships and economic basis used for these calculations are given in Table 1. Annual energy cost is the sum of the two reboiler duties in kJ/ year multiplied by an energy cost of $4.7 per 106 kJ. Annual

capital cost is the total capital investment of equipment divided by a 3-year payback period. Figure 4 gives results of the optimization study. Total annual cost is plotted against the number of stages in C2 for several values of the number of stages in C1. The TAC is minimized when there are 6 stages in C2 and 8 stages in C1. The first column in Table 2 gives details of equipment sizes and economics for the optimum design. These are compared with the pressure-swing flowsheet, to be discussed in the next section. The two-stripper/decanter flowsheet is much less expensive in terms of both capital investment and energy cost. It should be noted that the use of heat integration is possible in this system. A further increase in the pressure in C2 would be required to increase the temperature difference between the


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Figure 3. Two-stripper/decanter flowsheet for C5/methanol.

of the two components being separated. Recycling a liquid stream is less energy intensive than recycling a vapor stream, so maximum-boiling azeotropes can be separated less expensively.

Table 1. Basis of Economics

condensers heat-transfer coefficient differential temperature capital cost reboilers heat-transfer coefficient differential temperature capital cost column shell and decanter capital cost energy cost TAC payback period

0.852 kW/K · m2 13.9 K 7296 (area)0.65 [area in m2] 0.568 kW/K · m2 34.8 K 7296(area)0.65 [area in m2] 17 640 (D)1.066 (L)0.802 [diameter and length in meters] $4.7 per 106 kJ capital cost/payback period + energy cost 3 years

reflux-drum temperature in C2 and the reboiler temperature in C1. The potential energy savings is only about 15% (1.187 MW in the reboiler of C1 compared to 7.213 MW in the reboiler of C2). Heat integration has been discussed in a previous paper7 and is not explored here. 5. Alternative Flowsheets Two other flowsheets for the recovery of methanol have been studied in previous papers.2,3 A brief discussion of these alternatives is given in this section. 5.1. Pressure Swing. Separation of a binary azeotrope using pressure swing is feasible if the composition of the azeotrope is pressure dependent. The flowsheet has two distillation columns operating at different pressures. Feed is introduced into one of the columns. If the azeotrope is minimum boiling, the distillate streams from both columns will have compositions close to their azeotropic compositions and will be fed to the other column. The bottoms streams will be high-purity streams of the two components being separated. In those cases where the azeotrope is maximum boiling, the bottoms streams from both columns will have compositions close to their azeotropic compositions and will be fed to the other column. The distillate streams will be high-purity streams

The C5/methanol azeotropes are minimum boiling, so the recycles are the distillate streams. Details of the design are given in previous papers.2,3 Feed is introduced into the low-pressure column C1. The distillate D1 is fed to the high-pressure column. Its composition (22.8 mol % methanol) is near the azeotropic composition at the 2 bar pressure. Reboiler heat input and condenser heat removal are 19.5 and 24.5 MW, respectively. The column diameter is 4.2 m. The pressure in the high-pressure column is set at 10 bar, which shifts the azeotropic composition so that the distillate stream from this column D2 has a composition of 34.2 mol % methanol. Higher and lower pressures were explored to see their effect on the economics. The 10 bar pressure seems to be about the optimum since going above this pressure does not shift the azeotrope significantly and raises the base temperature, which would require higher temperature energy input. The separation is a fairly easy one, so using only 10 stages and a reflux ratio of 1 yield a bottoms impurity of 0.01 mol % methanol. This bottoms stream B2 is the C5 product stream. The distillate is recycled back to C1 at a flow rate of 1616 kmol/ h. Reboiler heat input and condenser heat removal are 27.2 and 20.2 MW, respectively. The column diameter is 6.2 m. The economics of this pressure-swing system are compared with the two-stripper/decanter process in Table 2. Clearly the latter flowsheet is much more economical. 5.2. Extractive Distillation. Another flowsheet discussed in the literature4 is the use of water as an extractive agent. A water solvent is fed near the top of the column. The feed is introduced lower in the column. The extractive agent is selected so that it pulls methanol out of the mixture and takes it out the bottom. The C5 components leave out the top of the column. This process was studied by Luyben4 and stated to be much better than the pressure-swing process. In this early study, the “valid phases” in the RadFrac model in Aspen Plus were

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Figure 4. Optimization of a two-stripper/decanter flowsheet.

6. Control

Table 2. Design Parameters and Economic Results two-stripper pressure decanter swing process process C1

sizes

capital costs

C2

energy sizes

capital costs

decanter total capital total energy TAC

energy capital cost

diameter (m) stages QC (MW) AC (m2) QR (MW) AR (m2) Shell cost (106 $) HX cost (106 $) total capital (106 $) energy cost (106 $/y) diameter (m) stages QC (MW) AC (m2) QR (MW) AR (m2) Shell cost (106 $) HX cost (106 $) total capital (106 $) energy cost (106 $/y) Shell cost (106 $) (106 $) (106 $/y) (106 $/y)

0.7509 8 0.7775 65.65 1.187 60.04 0.04819 0.2152 0.2634 0.1759 2.785 6 4.295 362.7 7.213 364.9 0.1488 0.6740 0.8228 1.069 0.3529 1.439 1.245 1.725

4.2 10 24.5 2065 19.5 985 0.339 1.69 2.02 2.89 6.2 10 20.2 1703 27.2 1374 0.508 0.304 2.23 4.03 4.25 6.92 8.34

specified to be VLE. This is an error in the original work because the C5/methanol system is heterogeneous, as shown in Figure 1. During a recent look at this system, the “valid phases” in the RadFrac model in Aspen Plus were specified to be VLLE. The resulting predicted separation is very poor. Significant amounts of C5 components go out the bottom with the water. The temperature profile is inverted with high temperatures in the top and low temperatures in the bottom. It is not clear if the Aspen simulation is valid or not. The possibility of numerical issues is always present. We continue to explore this problem, but at this point all we can say is that it is unclear whether or not extractive distillation using water as the solvent is feasible. The use of other solvents is being explored.

The two-stripper/decanter system for n-butanol/water was studied in a recent paper,5 and a simple but effective control structure was developed. This same control system was tested for the C5/methanol system considered in this paper. Figure 5 shows this control scheme. The feed is introduced into the decanter on flow control. The interface level in the decanter is controlled by manipulating the flow rate of the heavy methanol-rich liquid phase to the top tray in the stripper C1. The top liquid level in the decanter is controlled by manipulating the flow rate of the light C5-rich liquid phase to the top tray in the stripper C2. The pressure in each column is controlled by a valve in the overhead vapor line. Heat removal in the condensers controls the condenser exit temperature. Base level in each column is controlled by manipulating the bottoms product flow rates. The base of each column is sized to provide 5 min of holdup when half full, based on total liquid entering the base. The decanter is sized to provide 30 min of holdup based on the sum of the volumetric flow rates of the two liquid phases. A tray temperature in each column is controlled by manipulating the corresponding reboiler heat input. The trays are selected by finding the location where the temperature profile is steep. Figure 6 gives the temperature profiles in the two columns. Stage 4 is selected in stripper C1, and Stage 5 is selected in stripper C2. The temperature controllers have 1 min deadtimes. The PI controllers are tuned by running a relay-feedback test and using the Tyreus-Luyben settings. Table 3 gives controller parameters and tuning constants. All liquid levels are controlled by proportional controllers with gains of 2. Figure 7A shows the Aspen Dynamics flowsheet with all controllers installed. Figure 7B shows the controller faceplates. Figure 8 gives the responses to positive and negative 20% changes in the feed to the process occurring at 0.5 h. The solid lines are 20% increases in feed flow rate. The dashed lines are 20% decreases. The control structure maintains the purities of both product streams very close to their specification. The system reaches a new steady state in about 4 h.


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Figure 5. Control structure.

Figure 6. Temperature profiles.

Feed composition disturbances are also tested. The responses of the system to positive and negative changes in the methanol composition of the feed are shown in Figure 9. The solid lines are when the feed composition is changed from 28.14 mol % methanol and 44.46 mol % isopentane to 38.14 mol % methanol and 34.46 mol % isopentane. The dashed lines are when the feed composition is changed from 28.14 mol % methanol and 44.46 mol % isopentane to 18.14 mol % methanol and 54.46

Table 3. Controller Tuning Constants

set point (°C) transmitter range (°C) controller output (GJ/h) output range (GJ/h) KC τI (min)

C1

C2

97.0 50-150 4.14 0-8.23 0.35 9.2

114.6 50-150 26.0 0-100 0.81 9.2

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Figure 7. (A) Aspen Dynamics flowsheet. (B) Controller faceplates.

mol % isopentane. The control structure handles these very large disturbance quite well. 7. Decanter Simulation

pressure in the decanter are the usual variables specified. Then the required heat removal is calculated. This is unrealistic because most industrial decanters are simple drums with no heat exchange capability.

To conclude this study it may interest the reader to discuss the issue of pressure, temperature, and heat removal in the decanter. There are two different models in Aspen software that can be used for modeling a two-liquid phase vessel. The Decanter model is the more simple one and the one used in this study. The only streams exiting the decanter are the two liquid phases. There is no vapor stream. The temperature and

More importantly, the pressure remains constant during the dynamic simulation in the Decanter model. There is no modification of the pressure as the composition and/or temperature change in the decanter. This is certainly not what happens in a real decanter. However, this misrepresentation only affects the hydraulics of the flows around the decanter and probably has a minor effect on the performance of the system.


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Figure 8. Feed flow rate disturbances.

Figure 9. Feed composition disturbances.

The alternative and more rigorous approach is to use the Flash 3 model. The vessel has three exit streams: two liquid streams and one vapor stream. Now the pressure varies during the simulation as compositions and temperatures in the decanter change with time. A convenient configuration is to feed a small inert gas stream into the decanter and take a small vapor stream off the decanter. A pressure control holds the pressure in the decanter by positioning the two control valves in the gas feed and vapor lines. The use of split-ranged valve prevents the simultaneous addition and removal of gas from the decanter. When pressure is low, inert gas is added and the valve in the exit gas line is shut. When pressure is high, gas is vented and the valve in the inert gas addition line is shut.

8. Conclusion A two-stripper/decanter flowsheet has been demonstrated to be a very economical process for separating the C5/methanol azeotropes coming from the top of a TAME reactive distillation column. The process is also easily controlled by a simple PID control structure. Very large disturbances in throughput and feed composition are effectively handled. Two temperatures are controlled, one in each stripping column. The purities of both product streams are held very close to their specifications. Nomenclature AC ) heat transfer area of condenser (m2) AR ) heat transfer area of reboiler (m2)

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area ) area (m2) Bn ) bottoms flow rate from column n (kmol/h) Cn ) column n D ) diameter of vessel (m) Dn ) distillate flow rate from column n (kmol/h) HX ) heat exchanger L ) length of vessel (m) QC ) condenser heat removal (MW) QR ) reboiler heat input (MW) TAC ) total annual cost (106 $/y) ∆F ) change in feed flow rate ∆z ) change in feed composition

Literature Cited (1) Subawalla, H.; Fair, J. R. Design guidelines for solid-catalyzed reactive distillation systems. Ind. Eng. Chem. Res. 1999, 38, 3693.

(2) Al-Arfaj, M. A.; Luyben, W. L. Plantwide control for TAME production using reactive distillation. AIChE J. 2004, 50, 7–1462. (3) Luyben, W. L. Comparison of pressure-swing and extractive distillation methods for methanol recovery systems in the TAME reactivedistillation process. Ind. Eng. Chem. Res. 2005, 44, 5715. (4) Stichlmair, J. G.; Fair, J. R. Distillation: Principles and Practices; Wiley: New York, 1998. (5) Luyben, W. L. Control of the heterogeneous azeotropic n-butanol/ water azeotropic system. Energy Fuels 2008, 22, 4249–4258. (6) Doherty, M. F., Malone, M. F. Conceptual Design of Distillation Systems; McGraw-Hill: New York, 2001; p 383. (7) Luyben, W. L. Design and control of a fully heat-integrated pressureswing azeotropic distillation system. Ind. Eng. Chem. Res. 2008, 47, 2681.

ReceiVed for reView April 27, 2009 ReVised manuscript receiVed August 21, 2009 Accepted August 26, 2009 IE900670F

Decanter Flowsheet for Methanol Recovery in the TAME

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