Ind. Eng. Chem. Res. 2006, 45, 2707-2714
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Comparison of Reactive Distillation with Process Alternatives for the Isobutene Dimerization Reaction Ravindra S. Kamath,†,§ Zhiwen Qi,† Kai Sundmacher,*,†,‡ Preeti Aghalayam,§ and Sanjay M. Mahajani§ Max-Planck-Institute for Dynamics of Complex Technical Systems, Sandtorstrasse 1, D-39106 Magdeburg, Germany, Process Systems Engineering, Otto-Von-Guericke-UniVersity Magdeburg, UniVersita¨tsplatz 2, D-39106 Magdeburg, Germany, and Department of Chemical Engineering, Indian Institute of Technology Bombay, Powai, 400076 Mumbai, India
Reactive distillation is an attractive alternative to conventional multiunit structures. In the case of irreversible reactions involving several products, reactive distillation may effectively improve the selectivity toward a desired intermediate product. However, multiunit processes for such reactions cannot be simply discounted without a systematic comparison with reactive distillation. This work emphasizes such a comparison using the isobutene dimerization reaction as an example. The multiunit configurations considered include the conventional reaction-separation network with and without recycle and a distillation column coupled with a side reactor. Potential configurations of process alternatives are generated for a wide range of the conversion of isobutene and the alternatives are compared on the basis of the selectivity toward diisobutene achieved and cost incurred for identical isobutene conversion. Results show that, for a given isobutene conversion, the reactive distillation process is capable of providing the maximum selectivity among all the investigated alternatives. However, the conventional processes are competitive when costing is also involved in the evaluation. Introduction Alkylates are considered to be potential alternatives to MTBE, whose status as a gasoline additive is questionable considering its association with groundwater contamination. One of the sources of alkylates is the dimerization of isobutene (IB) to diisobutene (DIB; i.e., isooctene) followed by hydrogenation. There were claims that this indirect alkylation process is capable of utilizing the existing facilities with reasonable revamping and the feedstock used for MTBE production.1-3 The dimerization reaction is a typical series-parallel reaction network with the desired product being an intermediate in the network that ends with the undesired oligomers. Thus, both the conversion of the reactant (IB) and the selectivity toward the desired product (DIB) are desirable to be maximized simultaneously. However, in many of such reactions, the conversion and selectivity show an inverse relationship, i.e., as conversion increases selectivity decreases and vice versa. Since the reaction is fast and highly exothermic,4 temperature control by external cooling and addition of polar components have been traditionally used to improve the DIB selectivity. The reaction characteristics make it an interesting candidate for reactive distillation (RD). RD can maintain a low concentration of the dimer in the reactive zone by simultaneous distillation due to the difference of component boiling points, thereby suppressing the further oligomer-producing reactions. In our previous work,5 the feasibility of the dimerization reaction in reactive distillation has been illustrated. Results showed that a RD column structure similar to that commonly used for MTBE production can be successfully applied for the dimerization reaction and that high selectivity to DIB can be obtained with adequate temperature * Corresponding author. Tel.: +49-391-6110351. Fax: +49-3916110353. E-mail:
[email protected]. † Max-Planck-Institute for Dynamics of Complex Technical Systems. ‡ Otto-von-Guericke-University Magdeburg. § Indian Institute of Technology Bombay.
control, even in the absence of a polar component. These results strengthen the selection of the RD process for the dimerization reaction. Though RD seems suitable for the dimerization of isobutene from the technical point of view, practical issues, e.g., complicated designs of the column internals and inconvenience in catalyst handling during installation, removal, and regeneration, suggest that external reactors might be favorable because of their flexibility. However, external reactors do not possess the separation effect in the reactive zone, which may hamper the product selectivity. To overcome the problems of RD and to somewhat maintain the benefits of the combined process, application of the side-reactor concept6,7 or reactive-pump concept8 are suggested. This concept involves taking out a stream from the distillation column at a strategic location and introducing it into an external reactor or a series of reactors. In a sense, the reactor is coupled to the column but retains all the characteristics of an external reactor with reaction conditions very less limited by the distillation conditions. So far, systematic comparison of RD with other competitive processes is missing in the literature; the only notable exception worth mentioning is the recent work by Kaymak and Luyben9 which demonstrates this comparison for a hypothetical reversible reaction with ideal vapor-liquid equilibrium (VLE). For reactions such as dimerization, the final optimum combination of values for the IB conversion and DIB selectivity cannot be specified without doing a very rigorous optimization analysis involving process cost. At a preliminary stage to select a process alternative, one can expect that the desired operating point could lie anywhere within a wide window of the IB conversion or DIB selectivity. Hence, the comparison needs to be done bearing this fact in mind. The matter is complicated by the fact that RD can involve nonlinear behavior, such as multiple steady states, which can significantly influence or perhaps even tilt the balance in the favor of multiunit configurations if RD is not operated in the optimal window. Hence, this
10.1021/ie051103z CCC: $33.50 © 2006 American Chemical Society Published on Web 03/11/2006
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Table 1. Configuration of the Base Case for the Reactive Distillation Process total stages reactive stages pressure (MPa) reflux ratio reboiler duty (MW) Damko¨hler number
30 4-26 1.5 6.0 1.848 2.61
total catalyst loading (kg) C4 feed flowrate C4 feed composition (IB/nB, mol) ratio of polar to main feed (mol/mol) stage location of C4 feed stage location of polar feed
work begins with the process analysis of reactive distillation in order to study the influence of important design and operating parameters. For the other multiunit configurations being considered, unlike RD, the operating conditions for the reaction and distillation steps are independently analyzed and selected. Potential configurations are generated for all the selected process alternatives over a wide range of conversions. The configurations are then compared on the basis of the DIB selectivity achieved and cost incurred at identical conversions. Process Analysis of Reactive Distillation Modeling Fundamentals. The reaction is conducted in a hybrid reactive distillation column where the reactive zone lies between the rectifying and stripping sections. The widely used equilibrium stage model is applied. The activity and fugacity coefficients are calculated using the modified UNIFAC method10 and the Peng-Robinson-Stryjek-Vera equation of state,11 respectively. The kinetic model for the dimerization reaction is taken from Honkela and Krause,12 who considered the following two reactions simultaneously in their experiments:
Reaction 1: 2IB f DIB Reaction 2: DIB + IB f TIB The corresponding rate expressions are
r1 )
r2 )
kDIBaIB2 (aIB + BTBAaTBA + Bsolventasolvent)2 kTIBaIBaDIB (aIB + BTBAaTBA + Bsolventasolvent)3
(1)
(2)
The column model is formulated as a system of differentialalgebraic equations and solved using the DIVA simulation environment.13 Base-Case Configuration. A base-case configuration (Table 1) for the RD process is generated by trial and error, which is close to a set of optimum parameters and is considered as a potential RD configuration. Under such conditions, more than 90% IB conversion as well as DIB selectivity is achieved because of the synergetic effects of reaction and distillation. The corresponding composition profiles in the column are shown in Figure 1. Because of the high exothermicity of the dimerization reaction, for standard reactor types, it is extremely difficult to maintain the temperature within the permissible limits unless a sufficient amount of external cooling is provided. This fact is demonstrated through a sample temperature profile in an “almost-adiabatic” plug-flow reactor for a typical simulation with 40 mol % IB in the C4 feed. As can be seen in Figure 2a, the temperature increases continuously along the length/volume of the reactor. An amount of external cooling was intentionally provided so that temperature does not increase beyond permissible limits of ∼383 K. On the contrary, as shown in Figure 2b, RD provides an almost-constant temperature profile in the reactive zone even without any sort of external cooling. Thus,
700.64 100.0 1.5 2:100 26 6
RD eliminates the extraneous cooling requirement by effectively using the heat of reaction for vaporization of the liquid phase on the reactive trays. Although the base-case configuration and the resulting column profile shown here are obtained in the presence of polar components, they are similar to the case in the absence of polar components.5 This is not surprising, because the amount of polar component used throughout the analysis done in this work is very small. However, even small amounts of polar components can severely reduce the reaction rate, and hence, a higher catalyst loading is required in order to obtain the same conversion as in the nonpolar case. It is to be noted that the presence of polar components always leads to a better selectivity toward DIB than that for the nonpolar case, which coincides with the observation of the kinetics study.14 Hence, throughout this work, polar components are being used with the view that the increase in selectivity offered by the addition of polar components outweighs the additional costs incurred due to a higher catalyst loading and separation of the polar components from the final product. It should be noted that polar components have also been applied in the industrial plant NExOCTANE technology. For process development, the influence of various parameters on the RD configuration needs to be studied. Though similar work has been done with respect to a nonpolar base-case configuration,5 the influence of the most important parameters is restudied in this work using the polar base case as the reference. Effect of Number of Reactive Stages. For this analysis, the numbers of rectifying and stripping stages are fixed at 3 and 4, respectively, according to the separation requirements, and only the number of reactive stages is varied. The C4 feed is introduced at the bottom-most (last) reactive stage, while the polar feed is introduced at the topmost (first) reactive stage. Figure 3 shows the effect of number of reactive stages in the RD column on the IB conversion and the DIB selectivity. Both conversion and selectivity continue to increase as the number of reactive stages is increased. However, the improvement in selectivity is insignificant after 5 reactive stages, and this is the same case
Figure 1. Composition profile for the base case (shaded region indicates the reactive zone).
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Figure 4. Effect of the location of C4 feed on reactive distillation column performance.
Figure 2. Demonstration of the automatic temperature control in reactive distillation. (a) Temperature profile in a plug-flow reactor. (b) Temperature profile in reactive distillation column (shaded region indicates the reactive zone). Figure 5. Effect of the location of C4 feed on DIB selectivity (IB conversion fixed at 90% by suitably changing the catalyst loading).
Figure 3. Effect of the number of reactive stages on reactive distillation column performance.
with conversion beyond 20 reactive stages. To obtain the best possible combination of the DIB selectivity and IB conversion, 23 reactive stages were finally selected for the base case. Effect of Location of C4 Feed. Heuristics on RD generally recommend that, if the reactants are the most volatile compo-
nents in the system, as is the case here, the feed should be introduced near the bottom of the reactive zone. So for this analysis, while the location of the polar feed is fixed (as the base case), the C4 feed location is moved from stage number 16 (near the center of the reactive zone) down to stage number 28, i.e., two stages below the reactive zone of the base case. As seen in Figure 4, the IB conversion shows an optimum at stage number 23 while the DIB selectivity continuously increases. An optimum feed location cannot be selected based on this result. However, it could be reasoned from Figure 4 that the optimum feed location should lie anywhere between and including stages 23-28. As both conversion and selectivity change in the opposite direction from the 23rd stage onward, the following strategy is employed. While the C4 feed location is varied, the catalyst loading is changed in order to obtain the same IB conversion of 90%. As a result, the variation of DIB selectivity with C4 feed location is now shown in Figure 5. The maximum selectivity is obtained at the optimum C4 feed location, i.e., stage 26. Effect of Column Pressure. The dimerization reaction is generally conducted in the temperature range of 313-393 K, while C4 boils in that temperature range under the approximate pressure range of 6-18 atm. Hence, the effect of pressure on
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Figure 6. Effect of the operating pressure on reactive distillation column performance.
Figure 8. Effect of the catalyst loading on reactive distillation column performance.
components. For this analysis, the catalyst loading is represented in the form of Damko¨hler number (Da), defined as
Da )
Figure 7. Effect of the reflux ratio on reactive distillation column performance.
the performance of the RD column has also been studied in this pressure range, as shown in Figure 6. As pressure is increased, the IB conversion increases because of higher reaction rates. It worth noting that no sharp change in the DIB selectivity is observed. Instead, it remains close to 95% throughout the pressure range. For standard batch or continuous reactors, experimental investigations showed lower DIB yields as compared to simulation results of RD, but quantitative comparisons are not possible because of the absence of adequate information.15-17 Moreover, the bifurcation diagram shows a steadystate multiplicity in a narrow pressure range near 8.8 atm. A pressure of 15 atm was selected for the base case as it provides a better combination of DIB selectivity and IB conversion, both of which exceed 90%. Effect of Reflux Ratio. The effect of reflux ratio is studied for a fixed value of the reboil ratio of 10.0, as used in the basecase configuration. The column performance is shown in Figure 7, which shows lower IB conversion at lower reflux ratios. Hence, it is advisable to operate the reactive distillation column at a relatively high reflux ratio to ensure satisfactory performance. Effect of Catalyst Loading. Catalyst loading is the only parameter which is significantly affected by the presence of polar
WTkhet f,ref FT
(3)
where FT is the total feed flow rate to the column, WT is the het is the forward total amount of catalyst in the column, and kf,ref rate constant evaluated at a reference temperature which is conventionally the lowest-boiling pure component in the system (i.e., boiling point of isobutene at the column pressure). The catalyst loading is equally divided on the reactive stages. Da is varied from zero (i.e., nonreactive distillation) to 16 (∼99% conversion of IB), as shown in Figure 8. With an increase in catalyst loading, the IB conversion increases sharply in the initial period, but beyond Da ) 4.0, a large increase in catalyst loading is required for a small amount of further improvement in IB conversion. The selectivity to DIB decreases from near 100% at very small Da values to about 93% at Da ) 16. The final choice of Da ) 2.61 in the base case depends on a compromise between conversion and selectivity. Selection of Process Conditions for Multiunit Configurations The process alternatives other than RD that are considered in this work do not involve simultaneous reaction and separation. Hence, for these process alternatives, operating conditions for reaction and separation can be determined independent of each other. For a fair comparison, an effort must be made to optimize the conditions for the reaction and separation units for better performance. Reaction Temperature. As stated earlier, the dimerization reaction is highly exothermic such that external cooling is almost inevitable when RD is not utilized. If external cooling can be manipulated favorably, the question arises as to whether there is a need to maintain any particular temperature profile in the reaction unit. This question can be best answered through the use of optimization techniques. However, for a preliminary analysis, a constant temperature profile (i.e., isothermal conditions) within the reaction unit is assumed. Now, a constant operating temperature for the reactor needs to be selected that can give the best-possible performance. It can be expected, based on the nature of the dimerization
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Figure 9. Effect of the reaction temperature on DIB selectivity in a tanksin-series reactor (IB conversion fixed at 90% by suitably changing the catalyst loading).
reaction, that a lower temperature will lead to a lower IB conversion but a higher DIB selectivity and vice versa for a higher temperature. However, this fact does not allow us to select the best-possible operating temperature. Hence, in the following analysis, the variation of the DIB selectivity with temperature was studied using a tanks-in-series reactor model (eight tanks) while maintaining the DIB conversion constant of 90% by suitably changing the catalyst loading. The analysis suggests that, at the same IB conversion, a higher operating temperature always leads to a better DIB selectivity (Figure 9). This result seems a bit surprising, since the general belief is that oligomerization increases at higher temperatures and, hence, should reduce the selectivity. However, it can be reasoned that the use of higher operating temperatures leads to a lower catalyst loading for the same conversion. A lower catalyst loading implies that a lesser number of active catalytic sites are available for further oligomerization reaction and also that there is a better “wettability” of the catalyst because of the larger value of the weight ratio of the polar component to the catalyst. Thus, for the same conversion, the use of a higher temperature and lower catalyst loading seems more advantageous than using a lower temperature and high catalyst loading. Considering the thermal stability limit of the ion-exchange resin catalyst (120 °C), a safe design value of 90 °C is being used to represent the optimum temperature for the reactor unit in all the multiunit configurations considered in this work. Reactor Type. The type of the reactor to be used in the multiunit configurations also needs to be identified. The use of a single liquid phase is recommended for a better flexibility of operation, and to maintain liquid conditions, a sufficiently high pressure (>15 atm) needs to be applied. The simplest reactor model that can be used is a continuous stirred tank reactor (CSTR). For heterogeneously catalyzed reactions, the reactor model is expected to represent a typical industrial packed or fixed-bed reactor. In that case, the obvious choice is a pseudohomogeneous plug-flow reactor (PFR). Another good choice is the tanks-in-series model, which consists of several consecutive CSTRs in which the product of one reactor is led to the next one with or without controlling the intermediate conditions. All the discussed reactor models are compared on the selectivity vs conversion graph at a constant operating temperature of 90 °C, as shown in Figure 10. As expected, the ideal CSTR gives the worst performance and, hence, can be eliminated from all further analysis. The ideal PFR shows the best
Figure 10. Effect of the reactor type on performance (temperature fixed at 90 °C and changing the catalyst loading).
performance, and the performance of the tanks-in-series model is between those of the CSTR and the PFR. However, if a sufficient number of tanks (i.e., 8-10) are used, the performance of the tanks-in-series model is very close to that of the PFR. It was also observed that unsteady-state equations for the tanksin-series model created less convergence problems as compared to the PFR model, which involves partial differential equations in order to predict the variation with respect to time as well as length/volume of the reactor. Another advantage of the tanksin-series model is that it is very convenient to impose a userdefined temperature profile in the reaction unit by assigning the desired constant operating temperatures to the various CSTR units in the reactor train. Because of the above advantages, the eight-tanks-in-series model is used to represent the reactor unit in the multiunit configurations. It is to be noted that the tanksin-series model has also been used by other researchers, in particular, to represent the side reactor.6-7 Separation Unit. Even though the RD column employed in this work operates at a pressure of 15 atm, it is not necessary for the distillation column in multiunit configurations to be operated at such a high pressure. It is commonly known that, as pressure increases, the volatility difference between the components decreases. Therefore, it is advantageous to perform the separation at the lowest acceptable pressure, considering the limit imposed by the coolant used in the condenser. If cooling water at ambient conditions is to be used to minimize operating costs and the top product is almost pure C4, then a pressure of ∼6 atm (or higher) as adopted here was found to be appropriate for a reasonable heat transfer rate in the condenser. Because of the large difference in the volatilities of the reactants (C4) and the products, it is found that a column consisting of even 10 stages is sufficient to separate the C4 from the other products. The separation is based on the predefined specifications for the compositions of the bottom and top streams, i.e., almost the entire C4 cut leaves from the top while the dimer and other oligomers leave from the bottom of the separation unit. The exact numbers of stages in the rectifying and stripping sections depend on the location of feed(s) and side streams, if any. Structure of the Multiunit Configurations The first multiunit configuration considered is the conventional process (Figure 11). The C4 and the polar feeds are
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Figure 11. Conventional reaction-separation configuration with/without recycle.
Figure 13. Comparison of processes on the basis of performance.
Figure 12. Configuration of distillation coupling with side reactor.
introduced into the dimerization reactor operating at high pressure and liquid conditions. The reactor effluent containing the unreacted IB, inert C4, the desired dimer, and the undesired oligomers is led to a distillation column where C4 and the rest of the components are separated as the top and bottom products, respectively. A part of the top product, if necessary, may be recycled back to the reactor to increase the overall conversion. Since the difference between the operating pressure of the reactor and the distillation column is >7 atm and the C4 feed used contains ∼60% inerts, it is very likely that the reactor effluent will flash because of the reduction of pressure. The extent of flash depends on the amount of IB conversion that occurs in the reactor, because the formation of high-molecular products (DIB and oligomers) will increase the bubble-point temperature of the reactor effluent. To account for this eventuality, a flash unit is also placed between the reactor and the distillation column (not shown in Figure 11) which detects the occurrence of the two phases based on the bubble-point temperature of the reactor outlet evaluated at the distillation pressure. The second process alternative considered is the side-reactor configuration (Figure 12). It is to be noted that, the lesser the amount of DIB in the feed to the side reactor, the better will be the overall selectivity, because further loss of selectivity due to the oligomerization of the DIB in the feed can be avoided. It is evident that the top stage, i.e., the condenser, contains the least amount of DIB as compared to any other stage in the column. Hence, it is clear that recycling a part of the top product as the feed to the side reactor will give a better performance than using a side-draw from stages below the condenser. It is found that 10 stages in the distillation column are sufficient for the
separation even if the side reactor was applied. The polar component should always be introduced as a separate feed into the side reactor and not into the column, because if introduced into the column, due to its low volatility, it will never achieve a significant concentration in the top product, a part of which acts as a feed to the side reactor. However, the main C4 feed can be introduced either into the reactor or into the column. Since introducing the main feed into the side reactor will actually be the same as the conventional configuration with recycle, the main feed is introduced into the distillation column for the sidereactor configuration analyzed in this work. The multiunit flowsheets are also solved using the DIVA simulation environment. The modeling philosophy of DIVA takes into account the modular topology of chemical engineering flowsheets. Dynamic models for each of the process units are written and integrated into the overall plant model forming a large system of differential algebraic equations. Comparison of Process Alternatives As stated earlier, the presence of two criteria for performance, i.e., IB conversion and DIB selectivity, makes it difficult to compare the processes. One of the best approaches at such a preliminary stage is to generate potential configurations of the processes at various fixed values of conversion and then compare them on the basis of the best possible selectivity those configurations can achieve. Since, for a given conversion, a process alternative may give different values of selectivity depending upon the choice of design and operating parameters, the parameters need to be altered in such a way that constraints on the top and bottom product composition specifications are never violated, i.e., the dimer composition in the top should not exceed 1 ppm and almost the entire C4 cut is withdrawn from the top. Moreover, in this work, we attempted to generate the best possible combination of design and operating parameters for a particular process and a specified conversion using the aid of repeated simulations and knowledge gained from the trends and behavior of various parameters involved. The work finally results in a set of >10 configurations for each process, each of which is judiciously optimized for a particular IB conversion in the range of 50-95%. The performance of the processes is compared on the DIB selectivity vs IB conversion graph, as shown in Figure 13. It can be seen that, at low conversions (
