Synthesis and Hydrolysis of Methyl Acetate by Reactive Distillation

Although invented in 1921,3 the industrial application of reactive distillation took ... The reboiler duty is controlled with a transformer to within ...
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Ind. Eng. Chem. Res. 2001, 40, 1566-1574

Synthesis and Hydrolysis of Methyl Acetate by Reactive Distillation Using Structured Catalytic Packings: Experiments and Simulation Tim Po1 pken, Sven Steinigeweg, and Ju 1 rgen Gmehling* Carl von Ossietzky-Universita¨ t Oldenburg, Technische Chemie, P.O. Box 2503, D-26111 Oldenburg, Germany

Reactive distillation experiments have been performed for the synthesis and hydrolysis of methyl acetate using the structured catalytic packing Katapak-S with an acidic ion-exchange resin (Amberlyst 15) as the heterogeneous catalyst. Three different setups were used: one-feed and two-feed setups for the methyl acetate synthesis and another two-feed setup for the hydrolysis reaction. The influence of several variables, such as reaction kinetics, separation efficiency, residence time distribution, and heat loss, on the simulation results has been studied. Most important is the use of an adequate reaction kinetics, that is, an adsorption-based kinetic model that takes into account the selective swelling of the polymeric catalyst. Only using this model can the synthesis as well as the hydrolysis experiments be simulated within experimental accuracy. With pseudohomogeneous kinetics, only the experiments performed for the synthesis of methyl acetate can be modeled. An equilibrium-stage model is capable of describing the experiments in the packed column when the reaction kinetics, separation efficiency, and heat loss of the column are taken into account. Introduction Reactive distillation (RD) has received increasing attention in the last two decades because of its high potential for process intensification for certain types of reactions. The most important among these reactions are esterifications, hydrolysis reactions, transesterifications, and etherifications, whose maximum conversion is limited by chemical equilibrium. The methyl acetate synthesis and hydrolysis is of some commercial interest1,2 and serves as a model reaction for research in reactive distillation. Although invented in 1921,3 the industrial application of reactive distillation took place in the 1980s,4 i.e., 60 years later. Reactive distillation for the synthesis of methyl acetate has also been investigated experimentally by Sawistowski and Pilavakis,5 Agreda et al.,4,6 Krafczyk and Gmehling,7 Bessling et al.,8 and Kreul et al.9 For the special application of acetic acid removal from wastewater, investigations by Neumann and Sasson10 and Xu et al.11 are available. The hydrolysis was first investigated by Fuchigami,1 followed by Han et al.,12 Ge et al.,13 and Wang et al.14,15 Whereas in the first investigations,4-6 homogeneous catalysis using sulfuric acid was employed, all other investigations were performed using ion-exchange resins as the heterogeneous catalyst. Bessling et al., for example, used a catalytically active random packing made of 9-mm Raschig rings16 that consisted of porous glass impregnated with an acidic ion-exchange resin. Fuchigami employed selfmade pellets consisting of ion-exchange resin and PE powder with a diameter of 7 mm. In the work of Wang et al., only one stage in the middle of the column was equipped with ion-exchange resin as a packed-bed reactor. The only work available using structured catalytic packings for the synthesis of methyl acetate is that of Kreul et al. who used a packing similar to Katapak-S (Multipak by Montz) to carry out batch distillation experiments. * Author to whom correspondence should be addressed.

In the literature, many design and simulation methods can be found. Unfortunately, experimental data, which are needed for the verification of the results obtained by simulation or design methods, are scarce, especially for structured catalytic packings. Therefore, experiments were performed for the methyl acetate synthesis and hydrolysis reaction. For the measurements, a 50-mm column equipped with a laboratory version of the structured packing Katapak-S (Sulzer Chemtech Ltd.),17 which is made of corrugated wiremesh sheets, was used. Any heterogeneous catalyst with a size between 0.5 and about 2 mm can be immobilized between two of these sheets, forming a sandwich-like structure. In this work, an acidic ion-exchange resin (Amberlyst 15) was used. The experimental data will be used in developing criteria for the scale-up of the reactive distillation apparatus with Katapak-S.18 In parallel with the experiments, simulations were performed using the commercially available steady-state simulation environment Aspen Plus (version 10.1, Aspen Technologies, Cambridge, MA). One aim of this work was to determine whether the equilibrium-stage concept is sufficient to adequately describe the behavior of a packed reactive distillation column. Important parameters are identified by a comparison of the experimental findings with the simulation results. In the comparison, the influence of the reaction kinetics, residence-time distribution, separation efficiency of the reactive packing, and heat loss of the column were considered. For all of these parameters, independent experiments were performed. Thus, the simulation itself contained no adjustable parameters. Special attention was paid to the kinetics of the chemical reaction. The results of two different kinetic models were compared with the results obtained assuming chemical equilibrium throughout the reactive section of the column. Experimental Section All experiments were performed in a glass column with an inner diameter of 50 mm, equipped with a

10.1021/ie0007419 CCC: $20.00 © 2001 American Chemical Society Published on Web 02/16/2001

Ind. Eng. Chem. Res., Vol. 40, No. 6, 2001 1567

Figure 1. Schemes of the three setups used in this work. From left to right: one-feed setup, two-feed synthesis setup, and two-feed hydrolysis setup.

condenser and a reflux splitter. The condenser duty was calculated from measurements of the inlet and outlet water temperatures (to within (0.1 K) and the mass flow of the coolant ((1 g min-1). As a reboiler, a vessel with a low liquid holdup (approximately 1 dm3) was used to achieve stationarity in a reasonable period of time. The liquid was heated electrically with rod-shaped quartz heaters. The reboiler duty is controlled with a transformer to within (2%. The reactive section of the column consisted of Katapak-S elements (Sulzer Chemtech Ltd.), whereas in the nonreactive sections, BX packings by Sulzer Chemtech were used. The Katapak-S elements were filled with Amberlyst-15 (Rohm & Haas Co.), an acidic ion-exchange resin. Each column segment had a total height of 1.2 m and contained an effective packing height of 1 m. The column was insulated by a vacuum jacket. Flanges were insulated using mineral wool and aluminum foil. The heat loss of the column was measured as described below and taken into account during the simulation, as it was not negligible in the experiments with a very low reboiler duty. The column was controlled with commercially available process control equipment and software (Opto 22, Optoware; The Fix, Intellution). While the reboiler duty and the reflux ratio were set to a fixed value, the liquid level in the reboiler was kept constant to control the bottoms flow rate. All streams were measured by determining the mass flow using balances with an accuracy of (1 g h-1. The mass flow data of the feed streams were used to control the feed pumps (membrane pumps Gamma-4-RS, ProMinent) at the desired level. At the lower end of each column section, in the reboiler, and from the distillate stream, liquid samples were withdrawn by means of a syringe. Care was taken to avoid collecting any vapor in the syringe. The samples were cooled rapidly to 275 K and immediately analyzed by gas chromatography. The temperature was measured using Pt-100 thermometers with an accuracy of (0.1 K. Thermometers were installed at the lower end of each column section (except at the feed positions), in the reboiler, in the vapor rising into the condenser, and in the liquid reflux. The top and bottom pressures and the pressure drop were recorded using pressure transducers (Bosch, accuracy of (0.1%). All data were recorded by the process control system. The different setups used are shown in Figure 1. Below and above the reactive section (two-feed setup) or in the middle of the reactive section (one-feed setup), the liquid load of the column

was measured by recording the time needed to collect a specified amount (40 cm3) of liquid in a graduated vessel inside the column. The chemicals used were of reaction grade. Glacial acetic acid (99.5%) and methyl acetate (99.5%) were supplied by Celanese, methanol (99.5%) was supplied by Bayer, and the water used was deionized. The purity of all chemicals was verified by gas chromatography. All samples were analyzed by gas chromatography (HP 6890 with TCD; helium as the carrier gas at 2.8 cm3 min-1; HP-Innowax 30 m × 0.032 mm; split 30:1; temperature program ) 313.15 K for 4 min, heat at a rate of 65 K min-1 to 453.15 K, hold for 4.9 min; typical retention times ) 2.9 min for methyl acetate, 3.8 min for methanol, 5.7 min for water, and 8.0 min for acetic acid). (a) Separation Efficiency. Because it is known that the separation efficiency of structured wire-mesh packings diminishes with increasing surface tension of the liquid because of poorer wetting,19 the separation efficiency of the Katapak-S packing elements used was measured using the test system water-acetic acid with a column setup consisting of only two reactive sections with a total height of 2 m of Katapak-S. Between the two sections, the liquid load was measured as described above. Measurements were performed at total reflux using water-acetic acid mixtures of different compositions. The reboiler duty, and thus the liquid and vapor load, was varied in the range expected for the reactive experiments. After the column was allowed to equilibrate for approximately 2 h, liquid samples were withdrawn from the reboiler, the reflux, and between the two column sections. The liquid load between the two sections was also determined. The pressure was varied between 30 and 100 kPa. From the concentrations, the number of theoretical stages per meter (NTSM) were calculated using the known VLE of the system.20 The results are shown in Figure 2, along with a correlation as a function of the liquidphase mole fraction of the low boiler water at the lower end of the column segment. For the small range of liquid loads (approximately 0.5-4 m3 m-2 h-1, corresponding to F factors of about 0.1-1.5 Pa0.5) used in this work, the separation efficiency can be considered independent of the liquid load. Also, no significant dependence on the pressure was found. The following correlation for the dependence on the water content is obtained, which is applicable for 0.5 < xL,water < 0.97:

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Figure 2. Separation efficiency of Katapak-S packings for the test system water-acetic acid as a function of the liquid-phase mole fraction of water at the lower end of the column segment. Liquid loads were 1.5 m3/m2h (O) and 2.2 m3/m2h (0), corresponding to F factors ranged from 0.5 to 1 Pa0.5.

NTSMKatapak-S ) 0.75 ln(1 - xL,water) + 4.4

(1)

At water concentrations below a mole fraction of 0.5, a value for NTSM of 4 m-1 should be used. The separation efficiency at low to moderate water mole fractions is somewhat higher than that determined by the manufacturer for the 70-mm-diameter laboratory packing using the test system chlorobenzene-ethylbenzene (around 3 m-1).21 However, the higher NTSM value found in this work for the 50-mm-diameter laboratory packing correlates well with the higher total specific surface area available for mass transfer. Whereas for the 70-mm packing, the specific surface area is 356 m2 m-3 (including the packing and the inner column wall), the corresponding value for the 50-mm packing used in this work is 430 m2 m-3.22 (b) Heat Loss. By distilling pure components and measuring the condenser duty, the heat loss was determined by comparison with the set reboiler duty. Using different components (methyl acetate, bp 330.08 K; water, bp 373.15 K), the temperature difference across the column jacket was varied. It was found that the heat loss depends linearly on the temperature difference, as expected (Figure 3). Measuring the heat loss for different column setups allowed us to distinguish between the heat loss of individual column sections. The following relations hold for the column used in this work:

Q˙ loss,reboiler ) 2.0WK-1∆T

(2a)

Q˙ loss per segment ) 0.33WK-1∆T

(2b)

These relations were included in the simulation as described below. Simulation All simulations were carried out using the model RADFRAC from the commercial steady-state simulator Aspen Plus, version 10.1,23 which is based on a rigorous equilibrium-stage model for solving the MESH equations. Despite the fact that a column equipped with structured packings was used, very good results were obtained using this model instead of a rate-based model, as will be explained below. Stages are numbered from the top to the bottom, with stage 1 being the condenser and stage N being the

Figure 3. Heat loss for the two-feed column (0, synthesis setup with five sections; O, hydrolysis setup with four sections) as a function of the temperature difference across the column jacket.

reboiler. For each segment filled with BX packing, five theoretical stages were assumed, whereas for each Katapak-S segment, four theoretical stages were specified. This results in total stage numbers of N ) 10 for the one-feed setup, N ) 20 for the hydrolysis, and N ) 25 for the synthesis two-feed setup. For the simulation of the individual experiments, the reboiler duty and the reflux ratio were specified in addition to the feed rates and the column setup. In this way, the simulations are specified in the same way as the operating conditions of the column, and comparisons can be carried out directly. Key features of the simulation include the following: (a) VLE. For the vapor-liquid equilibrium calculations, the UNIQUAC model24 was used to describe the liquid-phase nonideality, and the Nothnagel equation25 was used to account for deviations from ideal behavior in the vapor phase. Such deviations are mainly caused by the dimerization of the acetic acid. Parameters for the Nothnagel equation were used from the Aspen Plus databases, whereas pure-component properties were taken from the from the Dortmund Data Bank,20 and temperature-dependent UNIQUAC parameters were fitted simultaneously to the different types of thermodynamic data (vapor-liquid equilibrium, heat of mixing, and activity at infinite dilution), also taken from the Dortmund Data Bank. These parameters have been published already, together with the kinetic data.26 During the fit of the UNIQUAC parameters, the vaporphase nonideality was taken into account. (b) Separation Efficiency. The known separation characteristics of the BX packing,27 which has a NTSM of 5 m-1 (at P ) 960 mbar and with an F factor of 1) were used to specify the number of stages in the nonreactive sections. The separation efficiency of the Katapak-S packing was measured using the binary test system water-acetic acid. This was included in the simulation by specifying a NTSM value of 4 m-1 (taken as the maximum value) and including a stage efficiency depending on the water content, as calculated from eq 1.

η)

min(4 m-1, NTSMKatapak-S) 4 m-1

(3)

Taking into account the influence of the water concentration on the separation efficiency showed only a small effect on the simulation results, because the water mole

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fraction was rarely above 0.5 in the reactive zone. Only for a few hydrolysis experiments was a significant influence observed. Nevertheless, the correlation was incorporated into the simulation. (c) Chemical Reaction. The chemical reaction kinetics have already been studied extensively.26 Parameters for two different kinetic models (pseudohomogeneous and adsorption-based) were suggested. Whereas in the pseudohomogeneous model, catalysis by solvated protons is assumed, the adsorption-based model takes into account the selective swelling of the polymeric catalyst. Both models will be compared with the assumption of chemical equilibrium on each stage. The pseudohomogeneous model used is

r)

1 1 dni ) k1aHOAcaMeOH - k-1aMeOAcaH2O mcat νi dt

(4)

with

k1 ) 2.961 × 104 mol gcat-1 s-1 exp(-49 190 J mol-1/RT) k-1 ) 1.348 × 106 mol gcat-1 s-1 exp(-69 230 J mol-1/RT) The adsorption-based model can be written as

r)

k1a′HOAca′MeOH - k-1a′MeOAca′H2O 1 1 dni ) mcat νi dt (a′HOAc + a′MeOH + a′MeOAc + a′H2O)2 (5)

with

a′i )

Kiai Mi

k1 ) 8.497 × 106 mol gcat-1 s-1 exp(-60 470 J mol-1/RT) k-1 ) 6.127 × 105 mol gcat-1 s-1 exp(-63 730 J mol-1/RT) and

KHOAc ) 3.15, KMeOH ) 5.64, KMeOAc ) 4.15, KH2O ) 5.24 The reaction is calculated on the basis of the catalyst mass on each stage. This allows for the neglect of the liquid holdup, which strongly depends on the operating conditions of the column, as the homogeneous reaction in the liquid phase is not important for the overall extent of reaction, as demonstrated in a previous paper.26 (d) Heat Loss. The heat loss of the column (eqs 2a and b) was included in the simulation model, because the aim was to accurately identify all influences on the column performance. (e) Residence Time Distribution (RTD). In the reactive distillation simulation using RADFRAC, by default, the liquid on each stage is assumed to be ideally mixed. With the measured NTSM values, this corresponds to a packing height of about 25-30 cm. The residence time behavior of Katapak-S has previously been measured by Go¨tze,28 Moritz and Hasse,21 and

Ellenberger and Krishna.29 Conditions close to plug flow for the liquid phase were found. Because the reaction is second-order, the RTD should have an influence on the extent of reaction. The importance of such an influence was assessed using the simulation by subdividing every stage into m “substages” with the catalyst mass and the vaporization efficiency being divided by m. In this way, the separation was affected only very slightly, but the RTD of each stage was that of a cascade of m ideal continuously stirred tank reactors. The effect on the simulated composition profile was found to be negligible for values of m between 2 and 10. Therefore, it was concluded that the RTD is of minor importance and can be neglected in further simulations. Moritz et al. came to the same conclusion.30 Results (a) One-Feed Column. To gain more insight into the complex interactions between reaction and separation that occur in reactive distillation processes, a column equipped with reactive packing only and one feed was used. This allows us to study the characteristics of this setup, which, from a research point of view, has certain advantages over the commercial two-feed hybrid (containing reactive and nonreactive sections) columns. With only one type of structured packing inside the column, the characteristics of this packing can be studied more easily, because this setup has fewer degrees of freedom than more complex setups. Furthermore, the reaction velocity can be changed by varying the column pressure, because the pressure directly influences the boiling temperature and, thus, the strongly temperature-dependent reaction kinetics. At the same time, the variation of the column pressure in a moderate range (20-100 kPa) has only a small effect on the separation efficiency of the laboratory packing in the range of operating conditions studied (see above and ref 30). This means that the extent of reaction in the column can be changed without replacing the catalyst. Additionally, a reactive one-feed column operated at high reflux ratios can be used to experimentally verify the concept of reactive residue curves. For the one-feed setup, 12 experiments were carried out. In Figure 4, the experimental and simulated liquidphase composition profiles are given for a typical experiment (run number 1-4), and in Table 1, the parameters for all experiments are listed. The reflux ratio, the pressure at the column top (P1), the feed flow rate (F˙ ), and the reboiler duty (Q˙ reb) were set, which resulted in the given distillate-to-feed ratio (D:F). Compositions are reported as mole fractions for the feed (xF), the distillate (xD), and the bottoms (xB). The resulting methanol conversion (XMeOH) is given, along with the column pressure drop (∆P) and the liquid load below the feed stage (wL,below feed). The temperatures of the feed (ϑF), the liquid reflux (ϑ1,L), the vapor rising into the condenser (ϑ2,V), the reboiler holdup (ϑ10,L), and the vapor rising from the reboiler (ϑ10,V) are also reported. In Figure 5, the methanol conversion is plotted versus the column head pressure. Experimental and simulation results show a good agreement if reaction kinetics are used. Clearly, the column does not operate at chemical equilibrium over the entire pressure range. Although the deviation between experiment and simulation is smallest for the simulation using adsorption-based kinetics, it is not clear from the results presented in

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Figure 4. Liquid-phase composition profiles for the one-feed column. Experimental data (4, MeOAc; 0, MeOH; ], H2O; O, HOAc) of run 1-4 (see Table 1) are compared with simulation results (- - -, MeOAc; - ‚ ‚ -, MeOH; - - -, H2O; s, HOAc) using adsorption-based kinetics.

Figure 5 alone whether pseudohomogeneous or adsorption-based kinetics describe the column performance better. (b) Methyl Acetate Synthesis Using a Two-Feed Column. A two-feed column entirely equipped with structured packings was built in the classical setup used by Eastman4 and in a previous work.7 Whereas in the Eastman process, homogeneous catalysis is used, with Katapak-S, the advantages of heterogeneous catalysis can be utilized. To develop a scale-up procedure for Katapak-S,30 several experiments were performed with the classical setup. A typical composition profile is given in Figure 6. As can be seen from the figure, the simulation results agree with the experimental composition profile, which is especially important in the range of the reactive zone (stages 11-19). Experimental details are given in Table 2. The reflux ratio, the top stage pressure (P1), the feed flow rates (F˙ HOAc and F˙ MeOH), and the reboiler duty (Q˙ reb) were set. The resulting distillate flow (D˙ ) and molar feed ratio are given in the table. Compositions are reported as mole fractions for the distillate (xD) and the bottoms (xB). The resulting conversion of the limiting reactant (X) is given, along with the column pressure drop (∆P) and the liquid loads above (wL,11) and below (wL,19) the reactive section. The temperatures of the feeds (ϑfeed), the liquid reflux (ϑ1,L), the vapor rising into the condenser (ϑ2,V), the liquid above (ϑ11,L) and in the middle (ϑ15,L) of the reaction zone, the reboiler holdup (ϑ25,L), and the vapor rising from the reboiler (ϑ25,V) are also reported. The concepts behind the setup have been described in detail already.4,6 Briefly, the system contains two binary azeotropes (methyl acetate-methanol and methyl acetate-water). The former is broken by the chemical reaction, whereas the latter is broken by extraction of the water by the acetic acid feed. Pure, dry acetic acid must be used, and in order to achieve a high purity, the addition of a small amount of acetic anhydride to the acetic acid feed is necessary to remove the last traces of water.31 In this work, the aim was not to produce pure

methyl acetate but rather to judge which kinetic expression is best-suited for the considered reaction and which complexity of the simulation model is sufficient. Therefore, no attempt was made to use dry acetic acid. The acid used contained about 0.3 wt % water. This results in a limited extraction capacity of the acetic acid feed, which in turn results in water contamination of the distillate because of the methyl acetate-water azeotrope. Also, the reactive zone was still too short to achieve complete conversion, and thus, the distillate always contained a little methanol. Nevertheless, one attempt was made to achieve high conversion and high distillate purity simultaneously. This was done to test whether the optimum conditions determined from the simulation can be reproduced experimentally. Using a 56% stoichiometric excess of acetic acid, a methanol conversion of 97.1% was achieved. The purity of the methyl acetate was 97.5 mol %, and the bottom stream was almost free from methyl acetate and methanol (see Table 2, experiment S-19). Simulation results and experimental data show excellent agreement for this run. The influence of relevant parameters on the results is shown in Figures 7-9. In Figure 7, a series of experiments are shown in which the total feed has been varied while the stoichiometric ratio and reflux ratio have been kept constant. As expected, the simulation assuming chemical equilibrium on every theoretical stage shows no dependence on the liquid throughput. The decline at very low total feed is due to the heat loss of the column, which is on the same order of magnitude as the reboiler duty at these very low feed flow rates. It is obvious that only at very low flow rates can the column be modeled with the assumption of chemical equilibrium on every reactive stage. Even at moderate liquid loads (F˙ ) 60 mol h-1 corresponding to approximately 3.5 m3 m-2 h-1 at these operating conditions), there is a significant difference between the results of the simulation using the simplifying assumption of chemical equilibrium and the simulation taking into account the kinetics of the chemical reaction. Only the latter is capable of describing the dependence of the conversion observed experimentally on the throughput, whereas the type of the kinetic model has only a small influence. When the reflux ratio is varied, as shown in Figure 8, a characteristic maximum in conversion around a value of about 1.8 is found. The behavior is similar for all models considered, although the maximum is much more pronounced when reaction kinetics rather than the assumption of chemical equilibrium are used. Agreda at al.4 and Bessling et al.8 found a similar maximum at reflux ratios of about 1.7 and 1.5, respectively. An explanation for this phenomenon can be derived from the separation characteristics of the column. At low reflux ratios, insufficient separation of the products from the reaction zone limits the conversion. Additionally, acetic acid can be found in the distillate,8 which causes decreasing conversion. At high reflux ratios, the reactants are separated too effectively from each other, thus limiting conversion is observed. For optimum reaction conditions, the concentrations of both reactants, i.e., methanol and acetic acid, must be high in the reaction zone. If a high reflux ratio is used, methanol leaves the column with the distillate, because the methyl acetatemethanol azeotrope is enriched at the top of the column. The optimal reflux ratio represents a compromise

Ind. Eng. Chem. Res., Vol. 40, No. 6, 2001 1571 Table 1. Experimental Data for the Methyl Acetate Synthesis Using the One-Feed Setupa run number

1-1

1-2

1-3

1-4

1-5

1-6

1-7

1-8

1-9

1-10

1-11

1-12

reflux ratio P1 (mbar) F˙ (mol/h) D:F (mol/mol) xF (HOAc) xF (MeOH) xF (MeOAc) xF (H2O) xD (MeOH) xD (MeOAc) xD (H2O) xB (HOAc) xB (MeOH) xB (MeOAc) xB (H2O) XMeOH (%) ∆P (mbar) Q˙ reb (W) wL,below feed (m/h) ϑ1,L (°C) ϑ2,V (°C) ϑF (°C) ϑ10,V (°C) ϑ10,L (°C)

1 247 149 0.41 0.517 0.473 0.000 0.010 0.648 0.352 0.000 0.623 0.155 0.014 0.208 25.6 3.5 934 20.6 22.6 46.6 50.6 55.6

1 297 149 0.31 0.517 0.473 0.000 0.010 0.545 0.455 0.000 0.608 0.143 0.016 0.233 43.5 3.1 1049 3.7 24.6 26.5 51.4 54.9 60.6

1 298 132 0.30 0.560 0.430 0.000 0.010 0.525 0.475 0.000 0.605 0.130 0.018 0.247 42.4 2.5 419 4.3 24.3 25.4 46.3 54.3 60.6

1 399 132 0.32 0.560 0.430 0.000 0.010 0.459 0.541 0.000 0.570 0.147 0.030 0.253 43.0 2.1 934 3.5 57.3 63.9 66.9

1 501 136 0.27 0.546 0.444 0.000 0.010 0.417 0.583 0.000 0.520 0.157 0.048 0.275 49.4 1.5 880 35.4 35.4 50.9 58.7 67.2

1 702 134 0.25 0.554 0.436 0.000 0.010 0.381 0.619 0.000 0.494 0.114 0.068 0.324 58.8 1.5 847 3.3 63.1 -

1 296 95 0.43 0.509 0.481 0.000 0.010 0.558 0.442 0.000 0.585 0.114 0.009 0.292 36.9 2.3 934 3.3 24.3 26.2 50.5 57.1 62.2

1.5 308 138 0.34 0.560 0.430 0.000 0.010 0.576 0.424 0.000 0.666 0.063 0.008 0.262 45.2 7.0 1568 4.9 24.5 26.5 47.9 64.7 67.4

2 301 138 0.42 0.560 0.430 0.000 0.010 0.651 0.349 0.000 0.708 0.010 0.000 0.282 35.6 10.2 1920 4.9 25.0 28.2 50.5 71.6 72.6

0.5 1024 100 0.38 0.459 0.481 0.014 0.046 0.482 0.500 0.017 0.354 0.105 0.089 0.453 48.2 0.8 934 53.4 54.7 80.7 69.9 81.4

1 1023 103 0.35 0.440 0.475 0.008 0.078 0.410 0.584 0.006 0.322 0.127 0.049 0.502 52.3 1.1 934 3.3 52.8 53.4 79.9 70.8 83.1

3 1022 45 0.46 0.437 0.457 0.016 0.090 0.347 0.647 0.006 0.244 0.003 0.000 0.753 64.5 0.9 1169 1.9 51.7 53.0 80.5 89.7 95.1

a

Numbers denote the corresponding stage. See text for a detailed description.

Figure 5. Methanol conversion for the one-feed column versus column head pressure. Experimental results (O) are compared with simulations performed using different kinetics (s, adsorptionbased; - - -, pseudohomogeneous) or the assumption of chemical equilibrium on each theoretical stage (- - -).

between these two cases. This phenomenon is not limited to two-feed columns but can also occur in onefeed columns.32 The maximum is more pronounced when kinetics are used, because the reaction velocity depends on the product of the reactant mole fractions. If the concentration profile in the column is unfavorable for reaction, not only is the attainable conversion low but so are the reaction rates. Figure 9 shows the results for the effects of a variation of the molar feed ratio of acetic acid to methanol on the acetic acid conversion. All models overestimate the acetic acid conversion slightly while showing the correct dependence on the molar feed ratio. The simulation using chemical equilibrium gives results similar to those obtained using kinetics at high reactant ratios. At ratios of unity and below, conversion is clearly overestimated by assuming chemical equilibrium. (c) Methyl Acetate Hydrolysis Using a Two-Feed Column. Because a comprehensive model for the methyl acetate system must be capable of describing not only the synthesis but also the hydrolysis and because of the

Figure 6. Liquid-phase composition profiles for the synthesis of methyl acetate using the two-feed setup. Experimental data (4, MeOAc; 0, MeOH; ], H2O; O, HOAc) of run S-1 (see Table 2) are compared with simulation results (- - -, MeOAc; - ‚ ‚ -, MeOH; - -, H2O; s, HOAc) using adsorption-based kinetics.

fact that the hydrolysis is of greater commercial interest than the synthesis, hydrolysis experiments were also carried out. Here, mainly the molar feed ratio of water to methyl acetate was varied. The data of the seven corresponding experiments are shown in Table 3. The reflux ratio, the top-stage pressure (P1), the feed flow rates (F˙ H2O and F˙ MeOAc, with the “methyl acetate” feed being premixed to contain methanol and water), and the reboiler duty (Q˙ reb) were set. The resulting distillate flow (D˙ ) and molar feed ratio are given in the table. Compositions are reported as mole fractions for the methyl acetate feed (xF), the distillate (xD), and the bottoms (xB). The methyl acetate conversion (XMeOAc) is given, to-

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Table 2. Experimental Data for the Methyl Acetate Synthesis Using the Two-Feed Setupa run number

S-1

S-2

S-3

S-4

S-5

S-6

S-7

S-8

S-9

S-10

S-11

S-12

S-13

S-14

S-15

S-16

S-17

S-18

S-19

reflux ratio P1 (mbar) F˙ HOAc (mol/h) F˙ MeOH (mol/h) D˙ (mol/h) molar ratio HOAc:MeOH xD (HOAc) xD (MeOH) xD (MeOAc) xD (H2O) xB (HOAc) xB (MeOH) xB (MeOAc) xB (H2O) X (%) ∆P (mbar) Q˙ reb (W) wL,11 (m/h) wL,19 (m/h) ϑfeed HOAc (°C) ϑfeed MeOH (°C) ϑ1,L (°C) ϑ2,V (°C) ϑ11,L (°C) ϑ15,L (°C) ϑ25,V (°C) ϑ25,L (°C)

2.1 1017 30.8 32.2 30.1 0.96

2.1 1024 11.1 10.2 9.6 1.09

2.1 1011 54.4 52.7 61.8 1.03

2 1008 107.7 108.1 75.9 1.00

0.477 986 32.6 31.4 51.9 1.04

1 992 32.2 29.2 31.7 1.10

4.4 993 32.1 32.1 30.7 1.00

2 1010 24.2 51.8 36.6 0.47

2 1022 31.0 40.3 36.4 0.77

2 1018 37.4 23.5 30.1 1.59

2 1017 40.9 25.5 32.0 1.60

2.1 1028 41.3 20.3 28.3 2.03

2 997 84.0 63.0 56.1 1.33

2 988 26.9 27.7 23.6 0.97

2 1020 30.4 34.0 35.6 0.89

2 1020 12.3 19.5 20.7 0.63

3 1006 34.5 33.2 32.8 1.04

1 1021 11.8 10.6 12.1 1.11

2.2 1010 24.0 37.4 23.5 1.56

0.000 0.107 0.877 0.016 0.104 0.070 0.002 0.824 87.6 2.98 1105 3.5 2.1 21.1 20.1 54.9 55.4 65.6 67.5 83.1 91.7

0.000 0.028 0.966 0.006 0.175 0.013 0.001 0.812 92.4 1.42 519 19.6 19.7 55.7 56.5 66.1 61.5 97.5 99.6

0.000 0.107 0.663 0.230 0.191 0.042 0.001 0.766 85.2 6.22 1728 6.7 3.5 18.2 18.4 54.0 54.8 65.2 70.8 94.4 97.1

0.000 0.079 0.907 0.014 0.257 0.236 0.004 0.503 65.0 12.94 2488 10.5 7.3 26.9 26.8 54.8 55.4 72.4 77.9 80.0 87.4

0.198 0.191 0.399 0.211 0.043 0.000 0.000 0.957 68.6 1.56 1081 1.8 1.1 20.6 20.7 62.7 63.3 90.1 84.6 99.3 99.8

0.009 0.065 0.846 0.080 0.108 0.058 0.001 0.833 92.4 1.88 847 2.2 1.7 20.4 20.3 55.5 58.7 87.2 81.2 86.7 92.6

0.000 0.196 0.786 0.018 0.214 0.029 0.001 0.756 77.8 5.52 1551 5.2 3.2 21.1 20.9 53.3 53.8 60.7 60.0 93.0 97.7

0.000 0.392 0.603 0.006 0.021 0.389 0.000 0.590 93.4 2.49 1190 20.7 20.3 53.3 53.2 69.2 71.8 70.9 76.6

0.000 0.263 0.734 0.003 0.059 0.033 0.000 0.908 94.3 3.15 1544 2.8 2.2 20.2 20.5 53.3 53.8 66.9 73.8 90.6 95.2

0.008 0.087 0.667 0.238 0.533 0.000 0.000 0.467 88.1 3.08 1678 3.1 2.3 23.9 24.5 56.7 64.0 98.8 97.0 104.2 105.2

0.005 0.056 0.731 0.207 0.482 0.000 0.004 0.514 93.4 2.83 1631 3.3 1.9 20.9 21.4 56.6 63.5 97.9 95.3 103.6 104.8

0.019 0.055 0.638 0.288 0.667 0.000 0.000 0.333 91.1 2.76 1292 3.3 2.2 20.4 20.4 59.3 72.0 100.4 97.8 106.1 107.5

0.000 0.072 0.915 0.013 0.345 0.071 0.006 0.578 83.1 8.08 2090 7.3 4.6 26.2 26.5 55.0 55.8 71.6 74.6 92.8 97.4

0.000 0.047 0.945 0.009 0.143 0.124 0.004 0.729 83.9 2.66 988 2.9 1.7 23.7 24.0 54.3 55.1 65.9 66.0 78.8 86.9

0.000 0.197 0.792 0.011 0.097 0.005 0.000 0.898 89.4 2.88 1358 3.3 2.0 21.3 21.7 54.0 54.5 66.1 72.8 99.5 100.3

0.000 0.429 0.567 0.003 0.063 0.000 0.000 0.937 89.8 0.83 988 0.9 0.6 21.7 22.2 52.4 53.6 79.7 87.0 100.4 100.8

0.000 0.235 0.729 0.036 0.268 0.000 0.003 0.729 75.9 2.56 933 2.7 2.1 21.1 21.5 53.4 54.0 94.5 92.6 101.1 101.8

0.004 0.021 0.839 0.136 0.153 0.000 0.001 0.846 96.6 0.92 463 0.7 0.4 24.6 24.8 54.1 56.2 95.8 91.0 100.9 101.5

0.000 0.018 0.975 0.007 0.367 0.007 0.001 0.624 97.1 2.68 919 3.7 2.0 22.0 22.2 56.4 56.1 70.4 68.7 100.1 102.3

a

Numbers denote the corresponding stage. See text for a detailed description.

Figure 7. Conversion versus total feed for the synthesis of methyl acetate using the two-feed column. Experimental results (O) are compared with simulations (v ) 2, D:F ) 0.48, F˙ MeOH ) F˙ HOAc) using different kinetics (s, adsorption-based; - -, pseudohomogeneous) or the assumption of chemical equilibrium on each theoretical stage (- - -).

Figure 8. Conversion versus reflux ratio for the synthesis of methyl acetate using the two-feed column. Experimental results (O) are compared with simulations (D:F ) 0.48, F˙ MeOH ) F˙ HOAc ) 32.5 mol h-1) using different kinetics (s, adsorption-based; - -, pseudohomogeneous) or the assumption of chemical equilibrium on each theoretical stage (- - -).

gether with the column pressure drop (∆P) and the liquid loads above (wL,6) and below (wL,14) the reactive section. The temperatures of the feeds (ϑfeed), the liquid reflux (ϑ1,L), the vapor rising into the condenser (ϑ2,V), the liquid in the middle (ϑ10,L) of the reaction zone, the reboiler holdup (ϑ20,L), and the vapor rising from the reboiler (ϑ20,V) are also reported. Whereas in most experiments, a very low distillate flow rate was used, one run (number 7) was conducted with a D:F ratio of 0.43. As a typical concentration profile along the column, the results for run number 3 are shown in Figure 10. The simulated and experimental profiles are obviously in a good agreement. In Figure 11, the methyl acetate conversions for several molar feed ratios of water to methyl acetate are shown. Because the chemical equilibrium conversion is very low for a stoichiometric mixture (28% at 50 °C), it would be a reasonable approach to use an excess of water to increase the methyl acetate conversion. The

drawback is a large amount of water in the bottom stream, which has to be separated from the products methanol and acetic acid. This set of experiments permits us to draw the conclusion that the adsorption-based kinetic model is the superior one, because the flattening of the curve at high molar feed ratios is described adequately only with the help of this model. The strong uptake of water and methanol into the ion-exchange resin versus the weak uptake of methyl acetate decreases the overall reaction rate. Raising the water content in the column, therefore, does not increase the conversion as strongly as expected from pseudohomogeneous kinetics. Conclusions A large number of experiments for various experimental setups have been performed in order to develop a reliable framework for simulating reactive distillation

Ind. Eng. Chem. Res., Vol. 40, No. 6, 2001 1573

Figure 9. Acetic acid conversion versus molar feed ratio for the synthesis of methyl acetate using the two-feed column. Experimental results (O) are compared with simulations (v ) 2, D:F ) 0.48, F˙ MeOH + F˙ HOAc ) 65 mol h-1) using different kinetics (s, adsorption-based; - -, pseudohomogeneous) or the assumption of chemical equilibrium on each theoretical stage (- - -). Table 3. Experimental Data for the Methyl Acetate Hydrolysis Using the Two-Feed Setupa run number

H-1

H-2

H-3

H-4

H-5

H-6

H-7

P1 (mbar) F˙ MeOAc (mol/h) F˙ H2O (mol/h) D˙ (mol/h) xF,MeOAc (HOAc) xF,MeOAc (MeOH) xF,MeOAc (MeOAc) xF,MeOAc (H2O) molar ratio H2O:MeOAc xD (HOAc) xD (MeOH) xD (MeOAc) xD (H2O) xB (HOAc) xB (MeOH) xB (MeOAc) xB (H2O) XMeOAc (%) ∆P (mbar) Q˙ reb (W) wL,6 (m/h) wL,14 (m/h) ϑfeed MeOAc (°C) ϑfeed H2O (°C) ϑ1,L (°C) ϑ2,V (°C) ϑ10,L (°C) ϑ20,V (°C) ϑ20,L (°C)

1023 12.8 9.3 0.34 0.000 0.008 0.816 0.167 1.09

1026 15.8 17.0 0.31 0.010 0.116 0.753 0.121 1.59

1026 15.3 21.4 0.30 0.010 0.116 0.753 0.121 2.02

1023 14.4 24.4 0.43 0.007 0.117 0.756 0.120 2.40

1017 16.1 40.7 0.31 0.009 0.115 0.763 0.114 3.46

1016 16.2 57.8 0.33 0.010 0.116 0.753 0.121 4.91

1005 15.6 9.5 10.77 0.009 0.115 0.762 0.114 0.95

0.000 0.179 0.579 0.060 0.171 0.175 0.298 0.355 35.7 1.05 247 1.0 2.2 15.8 15.9 50.1 51.5 57.8 58.8 63.7

0.000 0.359 0.638 0.002 0.151 0.200 0.214 0.434 40.1 1.07 314 1.6 3.0 19.8 20.3 53.0 53.4 58.6 59.5 65.3

0.000 0.351 0.646 0.003 0.168 0.182 0.168 0.482 47.3 1.07 314 1.7 3.1 21.1 20.4 52.5 53.4 59.0 60.1 66.3

0.000 0.354 0.634 0.011 0.146 0.185 0.131 0.538 50.8 1.12 367 1.9 3.3 19.2 19.7 53.6 53.2 59.0 61.2 68.6

0.000 0.324 0.662 0.014 0.123 0.150 0.087 0.640 55.1 1.27 425 2.7 4.1 19.7 20.4 53.0 53.2 59.4 61.9 71.1

0.000 0.194 0.774 0.032 0.096 0.122 0.064 0.718 57.7 1.15 367 2.5 3.9 19.7 20.1 53.7 54.0 60.4 63.3 72.1

0.000 0.149 0.800 0.051 0.202 0.189 0.043 0.565 23.4 1.4 511 2.1 3.5 19.6 19.9 54.1 54.2 56.1 64.6 79.5

a Numbers denote the corresponding stage. See text for a detailed description.

that is capable of modeling the methyl acetate synthesis as well as the hydrolysis reaction with the required accuracy. From a comparison of the experimental results with simulation results using the Aspen RADFRAC model, it has been shown that only with an adequate, that is, adsorption-based, expression for the rate of the chemical reaction can the experiments be described within the experimental accuracy. It was also found that the equilibrium-stage concept is sufficient to model the column behavior. However, the heat loss of the mini-plant column, which can be measured easily with the required accuracy, had to be included in the simulation. The residence time distribution has no significant influence. The difference in the composition profile obtained by assuming ideally mixed equilibrium stages or a cascade of two or more ideally mixed tank reactors per equilibrium stage is negligible.

Figure 10. Liquid-phase composition profiles for the hydrolysis of methyl acetate using the two-feed column. Experimental data (4, MeOAc; 0, MeOH; ], H2O; O, HOAc) of run H-3 (see Table 3) are compared with simulation results (- - -, MeOAc; - ‚ -, MeOH; - - -, H2O; s, HOAc) using adsorption-based kinetics.

Figure 11. Methyl acetate conversion versus molar feed ratio for the hydrolysis of methyl acetate using the two-feed column. Experimental results (O) are compared with simulations (v ) 50, D ) 20 g h-1, F˙ MeOAc ) 1 kg h-1) using different kinetics (s, adsorption-based; - -, pseudohomogeneous) or the assumption of chemical equilibrium on each theoretical stage (- - -).

Acknowledgment We thank the “Fonds der Chemischen Industrie” for providing a scholarship to T.P. and Celanese and Bayer for providing the chemicals required for our measurements. Also, we express our thanks to Sulzer Chemtech for the packing elements and additional financial support. Literature Cited (1) Fuchigami, Y. Hydrolysis of Methyl Acetate in a Distillation Column Packed with Reactive Packing of Ion Exchange Resin. J. Chem. Eng. Jpn. 1990, 23, 354. (2) Agreda, V. H. Acetic Anhydride from Coal. CHEMTECH 1988, 4, 250. (3) Backhaus, A. A. Continuous Process for the Manufacture of Esters. U.S. Patent 1,400,849, 1921.

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Received for review August 10, 2000 Accepted January 7, 2001 IE0007419