Microreactors new technology for modern chemistry pdf

State-of-the-art in microreaction technology : concepts, manufacturing and applications. Abstract Microreactors as a novel concept in chemical technology enable the introduction of new reaction procedures in chemistry, molecular biology and pharmaceutical chemistry. Today, microreactors … Expand.

View 1 excerpt, references background. Microreactor technology and process miniaturization for catalytic reactions—A perspective on recent developments and emerging technologies. Recent developments in microreactor technology MRT are reviewed within the context of discovery, development and commercialization of catalytic systems. Emerging trends and drivers for development … Expand. We present a new micro reactor concept which is suitable for investigations of heterogeneously catalyzed liquid phase reactions.

In the set-up two reactants are mixed in a commercially available … Expand. View 1 excerpt, references methods. The capillary-microreactor: a new reactor concept for the intensification of heat and mass transfer in liquid—liquid reactions.

Abstract The capillary-microreactor was used for studying the nitration of a single ring aromatic in an exothermic liquid—liquid two-phase reaction. In the capillary-microreactor, isothermal … Expand. Review: Microstructured reactors for distributed and renewable production of fuels and electrical energy.

Abstract The current paper provides an overview of recent and past research activities in the field of microreactors for energy related topics. The main research efforts in this field are currently … Expand. Microreactors in organic synthesis and catalysis. The intensification of rapid reactions in multiphase systems using slug flow in capillaries. Lab on a chip. A multiphase microreactor based upon the use of slug flow through a narrow channel has been developed.

The internal circulation, which is stimulated within the slugs by their passage along the … Expand. A supported aqueous phase catalyst coating in micro flow Mizoroki—Heck reaction. Knoevenagel condensation reaction between benzaldehyde and ethyl acetoacetate in microreactor and membrane microreactor. The Knoevenagel condensation reaction between benzaldehyde and ethyl acetoacetate was performed in microreactor using Cs-exchanged NaX catalyst and NaA membrane.

Similar considerations apply to heterogeneous catalysis, and this will be commented on later. At first, a slow reaction is considered. For residence time reasons a stirred vessel or, better, a cascade of stirred vessels is needed. Each vessel is supplied with energy to evaporate the component to be separated. If the relative volatility of this component is very high, a one stage evaporation is sufficient. At a lower relative volatility the separation requires more stages, so a column has to be put on top of the vessel.

If the separation is even more difficult, a stripping section must be added to the column and a reboiler is necessary. In the limiting case of a very low relative volatility, each stage of the cascade can be operated as a countercurrent stage. The first stage is additionally provided with a fractionating column to enrich the component to be separated. Such a setup is equivalent to a reaction column with a large holdup on each reaction stage.

The equipment therefore may be selected under the aspect of separation efficiency. If the relative volatility of the component to be removed is low, a considerable number of stages is necessary.

The only appropriate equipment is a column. Depending on the required residence time it may be a packed or a tray column. A relative volatility in the medium range allows the number of stages to be reduced, though the total hold up has to be kept constant. A tray column, perhaps with special bubble cap trays is possible.

At a certain low relative volatility the 41 Fig. Bessling homogeneous catalysis rectification column column with larger volume in the bottom stirred vessel with rectifying column strirred vessel with full column vessel cascade with column strirred vessel with evaporator evaporator In the range of medium reaction velocities, mixed constructions are the ideal solution, for example special column trays with large holdup or reactor cascades with a column on the first stage only.

The resulting principles for the choice of equipment for homogeneous catalysis are presented in Fig. As mentioned earlier, heterogeneous catalysis can be treated in a similar manner: Additional, separate reaction volumes are necessary to retain the catalyst.

These volumes can be arranged either within the equipment or in a side position [10], coupled by recycle systems. As in the case of homogeneous catalysis, principles for 2 Reactive Distillation Process Development in the Chemical Process Industries heterogeneous catalysis Equipment suitable for combining reaction and distillation heterogeneous catalysis Fig. Some, not all of the possibilities are shown in Fig. After the choice of the equipment, which is in principle determined by the reaction velocity and the relative volatility the next step follows: the proper design of the chosen equipment under the special conditions of the process.

Only if this step is done with care will the advantage of the proper choice of equipment, the optimum in energy consumption, be realized. This is of course also true for RD. In the case of autocatalytic reactions the reaction velocity can only be influenced by the reaction temperature, in other words for RD by the pressure of the equipment.

Homogeneous catalysis allows the reaction velocity to be influenced by changing the catalyst concentration. Thus the reaction velocity can be adapted over a wide range to the needs of the distillation equipment. Heterogeneous catalysis requires a structure to fix the catalytic particles in the reaction zone. This may cause construction and operation problems and is also factor that limits the catalyst concentration that can be achieved.

The reaction velocity can be enhanced only to the limit set by the attainable concentration range. Furthermore, the possibility of enhancing the reaction velocity by higher temperature or pressure is limited, because in general the catalyst consists of ion-exchange particles, whose temperature range is limited.

So, homogeneous catalysis is much more flexible but has its price in an additional separation step necessary for catalyst recycling and by demands for expensive materials in the case of mineral acids.

Heterogeneous catalysis is simpler in principle, but technical problems have to be solved. In general the equipment will need more volume, for example the columns must have a bigger diameter. It should be clear from these considerations that decisions have to be made in each case. For RD, the methods of design and synthesis have been developed to a considerable extent, partly with the aid of the many analogies to distillation.

A major focus of research and development in future years should be the scale-up of reaction columns. This is where great deficiencies still lie Fig. Also, the methods of choosing the best equipment will have to be improved. In the expert community, columns are often seen as the only possible choice. Work will have to be done to ensure that the benefit of combining reaction and distillation can be enjoyed to the full by employing the most suitable equipment.

This report would not have been possible without the joint work on these projects, some of which are still running. Further, we thank Dr. Ulrich Block, former BASF colleague, with whom the question of the choice of suitable equipment was worked out. Bessling References 1 M. Doherty, F. Michael, G. Buzad, 8 D. Block, Chem. DeGarmo, V. Parulekar, V. Pinjala, Chem. Bessling, J. Schembecker, K. Sundmacher, Chem.

Barbosa, M. Doherty, Proc. Stichlmair, Chem. Stichlmair, H. Offers, R. Potthoff, Ind. Massachusetts, Schoenmakers, W. Buehler, Ger. Jones, , EP Smith, M. Huddleston, Hydro- carbon Processing, , 3, Krafczyk, J. Gmehling, Chem. Hoffmann, K. Gorak, L. Kreul, , DE 01 A1. Frey, F. Nierlich, T. Reusch, J.

Stichlmair, and A. Tuchlenski 3. These streams contain valuable components, which can be either processed in situ or extracted purely. A typical composition of a steamcracker C4 cut is shown in Fig. Maximizing its value is a major objective for most petrochemical companies. Therefore, a variety of techniques exists for upgrading the C4 streams by removal of pure C4 components and conversion of low-value streams to higher-value products.

Over the last decade reactive distillation RD has become a key technology for meeting increased productivity demands. Typical composition of steam-cracker based C4 cut [18], reprinted from Chem. Tuchlenski umn to reach high conversions. Pure isobutene demand is forecast to reach 0. Production of polybutenes will probably dominate the demand for isobutene, with an average annual growth rate of 3.

Pure isobutene can be obtained from a number of sources: raffinate I, pseudoraffinate the product stream from selectively hydrogenated mixed C4s , FCC Fluid Catalytic Cracking C4 raffinate, isobutane dehydrogenation, etc.

In general, conventional distillation fails because of the very close boiling points of C4 components. In the case of streams like raffinate I or hydroraffinate, isobutene can be extracted either by cold acid extraction with sulfuric acid or by conversion to an oxygenated intermediate, for example MTBE, and subsequent back-cracking of this stream.

The production cost of pure isobutene depends mainly on the C4 source. However, MTBE decomposition is becoming the preferred way of producing isobutene since it can be integrated easily into refinery processes. An isobutene-containing stream, for example raffinate I, is fed to a standard MTBE synthesis plant, where the reactive component isobutene is converted to the intermediate MTBE while the inert C4 components are separated as distillate.

Isobutene, as the low-boiling component, is recovered at the top, while the bottom product methanol, acting as a reactive entrainer, is recycled to MTBE synthesis. This promising way of producing isobutene will be discussed below. A further objective is to illustrate our understanding of process design for RD from the industrial point of view. Surprisingly, the combination of reaction and distillation might lead to the formation of reactive azeotropes. This phenomenon has been described theoretically [2] and experimentally [3] and adds new considerations to feasibility analysis in RD [4].

Such reactive azeotropes cause the same difficulties and limitations in reactive distillation as azeotropes do in conventional distillation. On the basis of thermodynamic methods it is well known that feasibility should be assessed at the limit of established physical and chemical equilibrium. This might lead to different column structures depending on the severity of the kinetic limitations [5].

For reactive distillation with catalytic packing, there are only a couple of alternatives. The major task of catalytic packing structures is to ensure an adequate contact between catalyst surface and liquid phase. An evident requirement in catalyst selection is a fairly long lifetime since the entire catalytic structure has to be removed during catalyst replacement. For some processes, external reactors may be a promising alternative.

Similarly to conventional distillation, the determination of 51 52 T. Tuchlenski the required column height is one of the most difficult tasks in process design.

In general, it is based on the packing separation efficiency, which is obtained from non-reactive experiments with well-known test mixtures and chemically inert packing. Care must be taken because data from lab scale packing may differ appreciably from the packing applied on the industrial scale. According to Moritz and Hasse [6], the Sulzer Katapak-S has about three theoretical stages per meter on the lab scale, while on the industrial scale, the same packing provides only one to oneand-a-half theoretical stages per meter.

Besides these essential questions various criteria such as location of the reactive zone and catalyst mass must be taken into account. In the case of reactive distillation, the column height is influenced not only by the separation efficiency but also by the required residence time.

Further difficulties in scale-up calculations arise from complex mass-transfer phenomena and hydrodynamic effects. Simulation might be a decisive basis for process design when all major data are present. However, with the current state of the art it is unlikely that any company would build a new reactive distillation without any pilot plant tests or reliable references.

There are two minimum azeotropes in the non-reactive mixture, the first one between the high-boiling MTBE and the intermediate-boiling methanol, and the second one between methanol and the low-boiling isobutene. A border distillation line between the two azeotropes divides the entire concentration space into two distillation regions. The curve of the chemical equilibrium extends between the two products isobutene and methanol. According to Frey and Stichlmair [7], reactive azeotropes can arise wherever the concentration change due to distillation is co-linear to the concentration change due to reaction.

In the present system, this consideration gives rise to two curves on which reactive azeotropes may exist. At a pressure of kPa, the equilibrium curve does not intersect either of the two locus curves of reactive azeotropes anywhere in the entire concentration space. Accordingly, no reactive azeotrope exists in the system under these conditions. A feasible column design can be devised from thermodynamic principles, as illustrated in Fig.

The reactive section is located in the center between the 3 Application of Reactive Distillation and Strategies in Process Design Isobutene Isobuten min. Isobutene 0, 0,1 1,0 distillation tray 0,9 0,2 reaction tray 0,8 0,3 0,7 0,4 0,6 chemical 0,5 equilibrium 0,5 0,6 border distillation line 0,4 0,7 0,3 0,8 0,2 0,9 0,1 1,0 0,0 0,0 0,1 MTBE 0,2 0,3 0,4 0,5 0,6 0,7 0,8 0,9 1,0 MeOH Fig.

Conceptual design of an RD column for the decomposition of MTBE and the corresponding concentration profile calculated from thermodynamic considerations [18], reprinted from Chem. Tuchlenski non-reactive stripping and rectifying sections. The corresponding ternary plot shows the concentration profile as a sequence of phase- or phase-and-chemical equilibrium stages. In the rectifying section, conventional distillation takes place leading to an azeotropic concentration of isobutene and methanol in the distillate.

The concentration change in the reactive section coincides with the chemical equilibrium lines. Therefore, the composition of the ternary mixture crosses the border distillation line from the left to the right in downward direction.

As discussed above, thermodynamic simplifications allow an easy description and, in turn, a good understanding of the fundamental mechanisms effective in RD. The mass transfer can be described by either assuming an equilibrium or applying a kinetic model. By analogy, the reaction can be expressed as an established chemical equilibrium or with reaction kinetics. In the majority of cases the so-called equilibrium stage model is used for simulation of distillation without chemical reaction.

This has also been shown to be useful in reactive distillation. The equilibrium stage model is a pragmatic approach suitable not only in the early stage of process development. In general, it is implemented in all current process simulators [9]. The mass-transfer stage model is recommended e. Unfortunately, there are only few data available on suitable masstransfer kinetics.

Furthermore, Higler et al. This makes the application of these models fairly difficult. For this reason, the equilibrium stage model was used in this study.

Ion-exchange resins were found to be suitable for this reaction. Besides catalyst activity and selectivity, a major requirement is to provide a catalyst lifetime of the same scale of time as the period of plant turnaround.

Otherwise, catalyst replacement would require a shutdown of the entire unit. However, the results of run-time testing experiments have shown that the ion-exchange catalyst used provides high stability for more than h. Only slight decreases in catalyst activity have been observed.

Beginning with pure MTBE, or different mixtures of MTBE and methanol or isobutene, all experiments were started in the kinetic regime and were continued until chemical equilibrium was reached. Fitting the experimental data to the well-known kinetic approach suggested by Rehfinger and Hoffmann [11], good agreement was found between model prediction and experimental data Fig.

This holds for both the description of the kinetic regime and the approach to chemical equilibrium. All relevant side reactions have been included in the model, for example, diisobutene formation [12, 13], dimethyl ether formation, and hydration of isobutene. Results of a kinetic experiment performed in batch mode and comparison with model predictions based on the kinetic approach suggested by Rehfinger and Hoffmann [11] [18], reprinted from Chem. Special attention was paid to the exact description of the azeotropes including the temperature dependence , as they were found to have a strong influence on the calculated conversion.

Steady-State Simulation After the feasibility analysis had shown that the process for the decomposition of MTBE can be carried out using reactive distillation equipment, this conclusion had to be verified using a rigorous model.

Since no information on mass-transfer limitations is available, the equilibrium stage approach is the model of choice. Therefore, such a simulation was set up in Aspen Plus Version The chemical reaction was implemented as a user subroutine.

The amount of catalyst per theoretical stage was adjusted to match the properties of Katapak-S by Sulzer. The feed position was set to the top of the reactive section, since the reactant MTBE is high-boiling. The feed was pure MTBE.

Its amount was varied in order to assess the influence on column performance. Two calculated concentration profiles are depicted in Fig. It can be seen that both profiles are fairly similar, although MTBE conversion is very differ3.

The feed position was located above the reactive section [18], reprinted from Chem. For the low conversion case, MTBE is enriched in the stripping section while for the high conversion case, the desired product methanol is enriched. In conclusion, MTBE decomposition may be performed using reactive distillation equipment. The top product is fairly pure isobutene and will contain some methanol, which has to be separated downstream using conventional technology. The laboratory runs were performed using a pressure column having a diameter of 80 mm and up to 25 bubble cap trays.

The column set-up is shown in Fig. The catalyst pockets of the packing were filled with the ion-exchange resin described above. The column was operated at a pressure of kPa. The feed was introduced above the reactive section. Further experiments were performed with the feed position below the reactive section. In a set of various experiments stable operation was observed. The bottom product was free of isobutene.

Besides the expected mixture of isobutene and methanol, the distillate contained also small amounts of MTBE. This is due to the insufficient separation performance in the rectifying section. However, this can be alleviated easily by placing more trays or conventional packing in this section. Configuration of the lab-scale reactive distillation column.

The distillation section is equipped with bubble cap trays, the reactive section with Sulzer Katapak-S [18], reprinted from Chem. Tuchlenski Fig. Comparison of experimental data and model prediction for the concentration profiles of the lab scale column [18], reprinted from Chem.

A good agreement between model prediction and experiment can be observed. The simulation was able to describe all experimental data sets with the same accuracy as shown in Fig.

From these results, it was concluded that the equilibrium stage model is capable of describing the labscale experiments based solely on information on phase equilibrium, reaction kinetics and hardware set-up, giving the simulation a predictive character. One of the most important questions is whether the performance of large scale equipment can be predicted with the same approach as applied in the lab-scale.

For future developments, this would allow us to dispense with costly and time-consuming experiments on the pilot scale. The experimental data on the pilot scale were obtained using a column mm in diameter. The non-reactive stripping and rectifying sections contained wire gauze packing, while the reactive section of the column was equipped with Katapak SP from Sulzer. The temperature in the rectifying section was controlled by varying the reboiler load. Simulating the pilot column with the same model that was successfully used for the 80 mm column, the computed conversion was systematically too high Fig.

Reaction kinetics reduced reaction rates due to incomplete catalyst wetting, mass-transfer limitations, or maldistribution and separation efficiency of the reactive packing were considered as possible reasons for these deviations.

It was found that the sensitivity of the computed conversion against in these factors was small compared with the experimentally observed trends.

The best agreement was attained by significant reduction of the reaction rate constant. This implies that pilot-scale experiments provide the only basis for scale-up. This conclusion is supported by the data shown in Fig. For the upper feed positions, the agreement is good, the trend in the conversion for the lower feed positions is qualitatively correct. However, the observed reduction in reaction rate still remains unexplained. The pilot-scale RD column was operated at similar space velocity and hydrodynamic 59 T.

Experimental conversion as a function of feed position symbols. The concentrations in the reactive sections were likewise similar in both cases. Nonetheless, the simulation showed that the model was able to describe observed concentration and temperature profiles of the laboratory column very well but failed in the simulation of the pilot column.

A similar observation was made by other authors in the investigation of the synthesis of methyl acetate. Whereas a 50 mm laboratory column could be described very well [16] with the equilibrium stage model, this held no longer in the case of a pilot column [17]. It is known that the geometry of the reactive packing types for lab- and pilot-scale applications differs. This is accounted for in the present model by using the appropriate values for NTSM Number of Theoretical Stages per Meter and catalyst mass per stage.

A more detailed description of the hydrodynamics has hitherto not been included and remains necessary for future work. Possible explanations for the lower conversion compared to the lab column may be incomplete catalyst wetting due to maldistribution effects, mass-transfer phenomena, or the formation of temperature gradients in the reactive packing.

As long as this puzzle remains unsolved, a reliable scale-up of RD processes is not possible without expensive pilot testing. The results led to a reasonable column configuration. A confirmation of the thermodynamic suggestions was derived by simple steady-state simulations of the suggested column set-up. In the next stage of process design a suitable catalyst and column internals were selected. The final task was to develop a scale-up procedure based on experimental data in the lab and pilot scale.

While the lab-scale experiments could be described satisfactorily with a simple equilibrium stage model, the same approach failed in the case of pilot-plant experiments. In order to fit the pilot-plant data to the equilibrium-stage model, a reduced catalyst activity was introduced. As a result we obtained good agreement between model prediction and experimental data in a broad range of operating conditions. However, with the present state of the art it has to be concluded that experiments at the pilot-plant scale are indispensable for establishing any heterogeneously catalyzed reactive distillation process.

Reliable scale-up solely on the basis of laboratory scale experiments does not appear to be possible yet. References 1 E. Stein, A. Kienle, K. Sundmacher, 11 A. Rehfinger, U. Hoffmann, Chem. Tejero, M. Iborra, F. Ung, M. Song, R. Huss, M. Malone, M. Doherty, Ind. Taylor, R. Moritz, H. Hasse, Chem. Frey, J. Aspen Technologies Inc. Higler, R.

Gravekarstens, U. Abrams, J. Redlich, J. Kwong, Chem. Steinigeweg, J. Gmehling, Ind. Beckmann, F. Reusch, C. Tuchlenski, Chem. Hasse 4. In typical reactive separation processes such as reactive absorption or distillation the superposition of reaction and separation is deliberately used. In other cases, simultaneous reaction and separation simply cannot be avoided. This is, for instance, the case when side reactions occur in separation equipment or when intrinsically chemically reactive mixtures, such as solutions of weak electrolytes or formaldehyde solutions, have to be separated.

Furthermore, in many reactors products are directly removed, which is basically a reactive separation. Thermodynamics plays a key role in understanding and designing all these processes. The fact that reaction and separation occur simultaneously gives rise to special challenges both in experimental investigations and modeling the processes. It is these challenges that we will focus on in the present chapter. This will be done using case studies, which illustrate general facts, rather than by trying to cover the subject comprehensively.

There are several contributions of thermodynamics to the field of reactive separations. Thermodynamics provides the basic relations, such as energy balances of equilibrium conditions, used in the process models, which again are the key to reactive separation design. Furthermore, thermodynamics provides models and experimental methods for the investigations of the properties of the reacting fluid that have to be known for any successful process design. We will focus on equilibrium thermodynamics here but discuss relations to kinetic models.

The present chapter is aimed at readers interested in reactive distillation RD. The basic ideas can, however, easily be applied to other reactive separation processes. Two basic types of model are used: stage models, which are based on the idea of the equilibrium stage with phase equilibrium between the outlet streams, and rate-based models, which explicitly take into account heat and mass transfer. Similarly to the physical side of RD, the chemical reaction is either modeled using the assumption of chemical equilibrium or reaction kinetics are taken into account.

Note that a kinetic model, either for physical transport processes or for chemical reactions, always includes an equilibrium model. The equilibrium model is the stationary solution of the kinetic model, for which all derivatives with respect to time become zero. Hence, whatever model type is used, it has to be based on a sound knowledge of the chemical and phase equilibrium, which is supplied by thermodynamic methods. Starting from there, kinetic effects can be included. Only a limited amount of data is needed to develop models for RD processes if the assumption of both phase and chemical equilibrium is used.

Nevertheless, even the development of a reliable equilibrium model alone is a challenging task, which is often underestimated. The amount of information needed to develop reliable kinetic models greatly exceeds that for the equilibrium models. Finding reliable model parameters is often the bottleneck in model development. An excellent comprehensive survey on the fundamentals of different types of model for RD processes has recently been given by Taylor and Krishna [1].

In that paper the focus is on modeling concepts, we focus here on the application of such models, especially the comparison of model predictions with experimental data and the background of the complexity of the model and the effort needed for its parameterization. We do account for the fact that the results of any such comparison will depend on the chosen example, but emphasize that comparison with experimental data, especially in predictions, is the final test for any modeling strategy.

Unfortunately, there is only a very limited amount of data on RD experiments available in the open literature for such comparisons. Full details are reported by Moritz [4]. Methyl acetate production is often used as an example to demonstrate the advantages of RD processes [5]. The basic set-up of the heterogeneously catalyzed methyl 4 Thermodynamics of Reactive Separations Fig. The catalytic packing used in the present study was Katapak-S with an Amberlyst ion-exchange catalyst. Experiments were carried out both at the laboratory scale [6] and the semi-industrial scale [4].

Two major classes of modeling concepts, the equilibrium stage and the ratebased approach, were tested. The chemical reaction is assumed to occur only in the liquid phase. A survey of the four models used is given in Fig. Three variants of the equilibrium stage model were considered. In the simplest model Model 1 the RD column is modeled as a countercurrent multistage process with physical and chemical equilibrium on each stage. In Model 2 the assumption of chemical equilibrium is dropped and the reaction on each stage is described by second-order bulk reaction kinetics.

Model 3 was developed to take into account the residence time behavior of the liquid flow in the packing. For this goal the reaction zone is described by a relatively large number of stages, or, in terms of reaction modeling, stirred tank reactors. Neither phase nor chemical equilibrium is reached Fig. Hasse at these stages. Reaction kinetics are the same as in Model 2, Murphree efficiencies are used to give the desired separation capacity.

In the rate-based approach Model 4 heat and mass transfer are directly taken into account using two-film theory. Bulk reaction kinetics are included as in Models 2 and 3. Obviously, parameterizing these models meaningfully requires a very different amount of experimental data, increasing from Model 1 to Model 4.

The phase equilibrium model and the chemical equilibrium model should be consistent, as discussed in more detail below. Extending the equilibrium model to Model 2, which accounts for reaction kinetic effects, is not straightforward.

In the case of methyl acetate, the reaction kinetic model obtained from stirred-tank experiments is not in agreement with the results obtained from tricklebed experiments, so that empirical factors have to be introduced [4, 7].

Furthermore, the macrokinetic model requires information on the liquid flow velocity, so that the thermodynamic design and the fluid dynamic design, which can be done sequentially when using Model 1, are now directly coupled.

The reaction kinetic model should be consistent with the chemical equilibrium model. Going from Model 2 to Model 3, the only additional information needed is the residence time distribution of the packing, which is available for Katapak-S [3].

The residence time distribution of a packing segment of length L is modeled using a cascade of Nrea stirred tank reactors.

The same packing segment has a known number of theoretical stages Nsep. As Nrea is always larger than Nsep, the differences in separation and reaction capacity of one stage stirred tank can easily be accounted for by introducing a Murphree efficiency. This can, however, be done with an accuracy sufficient for most cases using standard methods [8]. The amount of additional information needed to be able directly to take into account heat and mass transfer in Model 4 is high.

Using the two-film theory, information on the film thickness is needed, which is usually condensed into correlations for the Sherwood number. That information was not available for KatapakS so that correlations for similar non-reactive packing had to be adopted for that purpose.

Furthermore, information on diffusion coefficients is usually a bottleneck. Experimental data is lacking in most cases. Whereas diffusion coefficients can generally be estimated for gas phases with acceptable accuracy, this does unfortunately not hold for liquid multicomponent systems. For a discussion, see Reid et al. These drawbacks, which are commonly encountered in applications of rate-based models to reactive separations, limit our ability to judge their value as deviations between model predictions and experimen- 4 Thermodynamics of Reactive Separations 20 21 22 23 21 22 23 Model III 21 19 20 Model II 20 18 17 17 17 19 16 16 16 Model I 15 15 15 19 14 14 14 18 13 13 13 experiment 12 12 12 18 11 11 11 10 10 10 1 0.

All model results are predictions in the sense that no adjustments to RD data were made. The predictions from the stage models, which take into account reaction kinetics Models 2 and 3 , are good and do not differ largely. Hasse Differences between Model 2 and Model 3 would only have been expected for residence time sensitive reaction systems, such as, for instance, ethylene glycol production from ethylene oxide and water, but not for the esterification studied here.

The very simple Model 1 still gives reasonable predictions, which would, for instance, be sufficient for conceptual design studies. The fact that the rate-based Model 4 gives entirely wrong predictions in the present case should not be overemphasized and may be due to the very limited amount of available thermophysical and fluid dynamic input data.

Obviously, the choice of an appropriate modeling level for RD processes strongly depends on the input available to parameterize the different models. Thermodynamics plays a key role in providing that input. As stated in the previous section, stage models with the simple assumption of chemical equilibrium in the streams leaving the stage and phase equilibrium between these streams often already provide a process model with reasonable predictive power.

Any more detailed model will include an equilibrium model as a limiting case and will therefore have to be based on reliable information about the equilibrium. We will give a brief outline of the basic concepts of thermodynamic modeling of simultaneous chemical and phase equilibrium here.

The focus is on the options provided by classical thermodynamics. A discussion of different GE models or equations of state is not within the scope of the present chapter. Share This Paper. Background Citations. Methods Citations. Citation Type. Has PDF. Publication Type.

More Filters. Synthesis and post-processing of nanomaterials using microreaction technology. A critical barrier to the routine use of nanomaterials is the tedious, expensive means of their synthesis. Microreaction technology takes advantage of the large surface area-to-volume ratios within … Expand. View 1 excerpt, cites background. Enzymatic Processing in Microfluidic Reactors. Ionic-liquid supported oxidation reactions in a silicon-based microreactor.

Abstract The combination of microfabrication and reaction engineering techniques has the potential to produce powerful microreactors.