Production of Butyl Acetate by Catalytic Distillation: Process Design

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Ind. Eng. Chem. Res. 2004, 43, 136-143

Production of Butyl Acetate by Catalytic Distillation: Process Design Studies Jignesh Gangadwala,† Achim Kienle,*,†,‡ Erik Stein,† and Sanjay Mahajani§ Max-Planck-Institute for Dynamics of Complex Technical Systems, Sandtorstrasse 1, D-39106 Magdeburg, Germany, Institut fu¨ r Automatisierungstechnik, Otto-von-Guericke Universita¨ t, D-39106 Magdeburg, Germany, and Department of Chemical Engineering, Indian Institute of TechnologysBombay, Powai, 400076 Mumbai, India

A simulation-based design method is employed to figure out the promising reactive distillation process configuration for the production of butyl acetate. The intrinsic kinetics developed for the esterification and the unwanted side reaction etherification over the Amberlyst-15 catalyst1 are utilized to evaluate the steady-state performance of different reactive distillation processes. A steady-state column model is developed and compared with experimental data from the literature.2,3 With this model, three different column configurations are investigated for the production of butyl acetate with the goal of eliminating the formation of byproduct dibutyl ether and achieving a high purity of the desired product butyl acetate. The following column configurations are explored: (a) a column with a nonreactive rectifying section and a reactive stripping section; (b) a column with a nonreactive rectifying section, a nonreactive stripping section, and a reactive middle section; and (c) a conventional distillation column with a cocurrent pump-around reactor. Configuration c is compared with the side reactor configuration, where the pump-around reactor is coupled to the column in a countercurrent fashion. 1. Introduction Reactive distillation is a state of the art technology for carrying out many equilibrium-limited chemical reactions. However, for some chemical systems, this alternative is not even feasible. In the cases where it is feasible, it can be very attractive compared to conventional technology for some suitable process designs. However, very often many different design alternatives are possible. For a recent review on the state of the art in reactive distillation, we refer to work by Sundmacher and Kienle.4 This work deals with simulation and design studies for butyl acetate synthesis by reactive distillation. Butyl acetate (BuAc) can be produced by esterification of n-butanol (BuOH) with acetic acid (AcH) in the presence of a suitable acid catalyst. Classical processes make use of a homogeneous acid catalyst either in continuous reactive distillation5 or in batch reactive distillation.6 However, severe corrosion problems associated with homogeneous acid catalysts (e.g., sulfuric acid or ptoluenesulfonic acid) make these processes less attractive. In recent years, solid heterogeneous catalysts are receiving attention because of their obvious engineering benefits such as ease of separation and fewer disposal and corrosion problems. Recently, the synthesis of BuAc has been studied by several workers in a continuous reactive distillation column (RDC) with solid acid catalysts.3,7,8 However, none of them have mentioned * To whom correspondence should be addressed. Tel.: +49-391-6110369. Fax: +49-391-6110515. E-mail: kienle@ mpi-magdeburg.mpg.de. † Max-Planck-Institute for Dynamics of Complex Technical Systems. ‡ Otto-von-Guericke Universita ¨ t. § Indian Institute of TechnologysBombay.

the formation of unwanted side product dibutyl ether (DBE), which forms by BuOH dehydration in the column section with high BuOH concentration. DBE is considered to be a poisonous substance, and because BuAc is a widely used solvent for many end products, DBE has to be completely eliminated from the final product. The authors have investigated the reaction kinetics for esterification and etherification over an Amberlyst-15 catalyst.1 In their work they have also studied the favorable conditions for side product formation and concluded that etherification is favored at high catalyst loading, higher temperature, and high BuOH to AcH mole ratio. In addition, catalyst drying has a significant effect on ether formation, which can also be depicted from the etherification rate expression. Because one is very likely to encounter such conditions in a reactive stripping section, knowledge of the side reaction kinetics is required for detailed simulations. In the present work, an attempt is made to find out a suitable reactive distillation configuration, which can not only minimize the formation of the unwanted side product but also give higher conversion of the limiting reactant. In this paper, at first an equilibrium stage model is formulated and compared with experimental data from the literature.2,3 Using this model, the influence of the most important design variables on the steady-state performance of three different configurations is studied. Conclusions are drawn on which of these configurations is most suitable. 2. Mathematical Model The mathematical model for the RDC is developed based on the following assumptions.

10.1021/ie021011z CCC: $27.50 © 2004 American Chemical Society Published on Web 12/03/2003

Ind. Eng. Chem. Res., Vol. 43, No. 1, 2004 137

Figure 2. RDC setup from Janowsky et al.2

Figure 1. Nonreactive residue curve maps in a BuAc system at 1 atm, adopted from Kienle et al.12 (b) represents azeotropes at 1 atm. x represents azeotropes at 650 mbar.

(1) An equilibrium stage model is used. (2) The stages are numbered from top to bottom. Stage 1 represents the uppermost column tray, and stage N represents the partial reboiler. A total condenser is assumed between the column and the decanter. (3) BuAc creates a low-boiling ternary heterogeneous azeotropic mixture with BuOH and water (see Figure 1), which is separated at the top of the column. Upon condensation, it separates into two liquid phases in the decanter. The organic phase consists of BuAc, BuOH, and water, which is refluxed back to the column, whereas the aqueous phase (mainly water) is removed as a distillate. Thus, the reflux consists of the organic phase only (no degree of freedom at the top), unless otherwise mentioned. For the decanter model, liquidliquid-phase equilibrium is assumed. Figure 1 shows the nonreactive residue curve map for the quaternary mixture consisting of AcH, BuOH, BuAc, and water, plotted at atmospheric pressure. The arrows show the direction of increasing temperature. As can be seen, there is a total number of six azeotropes, three binary minimum azeotropes, a binary maximum azeotrope, a ternary minimum azeotrope, and a ternary saddle azeotrope. The binary azeotropes between BuOH and water and between BuAc and water and a ternary azeotrope between BuOH, BuAc, and water are heterogeneous, whereas the rest of the azeotropes are homogeneous. As will be discussed later, most of the simulations were carried out at a pressure of 650 mbar. It is therefore necessary at least qualitatively to determine the behavior of the quaternary mixture at this pressure. It was found that all azeotropes remain present at 650 mbar (x in Figure 1) and that the qualitative behavior of residue curve map remains unchanged. (4) The activity coefficients of the liquid phase inside the column are calculated using the Wilson model. No liquid-liquid-phase splitting is assumed inside the column, whereas for the decanter UNIQUAC equations are used to calculate the activity coefficients for the liquid-liquid equilibrium. The Wilson and UNIQUAC binary interaction parameters are obtained from the DECHEMA database. They are shown in Gangadwala et al.1 (5) The chemical reactions only take place in the presence of the heterogeneous catalyst. The reaction kinetics for the esterification and the etherification are

obtained from Gangadwala et al.1 The reaction schemes are as follows.

BuOH + AcH h BuAc +H2O 2BuOH f DBE + H2O r1i )

1 dni ) νi dt aAcHaBuOH -

1 a a Ka BuAc H2O

(1) Mcat.KfKs,AcHKs,BuOH (1 + a′AcH + a′BuOH + a′H2O)2 r2i )

1 dni ) Mcat.KDBE νi dt (1 + a′

a′BuOH2

BuOH

+ a′DBE + a′H2O)2 (2)

where

a′i ) Ks,iai (6) Pressure drop and heat losses across the column wall are neglected. (7) Fluid dynamics of the column internals is ignored. Clearly, this is a fairly simple model. It, however, contains a reasonable description of the reaction kinetics and the phase equilibrium. Hence, it is suitable for some conceptual design studies to be considered in this paper. A detailed study including hydrodynamics, mass transfer, etc., is beyond the scope of this paper. 3. Model Comparison with Experimental Data The above steady-state model is solved by the DIVA process simulator.9 Because no experiments are conducted to support the simulation work, the model is compared with experimental data from the literature.2,3 3.1. Model Comparison with Experimental Data from Janowsky et al.2 Janowsky et al.2 carried out BuAc synthesis experiments to study the steady-state column performance at three different pressures ranging from 650 to 1105 mbar (Figure 2). At higher pressure, they observed a significant amount of 1-butene at the top of the column. Also the unwanted byproduct DBE was found with the main product BuAc in the bottoms. They were able to get rid of 1-butene formation by decreasing the column pressure up to 650 mbar but were unable to eliminate DBE from the bottoms. Because the model presented in the previous section does not include the kinetics of 1-butene formation, the

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Table 1. Inputs to the Column from Janowsky et al.2,a component FC (mole fraction) component FC (mole fraction) AcH BuOH BuAc a

0.1535 0.1676 0.4491

water DBE

0.2297 0.0001

Feed flow rate: 0.0215 kmol/h. Tf ) 343 K.

comparison of the model with experiments is restricted to the lowest pressure, i.e., 650 mbar. Janowsky et al. have used a packed column in their experiments. Each column section in the experimental column is equivalent to 15 theoretical stages. The stripping section filled with catalyst acts as a reactive section. The feed is provided at the top of the reactive section, i.e., on the 16th stage. Table 1 shows the feed composition (FC). The feed to the column comes from a prereactor. Hence, it contains a considerable amount of BuAc and water, which are the products of the esterification. The simulation has been performed using design specifications, like the number of stages, reactive stages, and feed location, from Janowsky et al.2 The reboiler heat duty (0.197 kW) and catalyst loading (1.4 kg/tray), which are the most sensitive design variables and which are not given in Janowsky et al., are adjusted to the experimental data. The decanter and reflux temperatures were found to be insensitive variables. They were fixed at 328 and 336 K, respectively. Table 2 shows results of the simulation run in comparison to the experimental data. It can be seen that bottom compositions and aqueous phase compositions

predicted by the model are in good agreement with the experimental data. The mismatch in some data, e.g., flow rates and organic phase compositions, can possibly be due to limited accuracy in the experimental measurements or the limited scope of the model. Nevertheless, we are able to reproduce the experimental results quite satisfactorily. Hence, it is justified to use this model for studies on conceptual process design for BuAc synthesis to be considered subsequently. 3.2. Model Comparison with Experimental Data from Hanika et al.3 Figure 3a shows the RDC setup used by Hanika et al., where a reactive middle section separates the nonreactive separation sections. The experimental column was a packed column having 20 theoretical stages in each separation section, with a feed introduced into the catalytic section at the 23rd stage. Figure 3b shows the comparison between the experimental temperature profile and the simulation results obtained for the column inputs given in Table 3. The feed flow rate and FCs given in Table 3 were selected in the range described by Hanika et al. The reboiler duty (0.5 kW) and catalyst loading (0.14 kg/ tray) are adjusted to the experimental data. The reflux flow that consists of 95% of the organic phase enters the column top at 360.15 K. The decanter temperature was fixed at 313.15 K. The agreement between the experimental data and the model is quite satisfactory. This indicates further applicability of the model for the conceptual design of BuAc synthesis opted for in this paper. It should be

Figure 3. (a) RDC setup from Hanika et al.3 (b) Comparison of temperature profile experimental data and model results. Table 2. Comparison between Experimental Data from Janowsky et al.2 and the RDC Modela aqueous phase from the decanter flow rate (kmol/h) AcH BuOH BuAc Water DBE temp (K) a

experiments

model

0.0119 0.0000 0.0040 0.0008 0.9952 0.0000 351.15

0.0083 0.0000 0.0090 0.0020 0.9889 0.0000 355.32

Compositions are in mole fraction.

organic phase from the decanter

bottoms

experiments

model

experiments

model

0.0000 0.1171 0.4221 0.4423 0.0186 351.15

0.0070 0.0000 0.2759 0.4759 0.2480 0.0000 355.32

0.0127 0.0096 0.0000 0.9609 0.0064 0.0231 384.15

0.0132 0.0155 0.0058 0.9649 0.0001 0.0137 384.13

Ind. Eng. Chem. Res., Vol. 43, No. 1, 2004 139 Table 3. Inputs to the Column from Hanika et al.3,a component FC (mole fraction) component FC (mole fraction) AcH BuOH BuAc a

0.1665 0.1797 0.3211

water DBE

0.3327 0.0000

Feed flow rate: 0.0188 kmol/h. Tf ) 360 K.

recalled that our model assumed byproduct formation by irreversible reaction, whereas Hanika et al. reports no byproduct formation. In their simulation study, Hanika et al. have used chemical equilibrium but did not incorporate the DBE kinetics, so catalyst loading may be considered as a nonsignificant parameter for their model. However, in our model catalyst loading has a strong influence on the DBE formation in the column. The catalytic section is mostly filled with BuOH and BuAc; therefore, it would produce more DBE if more catalytic stages or catalyst loading are provided. For the parameters given in Table 3, the model predicts a negligible amount (0.02 mol %) of DBE in the bottoms. The suppression of DBE formation, in this case, is mainly due to the introduction of the feed into the catalytic section because, in this case, more AcH is available in the catalytic section, which reacts away the BuOH and thereby “inhibits” the side reaction, producing DBE. 4. RDC Design The model developed in the previous section can be used to compare various possible RDC configurations for the production of BuAc and to determine the most suitable one with respect to elimination of DBE and AcH from the bottom products. Figure 4 shows three different RDC configurations to be considered subsequently: (a) a RD column having a reactive stripping section; (b) a RD column having a reactive middle section; (c) a distillation column with a cocurrent pump-around reactor. For configuration c, a continuous stirred tank

reactor is used. The catalyst loading of the reactor is comparable to that required in the RDC of configurations a and b. It has been noticed that, a very high pump-around flow rates, more than the feed flow rate is required to get a satisfactory performance. However, to maintain such a high pump-around rate, a very high reboiler duty is necessary. As was said earlier, the entire organic phase from the decanter is refluxed to the column, whereas the aqueous phase is withdrawn with the distillate. Recalling that at low pressures 1-butene formation is suppressed, all simulation runs are carried out at 650 mbar. Two FCs are considered: (1) a reaction mixture coming from the prereactor, for which compositions are given in Table 1 (the following discussion mainly focuses on this case) and (2) a stoichiometric mixture of BuOH and AcH with a slight excess of BuOH (this case is briefly described at the end). To determine suitable values of the design variables, viz., reboiler duty, catalyst loading, catalyst section length and location, and feed tray location, continuation methods implemented in the process simulator DIVA9 are used. The base case parameters for each configuration used in this study are given in Table 4. In the following discussion, because configurations a and b exhibit similar characteristics, they are grouped together, whereas configuration c is treated separately. At the end, a section is added to illustrate the important features of the countercurrent side reactor configuration (configuration d). Furthermore, in this study, no attempt is made to optimize the number of column stages for any configuration because it is found that the column performance is insensitive to the number of stages if more than 20 column stages are provided. Therefore, all the investigations discussed hereafter consider 22 column stages. 4.1. Reboiler Heat Duty. In general, for configurations a and b, increasing the reboiler heat duty increases the BuAc concentration in the bottoms but at the same time increases the DBE concentration as well. It is even

Figure 4. Reactive distillation configurations: (a) column with nonreactive rectifying and reactive stripping sections; (b) column with nonreactive rectifying, nonreactive stripping, and reactive middle sections; (c) conventional distillation column with a cocurrent pumparound reactor; (d) conventional distillation column with a countercurrent pump-around reactor. The reactive section is shown by the shaded area to distinguish it from the nonreactive section. Table 4. Base Case Parameters for Configurations a-d config.

total stages

reactive section/ reactor location

feed tray location

reboiler heat duty (kW)

catalyst loading (kg of catalyst/tray)

pump-around flow rate (kmol/h)

a b c d

22 22 22 22

12-21 trays 9-17 trays 6-7 trays 6-7 trays

11th tray 8th tray 6th tray 6th tray

0.4 0.4 2.0 1.0

0.2 0.2 2.0 4.0

0.2 0.2

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Figure 5. Effect of the reboiler heat duty for configurations a and b. The bottom compositions are plotted against the heat duty.

Figure 6. Effect of the reboiler heat duty for configuration c. The bottom compositions are plotted against the heat duty.

possible to get 100% conversion of AcH at higher reboiler duty. This is expected because increasing the heat duty increases the reactive section temperature, which enhances the conversion but reduces the selectivity (Figure 5). It should be noted that each point on the curve in Figure 5 corresponds to a steady-state solution for a given value of the reboiler heat duty. In further simulation runs, a heat duty of 0.5 kW will be used for configurations a and b. The model for configuration c predicts multiple steady states for fixed values of the reboiler duty as shown in Figure 6. For the range of reboiler duty shown in Figure 6, there exist three steady-state solutions. The upper steady-state solution branch gives higher conversion and selectivity (therefore desirable), whereas the other two branches give lower conversion and selectivity (therefore undesirable). However, unlike configurations a and b, the DBE concentration in the bottoms remains unchanged with increasing heat duty for the desirable multiple steady-state branch. In further simulation runs, a heat duty of 3 kW will be used for configuration c. 4.2. Catalyst Loading. Increasing the catalyst loading in configurations a and b initially increases the conversion and, hence, the BuAc and DBE concentrations in the product. However, when the catalyst loading is further increased, the BuAc concentration decreases, whereas the DBE concentration still increases. At higher catalyst loading, the side reaction becomes favorable, consuming more BuOH. Therefore, conversion of AcH decreases, and unreacted AcH can be found

Figure 7. Effect of the catalyst loading for configurations a and b. The bottom compositions are plotted against the catalyst loading per tray.

Figure 8. Effect of the catalyst loading for configuration c. The bottom compositions are plotted against the catalyst loading.

in the bottoms. In summary, it can be stated that providing more catalyst adversely effects the bottom product purity and the maximum in the BuAc concentration can be seen from Figure 7. The optimum catalyst loading is found to be 0.25 kg of catalyst/tray. Configuration c again exhibits multiple steady states for the catalyst loading as shown in Figure 8. Multiple steady states here involve three branches in the range of 0.88-3.55 kg of catalyst loading. Three steady-state solutions at Qr ) 3.0 kW and Mcat. ) 2.0 kg should be compared in Figures 6 and 8. To avoid any multiplicities, further simulation runs will be performed using the catalyst loading of 4 kg, for which only a single steady-state solution exists. 4.3. Feed Tray Location. Because the boiling point difference between BuOH (117.7 °C) and AcH (117.9 °C) is rather small, the idea of providing feed at two separate points is not very attractive. This has been observed from the simulation runs for configurations a and b. Therefore, in the following discussion, only a single feed case is considered. The effect of an increase in the feed flow rate is also investigated and, as expected, the trend is to decrease the conversion. As the feed tray is shifted down in the reactive section, the conversion of reactants decreases, leading to a decrease in the BuAc concentration in the products. Moreover, a high concentration of AcH in the reactive section inhibits the DBE formation. The same arguments can be applied to the shifting of the feed tray up in the rectifying section, which results in the high

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Figure 9. Effect of the number of reactive stages for configuration a. The bottom compositions are plotted against the number of reactive stages.

concentration of BuAc and DBE in the bottom products. Therefore, a tradeoff is required between the AcH and DBE concentrations in the bottoms. The best feed tray location for configurations a and b is observed at the top of the reactive section. For configuration c, shifting the feed tray up and down causes effects similar to those for configurations a and b. The best feed tray location is found at the tray from where the pump-around is withdrawn from the column. 4.4. Number of Reactive Stages and Reactive Section Location for Configurations a and b. As was said earlier, providing more catalyst degrades the product purity; therefore, the optimum number of catalytic stages must be provided in the reactive section. Figure 9 shows the effect of the number of reactive stages for configuration a with the feed tray at the top of the reaction section. It can be seen that 10 catalytic stages give a slightly better performance in terms of the product purity in the bottoms. For configuration b, with some effort it is found that 9 catalytic stages provided between the 9th and 17th stages can give a better column performance. 4.5. Pump around Location for Configuration c. It is observed that two stages between pump-around withdrawal and introduction gives a better performance compared to more stages between the pump-around. The pump-around position is then varied from 2-3 to 19-20 (withdrawal-introduction), maintaining the feed position at the same stage from where the pump-around is withdrawn. Figure 10 shows the bottoms composition vs pump-around location. As can be seen from Figure 10, the pump-around location between the 6th and 7th stages with the feed introduced into the 6th stage gives a comparatively good performance. 4.6. Pump around Flow Rate for Configuration c. As shown in Figure 11, multiple steady states exist for the pump-around flow rate. It should be recalled that multiple steady states were also observed for the catalyst loading of the reactor. This is expected because multiple steady states prevail for a fixed range of the Damko¨hler number, which can be obtained either by varying the catalyst loading or pump-around rate. Multiple steady states are observed for values between 0.055 and 0.173 kmol/h of the pump-around rate. For the desired branch, the pump-around flow rate of 0.2 kmol/h seems to give a satisfactory performance. 4.7. Configuration d. The high energy consumption of the cocurrent side reactor configuration is a well-

Figure 10. Effect of the pump-around position for configuration c. The bottom compositions are plotted against the stage from which pump-around is withdrawn.

Figure 11. Effect of the pump-around flow rate for configuration c. The bottom compositions are plotted against the pump-around flow rate.

Figure 12. Comparison between configurations c and d, plotted as the bottoms composition vs pump-around position.

known feature of that configuration. Further, it is known that this energy consumption can be reduced by using a countercurrent side reactor configuration.10,11 To demonstrate this, a comparison between countercurrent and cocurrent reactor configurations is given in Figure 12, showing bottoms composition vs location of pump-around. For simplicity, the other parameters, i.e., the catalyst loading and the distance between sidedraw and feedback of the side reactor, are the same. The feed tray is located at a tray where feedback is coming from the side reactor. The reboiler duty and pump-around

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Table 5. Comparison of the Results for the Best of Configurations a-c config. a

config. b

config. c

components

aqueous phase

organic phase

bottoms

aqueous phase

organic phase

bottoms

aqueous phase

organic phase

bottoms

AcH BuOH BuAc water DBE temp (K) flow rate (kmol/h)

0.0001 0.0130 0.0017 0.9852 0.0000 362.87 0.0085

0.0001 0.3746 0.2978 0.3273 0.0001 362.87 0.0300

0.0001 0.0013 0.9916 0.0000 0.0069 384.43 0.0131

0.0002 0.0126 0.0017 0.9855 0.0000 363.13 0.0085

0.0009 0.3639 0.3164 0.3186 0.0002 363.13 0.0300

0.0001 0.0003 0.9919 0.0000 0.0077 384.47 0.0130

0.0047 0.0119 0.0020 0.9814 0.0000 368.60 0.0085

0.0195 0.3161 0.3587 0.3055 0.0001 368.60 0.2230

0.0015 0.0000 0.9883 0.0000 0.0102 384.47 0.0130

Table 6. Operating Parameters for the Best of Configurations a-d parameter

config. a

config. b

config. c

config. d

reboiler duty (kW) catalyst loading (kg of catalyst/tray or kg of catalyst) feed tray location catalytic section/pump-around location pump-around flow rate (kmol/h)

0.5 0.25 11th tray 12-21 trays

0.5 0.25 8th tray 9-17 trays

3.0 4.0 6th tray 6-7 trays 0.2

1.0 4.0 6th tray 6-7 trays 1.0

includes operating conditions for configuration d. Because of very high energy requirements, configuration c is not suitable. Configuration d, on the other hand, seems to be promising and has a scope for further improvement upon rigorous optimization. The stoichiometric feed composed of AcH and a slight excess of BuOH was also studied in a similar manner. Again, configuration b was found to give a very good performance. Thus, all in all, out of the configurations studied, configuration b was found to be more effective than configurations a and c, whereas configuration d asked for detailed investigations. 6. Conclusion Figure 13. Effect of the reboiler heat duty for configuration d. The bottom compositions are plotted against the stage where pump-around is fed back.

flow rate were adjusted as described in the previous section. From Figure 12, it is concluded that a very similar product concentration can be obtained from the countercurrent side reactor with much lower energy input. Nevertheless, the energy requirement for configuration d is still higher than that for the other configurations (Table 6). However, we expect that further reduction is possible if all design variables are optimized simultaneously. A first step in that direction is taken by Baur and Krishna.11 However, such a detailed study is beyond the scope of this work. Finally it is worth noting that also for the countercurrent side reactor configuration multiple steady states will arise (Figure 13). This will also introduce an additional complexity in the determination of an optimal set of design parameters. 5. Discussion Table 5 gives a comparison of the results for the optimized configurations a-c found from the above simulation-based analysis. Table 6 shows the corresponding operating conditions. As can be seen from Table 5, configuration b gives the highest purity of BuAc (99.19 mol %) in the product accompanied by 0.78 mol of DBE. For configurations a and b, more than 99% AcH conversion is obtained, whereas configuration c gives around 98% conversion. For comparison, Table 6 also

In this paper, the conceptual process design for butyl acetate synthesis was studied. It was found that a column configuration with a reactive middle section and nonreactive stripping and rectifying sections is most suitable. Optimum values for various design parameters, like reboiler duty, catalyst loading, catalyst section length and location, and feed tray location, were determined using a one-parameter continuation technique, where only one parameter is varied at the same time. This provides useful insight into the qualitative behavior. However, it is not possible to get the “global optimal design” by this technique. Further improvement of the solution is possible by applying more advanced optimization techniques and by looking at more advanced solutions, with nonuniform distribution of the catalyst and multiple side reactors, for example. This, however, is beyond the scope of this work and will be subject to future work. Another interesting question that arises is related to steady-state multiplicity, which was observed for the configuration with the pump-around reactor over a wide range of operating conditions. In contrast to this, for the other configurations, multiplicity was only found in a small range of operating conditions outside the region of interest considered in the present paper. The physical sources of multiplicity for the pump-around reactor are not yet clear and also require further investigations. It is an interesting question whether steady-state multiplicity is a rather common property of pump-around reactor configurations.

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Acknowledgment

Literature Cited

The financial support of the German Science Foundation (DFG) under Grant KI 417/1 is greatly acknowledged. The authors thank one of the anonymous reviewers for drawing their attention to configuration d.

(1) Gangadwala, J.; Mankar, S.; Mahajani, S.; Kienle, A.; Stein, E. Esterification of acetic acid with butanol in the presence of ionexchange resins as catalysts. Ind. Eng. Chem. Res. 2003, 42, 21462155. (2) Janowsky, R.; Groebel, M.; Knippenberg, U. Nonlinear dynamics in reactive distillation-phenomena and their technical use; Final Report FKZ 03D0014 B0; Huels Infracor GmbH: experSCience, Marl, 1997. (3) Hanika, J.; Kolena, J.; Smejkal, Q. Butyl acetate via reactive distillation: Modeling and Experiments. Chem. Eng. Sci. 1999, 54, 5205-5209. (4) Sundmacher, K.; Kienle, A. Reactive DistillationsStatus and Future Trends; Wiley-VCH: Weinheim, Germany, 2003. (5) Hartig, H.; Regner, H. Verfahrenstechnische Auslegung einer Veresterungskolonne. Chem.-Ing.-Tech. 1971, 43, 10011007. (6) Venimadhavan, G.; Malone, M. F.; Doherty, M. F. A novel distillate policy for batch reactive distillation with application to the production of butyl acetate. Ind. Eng. Chem. Res. 1999, 38, 714-722. (7) Zhicai, Y.; Xianbao, C.; Jing, G. Esterification-distillation of butanol and acetic acid. Chem. Eng. Sci. 1998, 53, 2081-2088. (8) Bessling, B.; Welker, R.; Knab, J. W.; Lohe, B.; Disteldorf, W. Continuous preparation of esters and apparatus therefore. Ger. Offen. 1999, 6 (Chem. Abstr. 2003, 130, 11832v). (9) Mangold, M.; Kienle, A.; Mohl, K. D.; Gilles, E. D. Nonlinear computation using DIVA-Methods and applications. Chem. Eng. Sci. 2000, 55, 441-454. (10) Schoenmakers, H.; Buehler, W. Distillation column with external reactorssan alternative to the reaction column. Ger. Chem. Eng. 1982, 5, 292-296. (11) Baur, R.; Krishna, R. Distillation column with reactive pump arounds: an alternative to reactive distillation. Chem. Eng. Process. 2003, in press. (12) Kienle, A.; Mohl, K. D.; Gilles, E. D. Nichtlineare Dynamik bei Reaktivdestillationsprozessen. In Nichtlineare Dynamik bei chemischen Prozessen; DECHEMA eV: Frankfurt, Germany, 1997; Vol. B2, pp 1-6.

Nomenclature ai ) activity of the ith component in the liquid phase a′i ) product of the adsorption constant and activity for component i Ka ) equilibrium constant of esterification KDBE ) forward reaction rate constant for etherification, kmol/kg‚s Kf ) forward reaction rate constant for esterification, kmol/ kg‚s Ks,i ) adsorption constant of component i Mcat. ) mass of the catalyst, kg or kg of catalyst/tray n ) molar holdup, kmol P ) total pressure, Pa Qr ) reboiler heat duty, kW R ) gas constant, kJ/kmol‚K r1i ) esterification rate for component i, kmol/s r2i ) etherification rate for component i, kmol/s T ) temperature, K xi ) mole fraction of the ith component in the liquid phase Greek Letters γi ) activity coefficient of the ith component νi ) stoichiometric coefficient of the ith component List of Abbreviations AcH ) acetic acid BuAc ) n-butyl acetate BuOH ) n-butanol DBE ) di-n-butyl ether FC ) feed composition FTL ) feed tray location RDC ) reactive distillation column

Received for review December 11, 2002 Revised manuscript received September 17, 2003 Accepted October 27, 2003 IE021011Z