Mastering the Reaction Is the Key to Successful Design of

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Ind. Eng. Chem. Res. 2008, 47, 6014–6024

PROCESS DESIGN AND CONTROL Mastering the Reaction Is the Key to Successful Design of Heterogeneously Catalyzed Reactive Distillation: A Comprehensive Case Study of Hexyl Acetate Synthesis Markus Schmitt,† Sergej Blagov,‡ and Hans Hasse* Institute of Thermodynamics and Thermal Process Engineering, UniVersity of Stuttgart, 70550 Stuttgart, Germany

This work condenses the results of extensive experimental and modeling and simulation work on heterogeneously catalyzed reactive distillation and relates them to basic questions that must be answered in designing that complex integrated process. As a test case, hexyl acetate synthesis was chosen. The results summarized and evaluated here make it one of the most thoroughly and comprehensively studied examples of reactive distillation processes in the literature. In well-documented reactive distillation experiments, carried out in two scales, all relevant process parameters were studied. Modeling and simulation of the reactive distillation process, carried out both with the equilibrium stage and with the rate-based approach, are based on careful experimental and modeling work on the underlying vapor-liquid and liquid-liquid equilibria and the reaction equilibrium and kinetics. It is shown that the separation side of the process can be well described with conventional methods and that the equilibrium stage concept is sufficient. However, contrarily to other examples from the literature, the results for the system studied here clearly show that reaction kinetics measured in the laboratory, in this case with a plug flow reactor, cannot always directly be used to successfully describe the reactive distillation. Here, this was only possible after introducing a transfer factor, which proved to be constant over a broad parameter range. To clarify why the transfer factor is needed, trickle bed reactor experiments with the same packing as used in the reactive distillation were carried out. They show that a substantial part of the correction is caused by the fluid dynamic nonidealities in the packing. Using the transfer factor determined in laboratory scale allows a safe scale-up. Enlarging the reaction zone may give additional security, yet this is only an option if side reactions do not pose a problem, as shown by side product formation studies in the present work. A Damkoehler number is derived for the overall process, which can be used to support and assess the sizing of the reaction zone. All in all, this work shows that mastering the reaction is the key to successful design and scale-up and provides methods and experience needed for this end. 1. Introduction Heterogeneously catalyzed reactive distillation is an attractive process for the production of many chemicals, e.g., for producing ethers and esters.1–6 Integrating reaction and distillation in one column can be beneficial for both reaction and separation, resulting in increased conversion and selectivity or overcoming azeotropes. On the other hand, undesired effects may also occur, such as an increasing byproduct formation. More information on reactive distillation in general is given elsewhere.7–10 In recent years, a considerable number of papers have been published that contain well-documented information on reactive distillation experiments in the laboratory scale, accompanied by simulation results, which are usually in good agreement with the experimental data.3–5,11–15 These results give confidence that reactive distillation, though being a complex, highly integrated process, can be reliably designed. However, important questions remain open. What are the key factors for achieving a reliable simulation? For example, rarely the consequences of the often-lacking match between the * To whom correspondence should be addressed. Tel.: +49711-685-66103. Fax: +49-711-685-66140. E-mail: hasse@ itt.uni-stuttgart.de. † Present address: BASF AG, Department of Process Engineering, 67056 Ludwigshafen, Germany. ‡ Present address: INEOS Paraform GmbH & Co. KG, 55120 Mainz, Germany.

conditions under which the reaction kinetics are measured and those in the reactive distillation experiments are addressed. Rarely, also, the sensitivity of the process simulation regarding uncertainties of the reaction kinetic model is checked. Furthermore, there is the question of the modeling depth needed to describe the vapor-liquid mass transfer: is the equilibrium stage model sufficient or is it necessary to use a rate-based approach? Another important issue is that of the scale-up. Does a successful simulation of a 50 mm column allow the design of a column in production scale? All these topics are addressed in the present work. The basis of the discussion is comprehensive experimental data on heterogeneously catalyzed reactive distillation for the test case of hexyl acetate synthesis that was measured in the frame of the European Union project INTINT.16 Based on this large body of experimental evidence and the results of its modeling and simulation, in the present paper the key factors for the design and scale-up of heterogeneously catalyzed reactive distillation are analyzed. This is done on the background of longstanding cooperation between the authors and industrial partners. It is shown that the key to successful design of heterogeneously catalyzed reactive distillation is mastering the reaction side of the process, which is strongly influenced by a complex interaction of physicochemical and fluid dynamic effects in the catalytic column internal.

10.1021/ie0714504 CCC: $40.75  2008 American Chemical Society Published on Web 07/19/2008

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Figure 1. Left: Reactive distillation lines (-•-) in the system water + acetic acid + 1-hexanol + n-hexyl acetate at 32.5 kPa in transformed coordinates. Azeotropes (4; given along with their respective boiling temperature); reactive immiscibility gap (shaded; - -, tie lines). Right: Sketch of left-hand-side reactive distillation line diagram (without immiscibility gap).

This work is characterized by the comprehensiveness of the studies of the test system hexyl acetate synthesis. The experimental results obtained for that test system have been published before in a series of papers to which appropriate reference is given in the present work, but neither have they been discussed as an entire body before nor have they been used for a consistent modeling and process simulation. In the present contribution, first, the chemical system is shortly introduced and an overview of its physicochemical properties as well as of the reactive distillation experiments is given. Based on that information, the modeling is discussed and simulation results are compared to the experimental data. This sets the stage for addressing the key issues of the design and scale-up of the reactive distillation process. 2. Reactive Distillation Experiments n-Hexyl acetate, a nature-identical flavoring agent with a sweet-fruity, pear-like odor and taste, is used in the food industry and in fragrances with a world production of about 1000 tons per year (according to private information obtained from producers). The synthesis of n-hexyl acetate has been chosen as a test system within the European Union project INTINT16 due to its complex thermodynamics and similarity with other systems of interest to the industrial partners. n-Hexyl acetate (HexAc) is formed by the reaction of 1-hexanol (HexOH) and acetic acid (AC) with water (W) as additional product; see reaction I. This reaction is a typical equilibrium limited esterification. In the present work, the strongly acidic ion-exchange resin Amberlyst CSP2 (Rohm and Haas) was used for catalysis. Under the conditions present in hexyl acetate reactive distillation, 1-hexene (HEN) and dihexyl ether (DHE) formation take place in side reactions; see reactions II–IV. For more details on the catalyst and the side reactions, see Schmitt and Hasse.17 H+

1-hexanol + acetic acid 798 n-hexyl acetate + water

(I)

H+

2(1-hexanol) 798 dihexyl ether + water

(II)

H+

1-hexanol 98 1-hexene + water

(III)

H+

n-hexyl acetate 798 1-hexene + acetic acid

(IV)

The studied system shows strong liquid and vapor phase nonidealities, with a large miscibility gap, several homogeneous or heterogeneous azeotropes, and a reactive azeotrope being present, and acetic acid dimerization taking place in the vapor phase. The proper description of phase equilibria is of prime importance for setting up reliable models for reactive distillation. In the present test system, not only vapor-liquid equilibria for the description of the processes in the reactive distillation column are crucial but also liquid-liquid equilibria must be known, as an external decanter is used to separate the two liquid phases occurring upon condensation of the top product. Schmitt and Hasse18 published a comprehensive and critically surveyed database of pure component vapor pressures and mixture vapor-liquid and liquid-liquid equilibria for the present test system and its subsystems. Special attention was paid to fully cover the temperature, pressure, and concentration ranges relevant for the reactive distillation process. Existing gaps in the literature were systematically closed by appropriate experiments. The NRTL model was used for the description of the liquid phase nonidealities. Details on the phase equilibrium model and the model parameters are given by Schmitt and Hasse.18 The reaction side of the hexyl acetate system was studied by Schmitt and Hasse17 in the same detail. Chemical equilibrium and autocatalyzed reaction kinetics of the main reaction were measured using batch reactors, whereas a plug flow reactor was chosen to study the heterogeneously catalyzed main and side reactions. The study covered the entire concentration, temperature, and liquid load ranges relevant for reactive distillation. Based on the experimental results, a thermodynamically consistent pseudohomogeneous model was developed to describe the main and side reactions. The primary data and details on the chemical reaction model and its parameters are given by Schmitt and Hasse.17 The conceptual design of a reactive distillation column for hexyl acetate synthesis can be understood based on the reactive distillation line map for the system water + acetic acid + 1-hexanol + n-hexyl acetate given in Figure 1, which combines the above-mentioned thermodynamic information with that on chemical reaction equilibrium. For details on reactive distillation lines the reader is referred to the publications of Frey and

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Stichlmair19 and Bessling et al.,20 and for details on the applied coordinate transformation the reader is referred to the work of Ung and Doherty.21 The reactive distillation line map given in Figure 1 has been generated with an in-house simulation tool for a pressure of 32.5 kPa, which was chosen so that the temperature in the reaction zone of the real column does not exceed the maximal allowable temperature of the catalyst of 120 °C. The topology of the studied system (cf. Figure 1) shows two distillation regions: a small one in the corner of pure 1-hexanol, which is not important here, and a big one containing the two products n-hexyl acetate and water, in which the present reactive distillation takes place. In that area, all reactive distillation lines start from the heaviest boiling n-hexyl acetate and end in the lightest boiling reactive heteroazeotrope. In order to clarify the topology described, Figure 1 contains a small abstracted sketch, which outlines the two distillation regions mentioned above and shows selected reactive distillation lines. Though theoretically possible, it is difficult to obtain the pure products n-hexyl acetate and water in a single reactive distillation column. The difficulty is not to withdraw water as a saddle-point productsto do this one would need a proper product split ratio and a sufficiently high column. However, the corresponding liquid phase profile would in this case simply run along the edges n-hexyl acetate-acetic acid and acetic acid-water where there is practically no 1-hexanol so that the reaction rates would be negligible. This makes the withdrawing of water as distillate nonattractive. Instead of pure water, the typical overhead product of a hexyl acetate reactive distillation column is a mixture close to the reactive heteroazeotrope. This is fully sufficient since the liquid-liquid phase split occurring after total condensation in an external decanter allows withdrawing an aqueous phase close to pure water. In this process operation the corresponding liquid phase profile crosses the composition square in the middle where the reaction driving force is the highest. Because of the high boiling point difference of 40 K between the reactants, the lower boiling acetic acid is introduced below and the heavier boiling 1-hexanol above the reaction zone in order to improve their contact. A nonreactive rectifying section is not required: the reactive distillation line map (cf. Figure 1) shows that only a small distillation effort is needed to reach an overhead product within the large miscibility gap that allows withdrawing water of acceptable purity. The insight obtained from the reactive distillation line map leads to the reactive distillation process setup shown in Figure 2, which is a two-feed reactive distillation column without rectifying section (the rectifying section shown in Figure 2 was usually not active as indicated by the dashed reflux arrow) and a decanter, of which the aqueous phase is withdrawn as product while the organic phase is used as reflux. n-Hexyl acetate is obtained as bottom product. This setup was implemented in 55 mm diameter laboratory scale and 162 mm diameter pilot scale. All information about those installations, such as dimensions, internals types, dry catalyst mass, sample locations, and temperature measurements, are given by Schmitt et al.22 for the laboratory-scale experiments and the first pilot plant series and by Schmitt et al.23 for the second series of pilot-scale experiments with modified internals. The experimental procedure is described as well in these two publications. In total 32 experiments in laboratory scale and 12 in pilot scale (four in the first series, eight in the second series) were carried out. A detailed documentation of the experiments (process parameters, mass flows, concentration and temperature profiles) is given by Schmitt.24 The high quality of the reactive distillation experiments in the laboratory scale is shown by mass

Figure 2. Setup of laboratory- and pilot-scale reactive distillation for the production of n-hexyl acetate. Dashed lines represent options.

balance errors always below 0.5% and component molar conversions deviating from the stoichiometric mean value by less than 3%. Three reproduction experiments, with up to 33 days of different operations in between, show deviations of less than 2.5% in terms of conversion, mass flows, and product purity and of less than 0.5 K for the temperature profile. The liquid loads in the reaction zone were between 3.5 and 6 m3/m2 h (typically 4.3 m3/m2 h) and F factors between 0.6 and 1.2 Pa0.5 (typically 0.8 Pa0.5). In the pilot-scale experiments the mass balance errors are always below 3.5% (first series)/2% (second series) and the component molar conversions deviate from the stoichiometric mean by less than 10% (first series)/5% (second series). The fluid dynamics in both pilot-scale series is characterized by liquid loads between 4 and 6 m3/m2 h (typically 4.9 m3/m2 h) and F factors ranging from 0.6 to 1.7 Pa0.5 (typically 0.9 Pa0.5). In the 32 laboratory-scale experiments, the influence of all relevant hardware (type of catalytic internals, feed location, operation with/without prereactor, nonactive/active rectifying section) and process parameters (heat duty, feed ratio, mass flow of aqueous phase reflux and of organic phase purge, fluid dynamic load, pressure) was studied. Since all these parameters were varied systematically, their influence on the process can be understood and insight in the process can be gained. The results of these one-parameter studies are described in detail by Schmitt et al.22 In addition, these experiments provide an excellent basis for model validation. The laboratory-scale experiments are characterized by reactant conversions in the range from 80 to 90%, thus exceeding the equilibrium conversion of 65.7% achievable in a plug flow reactor with stoichiometric feed. The mean selectivity of 1-hexanol to n-hexyl acetate was 97.4% and that of acetic acid was 99.5%, thus also showing the key role of 1-hexanol as the starting point for side reactions. In most experiments Sulzer Katapak-S was used as the catalytic internal. Alternatively, also Montz Multipak-II was studied. Nowadays, Sulzer Katapak-S has been commercially replaced by Sulzer Katapak-SP, which is identical in construction with Montz Multipak-II. The 12 experiments in pilot scale were carried out in two series, using different catalytic internals (first series, four experiments with Sulzer Katapak-S 250.Y; second series, eight experiments with Sulzer Katapak-S 500.Y). Their objective was to obtain scale-up information and to examine the behavior of

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the side reactions. Details are presented by Schmitt et al. Again, conversions are typically between 80 and 90%, but the side product formation increased dramatically with selectivities of 1-hexanol to 1-hexene from 2% to more than 10%. With these experiments an experimental basis of unique width and depth for the discussion of heterogeneously catalyzed reactive distillation is available. The same quality was reached for the database on thermodynamic and chemical properties introduced above, so that a sound discussion of the process modeling is possible by comparing simulation with experimental reactive distillation data. In the following section, the influence of reaction kinetics, fluid dynamics, and modeling of vapor-liquid mass transfer is addressed. Then the scale-up by about a factor of 10 between the laboratory scale and the pilot-plant scale is discussed. Finally, a dimensionless number is derived, which allows assessing the reactive section size. It is aimed to give answers to the questions formulated in the Introduction and in this way to further increase confidence when applying reactive distillation. 3. Simulation of Laboratory-Scale Reactive Distillation Experiments 3.1. Model Validation. The simulations of the reactive distillation experiments are carried out using an equilibrium stage model. Results from a rate-based model (two-film model with rigorous Stefan-Maxwell equations; details are given by Klo¨ker25) will be used for comparison. Model details are given by Schmitt.24 Reaction kinetics is taken into account using the results of Schmitt and Hasse17 for the heterogeneously catalyzed reaction. The autocatalyzed reaction is neglected due to its very slow kinetics and the small liquid phase holdup in the columns equipped with packings. Side product formation is neglected in the laboratory scale because the selectivities were close to 100%; however, for the simulation of the pilot-scale experiments the preliminary side reaction kinetics given by Schmitt and Hasse17 are taken into account. Phase equilibria are modeled using the two different sets of NRTL parameters published by Schmitt and Hasse.18 For the simulation of the reactive distillation column, where no liquid-liquid phase split occurs under the conditions studied, the vapor-liquid equilibrium parameter set is applied. The liquid-liquid phase split in the decanter is described using a NRTL parameter set especially fitted for this type of phase equilibrium. Note that the simulation results shown by Schmitt et al.22,23 are based on an old, preliminary database. Figure 3 shows a typical example for the result of the comparison between simulation and experimental data from the laboratory-scale series. Results from three different types of models are presented. The assumption that chemical reaction equilibrium is reached on each stage is obviously inadequate. A fully predictive simulation taking into account reaction kinetics, though being closer to the experimental data, still yields results that are insufficient for the design of a reactive distillation process. This is unexpected as there are many reports in the literature on good agreement between predictive simulation with similar models and reactive distillation process data.3–5,11–15 The discrepancy is especially astonishing as exceptionally high efforts were made in the present work to base the simulation on reliable physicochemical property data models. Similar efforts are rarely made in academic studies and distinctly exceed what would be acceptable in industrial process design. What are the reasons for the observed discrepancies? They might arise from an inappropriate modeling depth or from deficiencies in the description of either the separation or the

Figure 3. Reactive distillation experiment no. 4 in laboratory scale using Katapak-S. Liquid phase concentration and temperature profiles: comparison between experimental data (4; from Schmitt24) and simulation ( · · · , reaction equilibrium; - -, reaction kinetics predictive; s, reaction kinetics with ΦPFRfRD ) 0.292). Shaded: reactive section.

chemical part of the model. An analysis of these factors showed that the description of the separation side is not the problem. This is discussed below in more detail. It could be shown that deviations result from the direct transfer of the reaction kinetic model determined in plug flow reactor (PFR) experiments to the reactive distillation (RD) conditions. To allow accounting for differences between these cases, a transfer factor ΦPFRfRD was introduced in the pseudohomogeneous reaction kinetic model published by Schmitt and Hasse.17 In eq 1 rI denotes the PH rate of reaction I, kf,I (T) is the forward kinetic constant of reaction I at temperature T, a values are the activities of the different components, and Ka(T) is the reaction equilibrium constant based on activities.

(

rI ) ΦPFR f RDk f,IPH(T) aHexOHaAC -

)

1 a a Ka(T) HexAc W

(1)

The transfer factor ΦPFRfRD was fitted in the process simulation so that the overall acetic acid conversion calculated there meets that of the reactive distillation experiments. Figure 4 shows the results for the individual fits of all laboratory-scale experiments. Over the whole, broad range of conditions studied the transfer factor is almost constant (see Schmitt24 for a discussion of that finding). The mean value is 0.292, with a standard deviation of only 6% (corresponding to less than 2% in conversion). That number for the transfer factor was subsequently used for modeling the entire set of reactive distillation experiments. Very good agreement between the

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Figure 4. Transfer factors ΦPFRfRD of the laboratory-scale reactive distillation experiments (×, individual experiments; s, mean value of 0.292) and ΦPFRfTBR of trickle bed reactor experiments (O, liquid load study at 120 °C, line: guide for the eye) as a function of liquid load. All results for Katapak-S. Experimental data from Schmitt.24

experiments and the simulation was achieved. An example is given in Figure 3. For more examples, see Schmitt.24 The success of the simple transfer factor concept introduced here strongly supports the assumption that the initially observed deviations result from problems in the adequate modeling of the reaction kinetics in the reactive distillation column. 3.2. Interpretation of the Transfer Factor. A correction of reaction kinetics determined in laboratory reactor experiments by 70% is enormous and requires a more detailed investigation. First, fluid dynamic differences between the laboratory reactor for the determination of reaction kinetics and the reactive distillation column will be discussed. The laboratory reactor used in the present study consists of tubes filled with the catalyst and has plug flow characteristics. Although also the flow inside the catalyst-filled channels of a catalytic internal shows similar characteristics (Moritz and Hasse26), liquid bypassing the catalyst and imperfect catalyst usage (i.e., not wetted or stagnant zones) may significantly impact the performance of the packing as a reactor and could be a reason for the need for using a transfer factor in the reaction kinetics modeling. Parada27 has set up a trickle bed reactor to examine this influence of fluid dynamics on reaction conversion that was also used in the present study. Similar approaches to study the impact of fluid dynamics were chosen by Go¨tze,28 Kreul,29 and DeGarmo.30 As shown in Figure 5, the trickle bed reactor is a 50 mm tube equipped with 1.4 m Katapak-Ssthe very same catalytic internal as used in the laboratory-scale reactive distillationsarranged in five sections, with sample collection points in between. Reaction mixtures with compositions typical for the reactive distillation process are pumped onto the reactor head with liquid loads covering the entire range relevant for reactive distillation. The liquid is trickling down the reactor in the same manner as in a reactive distillation column, with the only difference that the countercurrent vapor phase is omitted. In this way, this setup allows studying the impact of liquid phase fluid dynamic effects on the reaction. More details on the reactor setup and operation can be found elsewhere.27 Including one reproduction, 14 experiments at three different feed concentrations, temperatures between 80 and 120 °C, and liquid loads between 1 and 7 m3/m2 h were carried out in the present study. Experimental details are given by Schmitt.24 Figure 6 shows as an example the concentration profiles of five trickle bed reactor experiments carried out at different liquid loads but otherwise identical parameters. Also shown in Figure 6 is the result of a plug flow reactor experiment (i.e., one of the experiments used for the determination of the reaction kinetics;

Figure 5. Trickle bed reactor setup (B, balance; D, drum; HE, heat exchanger; P, pump; X, liquid phase sample collection point).

Figure 6. Experimental trickle bed reactor n-hexyl acetate concentration profiles for various liquid loads (0, 1.0 m3/(m2/h); O, 2.0 m3/(m2/h); ], 3.0 m3/(m2/h); 4, 5.0 m3/(m2/h); 3, 7.0 m3/(m2/h)) as a function of residence time (expressed as mass of dry catalyst divided by mass flow). +, Corresponding plug flow reactor experiment. Experimental data from Schmitt.24

see Schmitt and Hasse17), which was carried out under the same conditions as the shown trickle bed experiments. Two conclusions can be drawn from that data. First, the trickle bed reactor performance is always below that of the plug flow reactor. This can be attributed to fluid dynamic effects such as bypass and imperfect catalyst usage. Second, with increasing liquid load the performance of the trickle bed reactor approaches that of the plug flow reactor, but it levels off without reaching the latter (the curves at 5 and 7 m3/m2 h already almost coincide, indicating the reaching of the load point as defined by Moritz and Hasse26). That means that even for the optimal liquid loads of the catalytic packing in terms of catalyst usage its performance as a reactor is never as good as predicted based on the laboratory plug flow reactor experiments (to which reaction kinetics was fitted). It should be mentioned that the laboratory

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plug flow reactor experiments of Schmitt and Hasse proved the absence of external mass transfer limitations under the conditions in the trickle bed reactor (see Schmitt24 for more details), thus excluding this effect as an explanation for the observed trickle bed reactor results. Using the transfer factor concept as described above in the same manner as in the case of the simulation of the reactive distillation experiments, the trickle bed reactor results can be further evaluated. The associated transfer factor is labeled ΦPFRfTBR and describes the transfer of the reaction kinetics from the plug flow reactor (PFR) to the trickle bed reactor (TBR). This transfer factor was fitted individually to the concentration profiles of each trickle bed reactor experiment. The resulting transfer factors ΦPFRfTBR of the five liquid-load-study experiments presented in Figure 6 are displayed in Figure 4, where they can be compared directly with the transfer factors ΦPFRfRD of the reactive distillation experiments. The transfer factor for the trickle bed reactor experiments increases with increasing liquid load. This corresponds to the well-known fact that, below the load point of the packing, increasing liquid load improves the catalyst utilization.26 In the liquid load range of the reactive distillation experiments the transfer factor is about 0.7. This means that the fluid dynamic effects, the influence of which was studied with the trickle bed reactor setup, are responsible for performance losses of about 30% compared to the laboratory reactor kinetics, and thus substantially contribute to the transfer factor of 0.292 determined for the reactive distillation experiments. The other trickle bed reactor experiments confirm this conclusion; see also Schmitt.24 Still, there must be further effects hampering the reaction in the reactive distillation column. Besides liquid phase fluid dynamics, the second big difference between the conditions in the laboratory reactor for the determination of reaction kinetics and those in the reactive distillation column is the presence of a liquid phase and a vapor phase, leading to simultaneous vapor-liquid mass transfer and chemical reaction in the latter apparatus. First, the question is addressed of whether the simultaneously occurring vapor-liquid mass transport influences the chemical reaction in a way that cannot be adequately described by a simple equilibrium stage model including reaction kinetics on the stage. For that purpose, besides the equilibrium stage model also a rate-based model based on the two-film theory was applied, which resolves the transport processes at the vapor-liquid interface in more detail. Both models were used with consistent physicochemical property models. The additional information needed in the rate-based model, such as mass transport correlations, was determined within the INTINT project.16 For more details about the rate-based model, see Klo¨ker.25 Using the ratebased model without any adjustment of the reaction kinetics leads to large discrepancies compared to the experimental data that are very similar to those observed with the equilibrium stage model. Moreover, if the transfer factor concept is applied in the rate-based simulation in the same manner as described above, almost identical results for the transfer factors are obtained; i.e., the averaged result for the transfer factor is now 0.296 as compared to 0.292 with the equilibrium stage model. As can be seen from the typical example of the simulation of a reactive distillation experiment in Figure 7, also the results of the ratebased simulation (simulation with height of discretization element ∆z ) HETP/4; leading to simulation results independent of ∆z choice) and the equilibrium stage simulation agree well. Consequently, it can be concluded that the way of modeling vapor-liquid mass transfer is not the cause for the need for

Figure 7. Reactive distillation experiment no. 4 in laboratory scale using Katapak-S. Liquid phase concentration profile: comparison between ratebased simulation (- -, Klo¨ker25) and simulation with the equilibrium stage model (s) as well as experimental data (4, from Schmitt24). Shaded area: reactive section. The transfer factors ΦPFRfRD are 0.292 for the equilibrium stage model and 0.296 for the rate-based model.

introducing a transfer factor. This directs the attention once more toward the reaction kinetics as the key to successful simulation of heterogeneously catalyzed reactive distillation. Even more, it can be concluded, in agreement with the results of Baur et al.,31 Klo¨ker et al.,32 Noeres et al.,33 Peng et al.,34 Sundmacher et al.,35 Bhatia et al.,36 and others, that using an equilibrium stage model (with reaction kinetics) is sufficient for the modeling ofsat leaststhe reactive distillation systems studied by these authors and that of the present work. Second and still related to the presence of the vapor phase in reactive distillation, the influence of the boiling conditions on the heterogeneously catalyzed kinetics has to be discussed. The presence of boiling conditions is a major difference between the state in the reactive distillation column and that in the laboratory plug flow reactor. In the latter case the pressure is set high enough to prevent the mixture from boiling; thus the experiments are conducted in a subcooled liquid phase. To understand the possible negative influence of boiling conditions on the heterogeneously catalyzed reaction, one has to consider that in the esterification studied here, besides the heavy boiling n-hexyl acetate (normal boiling point 171 °C), water is formed as an additional product. In comparison with n-hexyl acetate, water is an extremely light boiling component.

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Figure 8. Reactive distillation experiment no. 7 in pilot scale with Katapak-S 500.Y. Concentration profiles of liquid and vapor phase: comparison between experimental data (4, liquid; O, vapor; from Schmitt24) and simulation with ΦPFRfRD ) 0.292 (s, liquid; - -, vapor). Shaded area: reactive section. Table 1. Conversion of 1-Hexanol in Reactive Distillation Experiments in Pilot Scale with Katapak-S 250.Y (Experiments 1-4) and Katapak-S 500.Y (Experiments 5-12), Respectively; Comparison between Experimental Results and Simulation (ΦPFRfRD ) 0.292) experiment XHexOH,exp/% XHexOH,sim/%

1 94.1 92.9

2 84.1 80.4

3 89.6 89.1

4 72.7 72.7

5 85.5 81.5

The consequences might be 2-fold. First, this leads to water concentrations inside the catalyst placed in reactive distillation that are distinctly lower than those inside the catalyst placed in the subcooled liquid phase in the plug flow reactor, and which could affect reaction performance. Second, knowing that the acidic ion-exchange resin used as catalyst consists of a macroporous network, in which the reaction takes place (Blagov et al.37), the extreme boiling point difference between the formed water and n-hexyl acetate might lead to a partial evaporation inside the catalyst’s macropores or directly at the outer surface of the catalyst. This again would hamper the transport of the reactants acetic acid and 1-hexanol to the catalytic active sites inside the catalyst’s pores, thus reducing the reaction capacity. Supporting evidence of similar phenomena occurring in catalyst pores is reported by Datsevich et al.38 Experiments are on the way to further elucidate this. Considering that the various possible causes for the need to introduce the transfer factor depend on system-specific properties, it is obvious that the transfer factor determined here is only valid for the hexyl acetate system under investigation. However, the order of magnitude of the correction that is due to fluid dynamic nonidealities should be transferable to similar systems using the same catalytic internal and a similar catalyst. 4. Scale-Up of Reactive Distillation The insight gained from laboratory scale will be applied now to the simulation of the pilot-scale reactive distillation experiments. The procedure suggested is to use the reaction kinetic model corrected by the mean transfer factor of 0.292, which was determined in the laboratory-scale experiments, and to apply it directly to the predictive simulation of the pilot-scale experiments. The specific characteristics of the pilot-scale catalytic internalswhich differ from those in laboratory scale since it is not possible to keep both separation efficiency and reaction capacity perfectly constant during scale-upsare taken into account in the simulation via the number of theoretical stages per meter of the packing (NTSM) and the dry catalyst mass per stage.

6 82.7 81.8

7 81.8 81.7

8 80.8 75.2

9 79.2 77.3

10 81.2 79.0

11 86.9 86.0

12 80.1 78.1

Figure 8 shows a representative comparison of simulated and experimental concentration profiles (vapor and liquid phases) of a pilot-scale experiment. Very good agreement is observed. In Table 1 the experimental and simulated 1-hexanol conversion is compared for all pilot-scale experiments, yielding an average underestimation of only 2.3% (relative deviation, of which 1-2% can be attributed to formation of side products other than 1-hexene and dihexyl ether and thus not considered in the simulation). Even more, the simulation is able to reliably predict the effect of the changing the catalytic internal type between the pilot-scale experimental series 1 and 2, just by taking into account the different dry catalyst masses and NTSMs; see Schmitt.24 However, the reaction zone of the pilot-scale experiments was amply sized. This can be seen by having a closer look at the liquid phase concentration profiles of the experiment shown in Figure 8: in the lower third of the 6-m reaction zone the 1-hexanol concentration is virtually zero, thus stopping the main reaction (hydrolysis does not take place due to the absence of water). Such an amply sized reaction zone leads to an insensitivity of the model predictions to uncertainties regarding reaction kinetics. The positive effect of this is obviously to (at least apparently) limit the risks of a scale-up. However, as a consequence, based on the pilot-scale data from the present project, it cannot be clearly decided whether a transfer factor determined in the laboratory scale also holds for the pilot scale. Following the arguments presented above, it may be assumed that the contribution of the boiling conditions to the transfer factor remains roughly unchanged. Some differences between the laboratory-scale and the pilot- or industrial-scale packing might occur regarding fluid dynamics. Here it is likely that a production-scale reactive distillation column with proper liquid distribution exhibits less fluid dynamic nonidealities, such as bypass and imperfect catalyst usage, than a laboratory-scale column. Thus, using the transfer factor determined in laboratoryscale reactive distillation experiments either should lead to a good match between simulation and reality in production scale or give a conservative estimate. For the reasons explained above, the pilot-scale experiments of the present study cannot be

Ind. Eng. Chem. Res., Vol. 47, No. 16, 2008 6021

Figure 9. Sketch of the reactive distillation process (left) and the idealized reference process (right) for the reaction A + B T C + D.

regarded as a proof of this, despite the very good match of experiments and simulation. This leaves the proof to be an aim of future research. Oversizing the reaction zone to ensure a reliable design and scale-up should be carefully thought about: not only may backward reaction occur in regions where only little educts are present but also increasing side reactions must be feared. Once more looking at the reaction zone concentration profiles shown in Figure 8, it can be seen (besides the fact that the lower reactive section is not used for the main reaction) that n-hexyl acetate is present in concentrations of about 0.8 mol/mol with catalyst only waiting for work. This is especially dramatic as the 1-hexene formation starting from n-hexyl acetate is the fastest side reaction out of reactions II–IV (see Schmitt and Hasse17). This underlines the industrial need for research into scale-up of reactive distillation under reaction kinetic sensitive conditions in order to allow tailor-sizing the reaction zone. Side reactions cannot always be avoided, but they can be reduced by choosing proper operation parameters. In order to obtain the necessary insight into the process, side-reaction kinetics have to be known. Often, however, these kinetics are difficult to measure in typical laboratory reactors since the rate of side-product formation is usually much lower than that of the main reaction, resulting in only small side-product concentrations and very low concentration changes of the main components due to side-product formation. Schmitt and Hasse17 carried out kinetic experiments to study the side reactions using pure n-hexyl acetate or pure 1-hexanol as feed, with the objective to determine at least gross kinetic information. For the case of the experiment shown in Figure 8, a simulation that includes these side-reaction kinetics (using the same transfer factor of 0.292 as determined for the main reaction also for the side reactions) predicts a selectivity of 1-hexanol to 1-hexene of 4.9% as compared to 7.0% measured and a dihexyl ether mole fraction in the bottom product of 0.0014 g/g versus 0.0030 g/g experimental. Considering the very limited database that was available for the modeling of the side-reaction kinetics and the considerable uncertainty in the transfer from the laboratory reactor conditions to those in the reactive distillation experiment, this agreement is quite good. Using the side-reaction kinetics, comparative design studies can be made. The absolute numbers should not be overinterpreted. 5. Dimensioning of Reaction Zone The laboratory-scale experiments of the present study turned out to be very sensitive to reaction kinetics, whereas the pilot-

scale experiments are not. This is due to differences in the sizing of the reactive section. To characterize the latter and to allow a comparison with other reactive distillation processes, a dimensionless number is proposed here. It is defined by the ratio of the installed dry catalyst mass in the reactive distillation process divided by the minimum dry catalyst mass required to obtain the same conversion in a fictive, idealized process. For the illustration of the following derivation a reaction of the type A+B T C+D (V) is used, but it holds also for other types of reactions. Note that also inert components are allowed for. The reactive distillation process shown on the left-hand side of Figure 9 is considered as a black box. The feed is given by the total molar feed rate n˙F and its composition xF. The conversion of component A in the black box XA is achieved real using a certain amount of catalyst mcat,dry at a certain mean R reaction zone temperature T . The distribution of the components between distillate and bottom stream is irrelevant for the method proposed here; only the conversion is important. The fictive, idealized reference process used for the comparison is shown on the right-hand side of Figure 9. The reference process uses the same feed (n˙F, xF) and has to achieve the same conversion XA (again, the detailed composition of the different exit streams is irrelevant) at the same reaction temperature TR as the studied reactive distillation process, but now with min minimum catalyst mass mcat,dry . The reference process consists of a reactor R, a separation unit S1 that perfectly separates products from remaining educts, and a separation unit S2, which purges each of the unreacted educts exactly in the amount required to obey the desired conversion and which provides an educt recycle back to the reactor. The recycle flow rate is set to infinity for reasons explained below. Analysis of the resulting material balances for the reference process shows that the composition of the educt recycle is a degree of freedom and can be chosen. Following the objective of having an idealized process that requires a minimum amount of catalyst, the recycle composition is set to the stoichiometric composition (as we assume here simple second-order kinetics, the maximal reaction rate corresponds to stoichiometric reactants). Due to the infinitely large recycle flow, mixing it with the finite amount of feed still results in a stoichiometric reactor feed composition. Even more, since the finite catalyst mass in the reactor leads to finite production rates, the product concentrations and thus the backward reaction are negligible along the reactor length. In fact, mixture composition does not change along the reactor, and the reactor type does not affect the operation (any reactor

6022 Ind. Eng. Chem. Res., Vol. 47, No. 16, 2008

can be considered as a simple stirred tank reactor in this case). To conclude, the reactor of the fictive, idealized process operates under optimal conditions, i.e., with maximum reaction rate possible at a given reactor temperature, which results in a minimum catalyst requirement to achieve the given conversion. min The minimum catalyst mass mcat,dry of the reference process can be easily calculated from the component material balance of the reactor, which degenerates to the CSTR equation: +

min H nAR,out - ˙ ˙ nAR,in ) -mcat,dry ccat,dry × max R rf,V (T , stoichiometric composition)

(2)

Herein n˙AR is the molar flow of component A into and out of + the reactor R, cH cat,dry is the capacity of the catalyst, i.e., the mole max number of acidic sites per mass of dry catalyst, and rf,V is the forward reaction rate of reaction V. In the case of the hexyl acetate esterification, the reaction rate is taken from Schmitt and Hasse.17 Together with the definition of the molar conversion XA of component A: XA )

∆n ˙AR ˙ nAF

)

nAR,in - ˙ ˙ nAR,out xAF ˙ nA

(3)

the ratio of real catalyst mass (given reactive distillation process) to minimum catalyst mass (idealized reference process) results in real mcat,dry min mcat,dry

) +

real H max R mcat,dry ccat,dry rf,V (T , stoichiometric composition)

xAF ˙ nF XA

≡ Da (4)

Having a closer look at the right-hand side of eq 4, this catalyst mass ratio can be interpreted as the ratio of reaction rate to product removal rate (using eq 3 and the stoichiometry relation ∆n˙AR ) -∆n˙CR allows substitution of the denominator accordingly) and thus as a Damkoehler number to characterize the overall reactive distillation process. It will simply be labeled as Da here. The Damkoehler number Da can, e.g., be used for assessing the amount of catalyst installed in existing reactive distillation processes: the closer the Da number is to 1, the closer the installed catalyst mass is to the absolute minimum, and the more efficient is the process. In addition, the Da number can be used to assess the process sensitivity with regard to the catalyst amount or activity (which might decrease in operation due to deactivation) and reaction kinetics (which might be different in the column from what was expected in the design for the reasons discussed above). The following discussion of these issues is based on Figure 10. It shows the Da numbers calculated for reactive distillation experiments published by various authors for different reactive systems.3–5,14,15,27,39 The choice of the individual experiments shown in Figure 10 was made to encompass the whole range of experiments carried out by each author. Looking at selected hexyl acetate reactive distillation experiments from the present work, which are also included in Figure 10, it can be seen that the Da number is in the range 7-10 for the kinetic-sensitive laboratory-scale experiments; i.e., the process operates with a catalyst mass approximately 10fold the minimum mass a “perfect” process would need. In contrast, the Da numbers of the pilot-scale experiments are higher by a factor of about 3-4, and lie in the range 30-40.

Figure 10. Da numbers for selected reactive distillation experiments from the literature3–5,14,15,27,39 and from the present work.

These experiments are only weakly sensitive regarding reaction kinetics due the amply sized reaction zone, as discussed above. Interestingly, the correlation between the Da number and the process sensitivity regarding reaction kinetics (or installed catalyst mass) holds also for other reactive distillation systems described in the literature. In the following some examples are discussed which stem either from previous studies in which our institute was involved or from studies within the collaborative project INTINT16 in which also the present investigations were carried out, so that information on the background of the experiments was available. These are data from Moritz et al.39 and Klo¨ker et al.5 (both: ethyl acetate), Moritz et al.39 and Kolodziej et al.3 (both: TAME), and Parada27 (n-butyl acetate). The Da numbers for the ethyl acetate reactive distillation experiments of Moritz et al.39 and Klo¨ker et al.5 are about 30, and these experiments turned out to be only weakly sensitive to reaction kinetics (or catalyst mass). This is supported by Klo¨ker et al.,5 stating that doubling the installed catalyst mass hardly shows any effect. The TAME experiments described by Moritz et al.39 and Kolodziej et al.3 were carried out at very high Da numbers of about 150. As a consequence, it should be assumed that the process is completely insensitive toward reaction kinetics due to a very large reaction capacity of the reaction zone. This insensitivity is explicitly confirmed by Kolodziej et al.3 On the other hand, the n-butyl acetate reactive distillation experiments carried out by Parada,27 with Da numbers in the range from 6 to 11, show a high sensitivity regarding reaction kinetics. For example, Parada27 reports that introducing a transfer factor of 0.8, i.e., a variation in the reaction rate of -20%, leads to significant differences in the column profiles and conversion. The limited information available for the reactive distillation experiments of the other authors shown in Figure 10 seems to support the present interpretation (see Schmitt24 for more details). To summarize, Da numbers in the region of 10 and below seem to characterize processes with tailored reaction zones, in which the reaction zone is well used for the main reaction and unwanted side-product formation is low. On the other hand, such a design is sensitive to reaction kinetics. In contrast, high Da numbers, in the region of 30-40 and beyond, indicate that the reaction zone is dimensioned in a manner that it will easily fulfill its task regarding the main reaction, having the capacity to buffer insecurities, e.g., in reaction kinetics, and that in this respect also scale-up should not cause problems. However, such large reaction zones may be the cause of problems with side-

Ind. Eng. Chem. Res., Vol. 47, No. 16, 2008 6023

product formation as demonstrated with the hexyl acetate pilot plant reactive distillation experiments from the present work. 6. Conclusions In the European Union project INTINT16 partners from academia and industry worked together to establish the synthesis of n-hexyl acetate from acetic acid and 1-hexanol as a new test system for heterogeneously catalyzed reactive distillation. All parameters relevant for the design and scale-up of this integrated process were thoroughly studied for that system. Schmitt et al.22,23 gave an overview of more than 40 reactive distillation experiments carried out in laboratory scale and in pilot scale; detailed experimental data were published by Schmitt.24 The database of thermodynamic and chemical reaction properties was established by Schmitt and Hasse.17,18 Based on these comprehensive and detailed studies, the present work addresses the key factors for successful model-based design and scale-up of heterogeneously catalyzed reactive distillation. Predictive simulations of the laboratory-scale hexyl acetate reactive distillation experiments do not yield satisfying results. This is surprising, as state-of-the-art models (an equilibrium stage model including reaction kinetics and a rate-based model) are used, which are based on a comprehensive, thoroughly validated database. Only negligible differences between the ratebased and the equilibrium stage simulations are observed, which leads to the conclusion that the type of modeling the vapor-liquid mass transfer is not crucial. Rather, it is shown that the deviations between the model and the experiments result from problems in transferring the reaction kinetics determined in a laboratory reactor to the conditions present in reactive distillation. In other words, mastering the reaction is the key to the successful design of heterogeneously catalyzed reactive distillation. A simple transfer factor concept is introduced that allows carrying out that transfer in a pragmatic way. That factor describes the differences between the conditions in the laboratory reactor and the catalytic packing in the reactive distillation. These differences are 2-fold: first, there are fluid dynamic differences as was shown and quantified in trickle bed reactor experiments. Second, the experiments in the laboratory reactor are carried out in the subcooled liquid region whereas the reaction in the column takes place under boiling conditions, which seems to reduce reaction capacity strongly. The importance of that effect in the system studied in the present work may be related to the fact that the boiling point differences of the different substances are large. The insight obtained on the physicochemical background of the transfer factor still does not allow its reliable prediction. Thus, the procedure proposed here is to determine the transfer factor with laboratory-scale reactive distillation experiments and to apply it directly in scale-up. It is shown that the transfer factor determined in laboratory scale is virtually constant in a broad parameter range. The results of this work show that scale-up is safe when the system allows amply sizing the reaction zone. This, however, is an option only if side reactions do not pose a problem. If side reactions are critical, the reactive section must be tailored. Side-reaction kinetics, even if only grossly determined, are helpful for this. As the reaction zone in the pilot plant experiments of the present study was large, leading to results rather insensitive regarding reaction kinetics, the present data do not allow fully validating the application of the transfer factor in the case of scale-up with a tailor-sized reaction zone. It is likely that directly applying the transfer factor determined in the laboratory reactive distillation in such a case will lead to a

conservative design. However, additional research is proposed to validate these statements. A Damkoehler number was derived for the overall process, which supports the sizing of the reaction zone. To conclude, the analysis of the comprehensive data from the present project shows that mastering the reaction is the key to successful design and scale-up of heterogeneously catalyzed reactive distillation. This work provides methods and experience needed for this end. Acknowledgment The authors gratefully acknowledge financial support of the work by the European Union under the Competitive and Sustainable Growth Program (GROWTH), Project GRD1CT1999-10596 “Intelligent column internals for reactive separations (INTINT)”. Within this project, the fruitful cooperation with the following colleagues contributed to the present work: Markus Klo¨ker, Eugeny Kenig, and Andrzej Go´rak from University of Dortmund; Klaus Althaus and Hartmut Schoenmakers from BASF AG; and Peter Moritz, Oliver Bailer, and Claudia von Scala from Sulzer Chemtech Ltd. Literature Cited (1) Smith, L. A.; Huddleston, M. N. New MTBE design now commercial. Hydrocarbon Process. 1982, 61 (3), 121. (2) Stadig, W. P. Catalytic distillation: Combining chemical reaction with product separation. Chem. Process. 1987, 2, 27. (3) Kołodziej, A.; Jaroszyn´ski, M.; Sałacki, W.; Orlikowski, W.; Fra¸czek, K.; Klo¨ker, M.; Kenig, E. Y.; Go´rak, K. Catalytic distillation for TAME synthesis with structured catalytic packings. Chem. Eng. Res. Des. 2004, 82 (A2), 175. (4) Po¨pken, T.; Steinigeweg, S.; Gmehling, J. Synthesis and hydrolysis of methyl acetate by reactive distillation using structured catalytic packings: experiments and simulation. Ind. Eng. Chem. Res. 2001, 40, 1566. (5) Klo¨ker, M.; Kenig, E.; Go´rak, A.; Markusse, P.; Kwant, G.; Go¨tze, L.; Moritz, P. Investigation of different column configurations for the ethyl acetate synthesis via reactive distillation. Chem. Eng. Process. 2004, 43 (6), 791. (6) Hanika, J.; Kolena, J.; Smejkal, Q. Butyl acetate via reactive distillations modeling and experiment. Chem. Eng. Sci. 1999, 54, 5205. (7) Doherty, M. F.; Buzad, G. Reactive distillation by design. Trans. Inst. Chem. Eng. 1992, 70 (A), 448. (8) Stichlmair, J.; Frey, T. Reactive distillation processes. Chem. Eng. Technol. 1999, 22 (2), 95. (9) Barbosa, D.; Doherty, M. F. The simple distillation of homogeneous reactive mixtures. Chem. Eng. Sci. 1988, 43 (3), 541. (10) Taylor, R.; Krishna, R. Modeling reactive distillation. Chem. Eng. Sci. 2000, 55, 5183. (11) Hoffmann, A.; Noeres, C.; Go´rak, A. Scale-up of reactive distillation columns with catalytic packings. Chem. Eng. Process. 2004, 43, 383. (12) Steinigeweg, S.; Gmehling, J. Transesterification processes by combination of reactive distillation and pervaporation. Chem. Eng. Process. 2004, 43 (3), 447. (13) Singh, A.; Hiwale, R.; Mahajani, S. M.; Gudi, R. D.; Gangadwala, J.; Kienle, A. Production of butyl acetate by catalytic distillation: Theoretical and experimental studies. Ind. Eng. Chem. Res. 2005, 44 (9), 3042. (14) Steinigeweg, S.; Gmehling, J. n-Butyl acetate synthesis via reactive distillation: thermodynamic aspects, reaction kinetics, pilot-plant experiments and simulation studies. Ind. Eng. Chem. Res. 2002, 41, 5483. (15) Steinigeweg, S.; Gmehling, J. Esterification of fatty acid by reactive distillation. Ind. Eng. Chem. Res. 2003, 42, 3612. (16) European Union project, Intelligent column internals for reactive separations (INTINT), Competitive and sustainable growth (GROWTH) program, Project No. GRD1-CT1999-10596, 2000. (17) Schmitt, M.; Hasse, H. Chemical equilibrium and reaction kinetics of heterogeneously catalyzed hexyl acetate esterification. Ind. Eng. Chem. Res. 2006, 45, 4123. (18) Schmitt, M.; Hasse, H. Phase equilibria for n-hexyl acetate reactive distillation. J. Chem. Eng. Data 2005, 50, 1677. (19) Frey, T.; Stichlmair, J. Thermodynamic fundamentals of reactive distillation. Chem. Eng. Technol. 1999, 22 (1), 11.

6024 Ind. Eng. Chem. Res., Vol. 47, No. 16, 2008 (20) Bessling, B.; Schembecker, G.; Simmrock, K. H. Design of processes with reactive distillation line diagrams. Ind. Eng. Chem. Res. 1997, 36 (8), 3032. (21) Ung, S.; Doherty, M. F. Vapor-liquid phase equilibrium in systems with multiple chemical reactions. Chem. Eng. Sci. 1995, 50 (1), 23. (22) Schmitt, M.; Hasse, H.; Althaus, K.; Schoenmakers, H.; Go¨tze, L.; Moritz, P. Synthesis of n-hexyl acetate by reactive distillation. Chem. Eng. Process. 2004, 43, 397. (23) Schmitt, M.; Scala, C. von.; Moritz, P.; Hasse, H. n-Hexyl acetate pilot plant reactive distillation with modified internals. Chem. Eng. Process. 2005, 44, 677. (24) Schmitt, M. Heterogen katalysierte Reaktivdestillation: Stoffdaten, Experimente, Simulation und Scale-up am Beispiel der Synthese von Hexylacetat. Ph.D. Thesis, University of Stuttgart, 2006. (25) Klo¨ker, M. Modellierung und Simulation heterogen katalysierter Reaktivtrennverfahren. Ph.D. Thesis, University of Dortmund, 2007. (26) Moritz, P.; Hasse, H. Fluid dynamics in reactive distillation packing Katapak-S. Chem. Eng. Sci. 1999, 54, 1367. (27) Parada, S. Nebenreaktionen bei der Reaktivdestillation zur Synthese von Butylacetat. Ph.D. Thesis, University of Stuttgart, 2008. (28) Go¨tze, L. Makrokinetik der heterogen katalysierten Veresterung von Methanol und Essigsa¨ure sowie der Hydrolyse von Methylacetat in strukturierten Packungselementen. Ph.D. Thesis, University of Oldenburg, 1998. (29) Kreul, L.-U. Discontinuous and reactive distillation. Ph.D. Thesis, University of Dortmund, 1998. (30) DeGarmo, J. L.; Parulekar, V. N.; Pinjala, V. Consider reactive distillation. Chem. Eng. Prog. 1992, 88 (3), 43. (31) Baur, R.; Taylor, R.; Krishna, R. Development of a dynamic nonequilibrium model for reactive distillation tray columns. Chem. Eng. Sci. 2000, 55, 6139. (32) Klo¨ker, M.; Kenig, E.; Hoffmann, A.; Kreis, P.; Go´rak, A. Ratebased modeling and simulation of reactive separations in gas/vapour-liquid systems. Chem. Eng. Process. 2005, 44 (6), 617.

(33) Noeres, C.; Dahde, K.; Gesthuisen, R.; Engell, S.; Go´rak, A. Modelbased design, control and optimisation of catalytic distillation processes. Chem. Eng. Process. 2004, 43, 421. (34) Peng, J.; Edgar, T. F.; Eldridge, R. B. Dynamic rate-based and equilibrium models for a packed reactive distillation column. Chem. Eng. Sci. 2003, 58, 2671. (35) Sundmacher, K.; Uhde, G.; Hoffmann, U. Multiple reactions in catalytic distillation processes for the production of the fuel oxygenates MTBE and TAME: Analysis by rigorous model and experimental validation. Chem. Eng. Sci. 1999, 54, 2839. (36) Bhatia, S.; Ahmad, A. L.; Mohamed, A. R.; Chin, S. Y. Production of isopropyl palmitate in a catalytic distillation columnsexperimental studies. Chem. Eng. Sci. 2006, 61, 7436. (37) Blagov, S.; Parada, S.; Bailer, O.; Moritz, P.; Lam, D.; Weinand, R.; Hasse, H. Influence of ion-exchange resin catalysts on side reactions of the esterification of n-butanol with acetic acid. Chem. Eng. Sci. 2005, 61, 753. (38) Datsevich, L. Alternating motion of liquid in the catalyst pores at a liquid/liquid-gas reaction with the heat and gas production. Catal. Today 2003, 79-80, 341. (39) Moritz, P.; Go¨tze, L.; Schildhauer, T.; Kapteijn, F.; Majchrzak, M.; Salacki, W.; Markusse, M.; Kwant, G. ReactiVe pilot plant experiments for different systems in different equipment, known geometry; Deliverable 44, European Union project Intelligent Column Internals for Reactive Separations (INTINT), competitive and sustainable growth program (GROWTH), Project No. GRD1-CT1999-10596, 2002.

ReceiVed for reView October 26, 2007 ReVised manuscript receiVed April 19, 2008 Accepted April 29, 2008 IE0714504