Reactive Distillation for FischerâTropsch Synthesis: Simulation-Based

Ind. Eng. Chem. Res. 2010, 49, 9673–9692

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Reactive Distillation for Fischer-Tropsch Synthesis: Simulation-Based Design Methodology Using Aspen Plus Seethamraju Srinivas, Sanjay M. Mahajani, and Ranjan K. Malik* Department of Chemical Engineering, Indian Institute of Technology, Powai, Mumbai 400076, India

In a series of earlier papers, it has been shown through simulations in Aspen Plus that reactive distillation (RD) is feasible for Fischer-Tropsch Synthesis (FTS) [Srinivas et al., Ind. Eng. Chem. Res. 2009, 48, 4710-4718]. The flexibility offered by changing parameters such as reflux ratio, etc. has also been investigated through parametric studies [Srinivas et al., Ind. Eng. Chem. Res. 2009, 48, 4719-4730]. As an extension of the previous works, a methodology is now proposed to design a RD column for FTS in Aspen Plus, utilizing the kinetic and thermodynamic models reported previously. Slurry reactor simulations are performed initially to form a design basis and a simple RD column is configured. This is followed by catalyst redistribution, addition of coolers, nonreactive stages, and side draws. The methodology is illustrated step-by-step for three examples in a systematic manner. The possibility of multiple designs, expected difficulties in execution, and limitations of the algorithm are discussed. Introduction In relevance to the current energy scenario, Fischer-Tropsch synthesis (FTS) is gaining importance to convert coal or biomass to liquid transportation fuels via the syngas route. While most of the improvements in FTS are aimed at modifying the catalyst for improved selectivity, work is in progress by a few researchers to improve product distribution by altering the reactor and/or operating conditions. Reactive distillation (RD) is a well-known process intensification technology that has been demonstrated to be successful in applications such as esterification (MeOAc), etherification (TAME, MTBE), acetalization, etc. The primary motives behind the use of RD include overcoming equilibrium limitations, improving product selectivity by simultaneous separation, and utilizing reaction exotherm in maintaining saturation conditions inside the column. A few other advantages include improved catalyst life, reduction in plant size (because of the combination of reaction and separation), etc. The major benefits of using RD for FTS are expected to be the following: utilizing reaction exotherms to reduce energy consumption; simultaneous product separation that enables side-draw provisions in column sections, giving streams rich in a single fraction such as gasoline, etc.; reduction in load on the downstream processing equipment; and a broader choice of both operating (reflux ratio, side-cooler duties, etc.) and design (number of reactive stages, catalyst loading per stage, etc.) parameters to alter the product yields and selectivity. The option of feeding the reactants at different locations is also possible. It may be noted that the use of RD for systems involving noncondensibles such as H2 has been reported for hydrogenation and desulfurization reactions. Through the use of simulations, the authors have previously demonstrated the feasibility of using RD for FTS1 and carried out parametric studies2 to ascertain the effect of various design and operating parameters on the column performance. The present work involves carrying this study a step further, to propose a methodology for designing such a column, as suggested in our earlier works. Most of the design methodologies for RD reported in literature are based on calculations using the ternary diagram (using the works of Barbosa and Doherty,3 Doherty and Buzad,4 Perez-Cisneros * To whom correspondence should be addressed. Tel.: (022) 2576 7796. Fax: (022) 2572 6895. E-mail: [email protected].

et al.,5 Mahajani,6 and Thery et al.7) or an extension of graphical methods such as McCabe-Thiele to RD (taken from SanchezDaza et al.8). Dragomir and Jobson9 reviewed the work done on the boundary value method and extended it to check the feasibility and determine operating conditions for single-feed kinetically controlled RD columns. Almeida-Rivera et al.10 presented a concise overview of all three existing design methodssnamely, graphical, optimization-based, and heuristicbasedsby discussing the relative merits and demerits of each. Guidelines for designing a RD column using an equilibrium model have been proposed by Bessling et al.11 and Subawalla and Fair;12 however, their methodology is limited to systems that have equilibrium reactions and liquid-phase reactants. Contrary to the usual practice, Melles et al.13 relaxed a few conventional assumptions by considering different tray holdups per column section, heats of reaction, and a nonzero stoichiometric sum to design a RD column, and they illustrated it for a ternary system with ideal vapor-liquid equilibria (VLE). The boundary-value design methodology has some important assumptions that do not hold true in the case of the FischerTropsch reaction system. The important assumptions that must be relaxed are the following: (i) Because of the high heat of reaction, the constant molal overflow assumption cannot be applied. (ii) Feed is in the form of gas, which usually has not been considered. However, Viveros-Garcia et al.14 presented an interesting methodology of conceptual design for RD using a ternary diagram for a hydrodesulfurization system. This considers at least one feed component to be in the gaseous phasesnamely, H2swith the second feed being a liquid stream, and the ternary diagram is plotted on the basis of weight fractions and pseudocomponents, rather than mole fractions and actual components. Cardenas-Guerra et al.15 analyzed the operational and multiplicity aspects for a similar type of system using Aspen Plus. In FTS, both of the feed components are in the gaseous phase. (iii) For systems with more than three components, a transformation of variables has been used to reduce the composition space and provide insights. However, the large number of components in FTS might make a

10.1021/ie100108p  2010 American Chemical Society Published on Web 09/10/2010

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Ind. Eng. Chem. Res., Vol. 49, No. 20, 2010

transformation of variables a difficult task. To add to this, the transformed variable concept is useful for fast and equilibrium reactions, whereas the FT reactions are irreversible and kinetically controlled. (iv) Most of the systems that have been addressed in the literature do not consider the presence of a vapor distillate, which, however, is inevitable in the FischerTropsch reaction scheme. Thus, the design of a RD column for FTS at the conceptual level itself needs the help of a simulator to consider the effects stated previously in assumptions (i)-(iv). The simulation-based design is also a challenging task because of the complexity of the system (detailed nonlinear kinetics with a large number of reactions and components, phase equilibrium calculations involving a wide boiling range, viz., noncondensibles to C30 paraffin, choice of the proper location of side draws and side coolers, etc.) and the convergence issues such as the generation of a proper set of initial estimates, the possibility of multiple solutions and varying degrees of stiffness of each independent variable (such as pressure, reflux ratio, etc.) to performance parameters (such as conversion, product selectivity, etc). Given suitable inputs such as feed throughput, product yields, and selectivity, etc. and constraints such as maximum-possible catalyst loading, purity of the products, etc., the design methodology is expected to provide a configuration specifying the number of stages (reactive and nonreactive) and their appropriate location; the number of coolers, in addition to their duties and location; the number and rate of side draws; the reflux ratio; etc. In the following sections, such an algorithm is proposed to design a RD column for FTS with the minimumpossible inputs and is illustrated with examples. In this case, the design basis is defined through targets on conversion and the chain-growth parameter (R), the latter of which is defined as the probability of a given alkyl species being formed during the course of the reaction to add itself to an existing chain and grow further, as opposed to termination into a product. Thus, the higher the value of R, the greater the molecular weight of the product formed. As R varies from 0 to 1, the product distribution shifts from methane to waxes. It is sufficient to choose R to obtain a desired product distribution in the first approximation of the design. As a first step, slurry reactor simulations are performed to generate data for the design basis as described in the next section. In the second step, a simple configuration is evolved for FTS in RD that closely follows the target sets. In the subsequent steps, the number of reactive stages is increased and coolers are added to address the design constraints/ inputs. In the last step, nonreactive trays and side draws are added as per the necessity to complete the column design. A schematic of the envisaged column configuration is shown in Figure 1. The process simulator, Aspen Plus,16 is chosen for this exercise, and its choice has been justified in our earlier work.17 Reaction kinetics and phase equilibrium model are the two important inputs needed to simulate a RD column. The reaction kinetics used is the model proposed by Wang et al.,18 and the phase equilibrium calculations are addressed through the use of the “PRMHV2” model, which is the same as that reported in the work of Marano and Holder.19 A short discussion on the kinetic and VLE models used can be found in Srinivas et al.1 Slurry Reactor Simulations The objective of these simulations is to generate a basis for the design. The Fischer-Tropsch slurry reactor, which is simulated using the “RCSTR” module, is considered to be an isothermal reactor operating at 25 atm and 250 °C, having a

Figure 1. Hybrid reactive distillation (RD) column with side draws and side duties. 1. Syngas feed; 2. vapor distillate; 3. liquid distillate; 4. side draw; 5. bottoms product; 6. non-reactive stage; 7. reactive stage; 8. side cooler.

Figure 2. Schematic of the slurry reactor.

volume of 24.53 L with 4 kg of catalyst. Two streams exit the reactor: a vapor stream and another stream that consists of liquid and free water (see Figure 2). In addition to the syngas feed, a second feed stream that contains solvent is also fed continuously. The solvent fed is the C16 paraffin, and its flow rate is set to the minimum needed to ensure “slurry” conditions in the reactor during simulations. Vapor-liquid-free water calculations are enabled in all the simulations to address water, which is a major reaction product. The feed flowrate to all of the simulations was 1000 mol/h of syngas (CO + H2) at 250 °C and 25 atm, with a H2/CO ratio of 2.03. Simulations are performed at different temperatures in the range of 150-400 °C. We note here that the reported operating temperatures for Fischer-Tropsch reactors lie within a temperature range of 200-300 °C. Temperatures within the chosen range are considered to cover the entire spectrum for R ) 0-1. Furthermore, at a constant feed rate of 1000 mol/h, five different catalyst loadings are used: 4, 8, 12, 16, and 20 kg, which represent different space velocities. Figure 3 depicts the results of these simulations. The values shown in the legend in Figure 3 represent the space velocity, expressed in units of (m3/day)/kg. As expected, in Figure 3a, conversion increases with increasing reactor temperature, as well as with decreasing space


Figure 3. (a) Conversion as a function of temperature at different space velocities (expressed in units of (m3/day)/kg). (b) Variation between R and conversion at different space velocities (expressed in units of (m3/day)/ kg).

Figure 4. Typical Fischer-Tropsch product spectrum from Aspen CSTR simulations.

velocity. The change in conversion at longer residence times is not appreciable. R is a function of temperature, and hence, we plot R vs conversion in Figure 3b. Note that conversion varies as R changes, with the variation being steeper at higher values of R. Also, once the catalyst loading is fixed, there is a oneto-one correspondence between conversion and R. The design basis can now be set as follows. Based on the desired product distribution, one can choose a value for R from Figure 4. Once R is fixed, the relative amounts of products such as gasoline (C5-C11), diesel (C12-C18), light wax (C19-C23), medium wax (C24-C35), and heavy wax (C35-C50) can be predicted. RD can be considered as a cascade of reactive flashes, viz, each stage in RD acts as a slurry reactor that has two inlets and two outlets (one each for vapor and liquid), as shown in Figure 2. Hence, Figure 4 is generated by performing CSTR simulations at

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different temperatures with a fixed catalyst loading, to cover the range from R ) 0 to R ) 1. The nature of the product spectrum in Figure 4 remains the same when a fixed-bed reactor is used for the analysis. A target on the conversion can then be set, depending on the end-use for the unreacted syngas, such as power generation, use as a fuel, or use as a raw material for the production of other chemicals or recycle back to the main reactor. Once the conversion and R values are chosen, Figure 3b can be used to determine the residence time needed and, hence, the catalyst requirement for a given feed rate. Finally, Figure 3a can be used to set the reactor temperature needed using the values of conversion and space velocity obtained in an earlier step. The idea is to try and operate the RD column for FTS at an average temperature that is as close as possible to this set value from Figure 3a. This may be achieved by manipulating parameters such as reflux ratio, pressure, duties of side coolers, number of reactive stages, etc. The important conclusions from an earlier parametric study for FTS in RD2 are as follows: • An increase in reflux ratio helps to increase the conversion at the expense of liquid yields, i.e., the product spectrum shifts from heavy products to light products. • The nature of the catalyst distribution also affects the conversion and product selectivity. • With an increase in the number of reactive stages with the catalyst weight fixed, there is a change in conversion and product selectivity. • With an increase in heat removal, there is a slight decrease in the corresponding stage temperature, leading to a decrease in conversion and an increase in heavy product selectivity. • There is no significant change in conversion or product selectivity on altering side draws or by adding nonreactive stages below the last reactive stage in the column; it only changes the purity of the bottom product. In a CSTR, three parameters can be used to alter conversion: reactor temperature, pressure, and inlet H2/CO ratio. A discussion on sensitivity analysis for these three parameters is given in Section A of the Supporting Information. It is concluded that temperature is the dominant parameter among the three that strongly affects conversion and product distribution. Hence, note that the emphasis in this article is on the average reactive stage temperature of the RD column. Furthermore, each stage in RD will have an R value that is representative of the temperature on that stage. Hence, the net product distribution in RD will be a combination of many such R values and cannot be described by a unique value, as is usually done in the case of a CSTR. The enhancement in reaction due to separation is difficult to interpret, because of the complex interaction between temperature, catalyst loading, and H2/CO ratio (see the work of Srinivas et al.17) and is not discussed further. Design Basis. Two examples are considered in this work with the following design basis and constraints: Design1: R ) 0.92, Xco ) 52.09 ( 5%, 500 000 m3/day of feed at 25 atm and 250 °C. Design2: R ) 0.84, Xco ) 81.6 ( 5%, 50 000 m3/day of feed at 25 atm and 250 °C. Both designs in the RD column have an inlet feed ratio (H2/ CO) of 2.03. The maximum-permissible catalyst weight per stage is assumed to be 1000 kg and it is desired to have the maximum possible purity of a single fraction (for, e.g., gasoline) in a side draw. The column pressure is set at 25 atm and the condenser at 35 °C. This leaves us with a reflux ratio that can be manipulated as required to meet the design targets, besides adjusting the side-cooler duties and side-draw rates. Changes

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in product selectivity by altering the catalyst distribution, with the total catalyst weight remaining constant, has been shown in our earlier work.2 However, since this is a first attempt to design a RD column for FTS, uniform catalyst loading per stage is assumed in this work to keep the algorithm simple and to limit the number of possible solutions that must be explored. Furthermore, the design problem statement in our case is rather different, because we have not specified the production rate of gasoline, diesel, etc.; however, it is based on feed conditions, conversion, and R. However, note that the given data in the design basis is sufficient to calculate the product rates. Knowing the feed rate and conversion will give the amount of hydrocarbon products formed. Since R is fixed, the product distribution is known. Both of these factors together can be used to find the approximate product yields. Alternatively, defining the total amount of liquid products and R in the problem statement, one can find the amount of CO (in moles) needed for generation of each component. Assuming a value for conversion based on other factors (such as the end-use of syngas, etc.), the net molar flow rate and, hence, the volumetric feed rate of syngas can be calculated. In addition, the catalyst requirement can be calculated using Figure 3. Furthermore, one may compare this design problem with the crude column design in a refinery where design calculations are done from left to right. In other words, given a specific quality and quantity of crude and its assay, the column is configured to obtain a desired product slate. Note that per-pass conversions for industrial-scale FischerTropsch reactors are not available in the literature, to the best of the authors’ knowledge. The conversion values chosen here are for the purpose of illustration only. Design1 is so chosen as to have a reasonable feed flow rate, conversion, and good yields of liquid fuels, which will enable us to predict the typical size of a RD column that one can expect under situations of practical interest. Figure 3b suggests that the increase in conversion for a fixed R and feed flow rate is only incremental, despite a considerable increase in the catalyst weight (residence time). To verify the difficulty in configuring a column that has a high residence time (∼3 times that of Design1) and almost similar product distribution as Design1, Design2 is chosen to have a high conversion. The feed rate is reduced in Design2 to keep the catalyst weight low. Corresponding to the conversion and R values that are chosen, Figure 3a indicates that a temperature of ∼225 °C must be maintained in the reactive zone of the column to attain the desired product selectivity. Figure 3b is used to calculate the catalyst weights needed for both the designs, which correspond to 48 and 14 metric tons, respectively. Note that Dry20 mentioned a catalyst charge of 40 m3 and a design with a fresh feed space velocity of 500 h-1 in the first fixed-bed reactors of SASOL (∼500 000 m3/day and 32 metric tons of catalyst, assuming a catalyst density of 0.8 g/cm3). Design Algorithm. The following steps are proposed to evolve a RD column that can meet the specified objectives. Note that this algorithm is proposed to obtain the feasible design at a conceptual level. It is modified further at the end of this paper, based on some practical considerations: (1) Start with a simple configuration that has no side duties, side draws, or nonreactive stages, with the feed rate and catalyst weight decided as the base. The number of reactive stages can also be kept at the minimum possible, to avoid convergence problems.1 (2) Increase the number of reactive stages and progressively decrease the catalyst loading per stage until the maximum desired loading on each stage is achieved. In this case, the desired loading is assumed to be 1000 kg per tray.

(3) Calculate the average reactive tray temperature for the configuration at the end of Step 2. If it is close to the value chosen as the basis, go to Step 4. Otherwise, add heaters/coolers to the stages gradually, in steps, so that the desired temperature can be achieved. (4) Examine the composition profiles in the fully reactive column at the end of Step 3. If there are sections in the column from which side draws with a desired purity for each fraction can be removed, go to Step 5. Otherwise, add nonreactive stages at the proper location(s) to enrich the composition of the desired fraction. (5) Add side draws on stages where the purity of the desired fraction meets the specifications. The rate of removal of the side draws can be calculated from the product distribution of all of the outlet streams together. During the course of this evolution, the conversion and product distribution may move away from the desired values. In such cases, parameters such as reflux ratio, location and number of coolers added, etc. can be manipulated to bring the design objectives close to their target values. The basis for this manipulation is the parametric study performed in the earlier work.2 Modification of the catalyst distribution also helps to attain the design objectives; however, that has not been addressed in this work, for reasons stated earlier. In the following sections, the evolution for Design1 and Design2 are illustrated. Illustration 1: Design 1 The first case chosen has a syngas feed rate of 500 000 m3/ day with 48 000 kg of catalyst. The feed ratio (H2/CO) is 2.03. This corresponds to a base case of 1000 mol/h of feed rate and 4 kg catalyst, for which a solution already exists.1 Since the pressure of the existing solution is 25 atm, the new design is also developed at a pressure of 25 atm. Besides, sensitivity analysis that has been reported elsewhere21 predicted that the maximum number of solutions is possible in the pressure range of 24-26 atm. Step 1. It is advisible to start with a column that has a small feed rate (1000 mol/h). Hence, the flowsheet with an existing solution in our case was modified to reflect the changes in the feed flow and catalyst distribution by increasing them proportionately so that the Damko¨hler number in either case is constant. The catalyst distribution in the lower feed case is 1 kg per stage on stages 2-5. Forty eight metric tons of catalyst in the higher feed case is distributed in the same proportion, giving 12000 kg of catalyst each on four reactive stages. Convergence is observed at the same reflux ratio as the base case. The column consists of 6 stages in total, with 2-5 being reactive stages. The conversion of CO is 55.65%, the average reactive stage temperature is 245.24 °C, and R ) 0.817. While the conversion is more than the design target, R is lower than the target value, because of the higher temperature. We name this configuration as Design11. Step 2. Reactive stages are added in this step to reduce the catalyst loading per stage to the desired value. Figure 5 and Table 1 represent the evolution as reactive stages are added to Design 11. The configuration resulting at the end of this step is termed Design12. The terms shown in the legend in Figure 5 indicate the per stage catalyst loading. For example, D2000 represents a column that has 2000 kg of catalyst per stage on 24 reactive stages, with a total of 26 stages. Note that the partial vapor-liquid (V-L) condenser is included as a stage. Also, D12000 corresponds to Design11 and D1000 to Design12. The design basis restricts the maximum catalyst loading possible on a single stage to 1000 kg. Design12 has 48 reactive stages,


Figure 5. (a) Temperature profile evolution for Design11 to Design12. (b) Reaction rate (CO) profile evolution for Design11 to Design12. (c) R profile evolution for Design11 to Design12.

with 1 metric ton of catalyst on each stage. While evolving the design, it is further necessary to devise some strategy to avoid convergence problems. Consider the following ways in which Design11 can be modified: (i) One can add 44 nonreactive stages and then try shifting the excess catalyst from the top stages to the bottom stages.

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(ii) Add one stage per step and shift 1000 kg onto this stage via removal of the entire weight from one of the top stages alone. (iii) Add one stage per step and shift 1000 kg onto this stage via the equal removal of weights from the top stages having excess catalyst. (iv) Add stages in each step and redistribute the total catalyst loading, so that all the stages, including the new stage added, have equal catalyst weights on them. It is observed that option (iv) works the best in evolving Design12 from Design11. During the addition of stages and the redistribution of catalyst, the reflux ratio was manipulated to obtain convergence while keeping the pressure fixed at 25 atm. As the catalyst distribution per stage decreases, it can be seen from Figure 5a that the maximum temperature in the column increases, and the minimum temperature decreases. The average reactive stage temperature (Table 1) goes through a maximum and finally reaches a value of 241.4 °C in Design12. The reaction rates have also become more uniform in the 50-stage column, compared to the 6-stage column, as is evident from Figure 5b. This sounds interesting, as well as contradictory, if one were to look at the temperature profile in the case of Design12. It is suggested that the number of reactive stages is equal to the total quantity of catalyst divided by the maximum catalyst loading permitted per stage. Despite the top section of the column in Design12 having temperatures in excess of 250 °C and the lower section having temperatures of 85), compared to a minimum recommended value of 51. Figure 11 compares the ASTM D86 curves for the cases in Table 8 to those of straight-run naphtha (SRN), heavy naphtha (HN), and medium gas oil (MGO) for a typical crude processed in a refinery. The gasoline fractions from FTS in Figure 11a are similar enough to the HN fraction. Thus, this stream can possibly be blended with the HN stream and processed further to get a high-octane gasoline product. The Fischer-Tropsch diesel component in Figure 11b lies quite close to the MGO fraction. Therefore, it can act as a diesel blend, in combination with the MGO cut for further finishing. Note that there is a considerable difference in Figure 11b at the 5% distill-off point, because of the presence of light hydrocarbons in the diesel fraction, which are normally isolated from the MGO cut by the use of a side stripper. To summarize, the fractions from FTS in this study that are termed “gasoline” and “diesel” are not products by themselves, but possess properties (MABP, API gravity, ASTM D86) that are comparable to those of typical refinery cuts and can be treated as “blends” leading to the respective products. Practical Aspects Column Sizing. To gain some insight into the practical feasibility of the designs illustrated, column sizing calculations were performed using Aspen Plus. The default settings were made use of with a tray spacing of 2 ft, a fractional approach to flooding of 0.8, a foaming factor of 1, an overdesign factor of 1, and a minimum downcomer area (fraction of total area)

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Table 7. (a) Important Parameters, (b) Product Selectivity, and (c) Draw Rates and Purity before and after Reducing the Number of Coolers (a) Important Parameters before and after Reducing the Number of Coolers Temperature (°C)

D15 D15alt D15RxnRate D15RxnRateAlt D16 D16alt D16RxnRate D16RxnRateAlt D16LowCat D16AltLowCat

Liquid Yields (kg/h)

reflux ratio

Xco (%)

Qc (MW)

avg reac

max

min

gasoline

diesel

wax

0.113 0.072 0.11 0.106 0.113 0.072 0.11 0.112 0.113 0.072

54.31 47.8 53.74 53.39 54.27 48.09 54.25 54.24 54.15 48.14

-28.31 -21.49 -28.48 -27.42 -28.26 -21.82 -28.35 -28.3 -28.39 -22.03

223.86 217.76 224.73 224.54 223.78 217.95 225.29 225.3 224.19 218.53

245.84 242.79 241.43 240.02 247.68 246.97 244.9 244.84 247.72 247

156.12 146.98 158.48 155.99 155.89 148.81 158.11 156.69 156.35 150.05

1976 1619 1699 1685 1973 1776 1760 1689 2143 1961

4708 4260 4988 4857 4671 4089 4885 4891 4636 4062

3102 3160 3384 3435 3056 3018 3350 3324 2964 2945

(b) Product Selectivity before and after Reducing the Number of Coolers Liquid Selectivity (%)


Net Product Selectivity (%)

gasoline

diesel

wax

light gas

naphtha

gasoline

diesel

wax

alpha

20.18 17.89 16.82 16.88 20.33 19.89 17.54 17.02 21.9 21.68

48.07 47.05 49.4 48.65 48.12 45.76 48.69 49.36 47.38 44.92

31.67 34.9 33.51 34.4 31.48 33.8 33.38 33.54 30.29 32.57

13.03 10.62 11.73 11.63 13.36 11.86 12.5 12.7 13.43 11.89

15 13.16 13.88 13.87 15.18 14.04 14.39 14.52 15.29 14.09

20.45 20.27 18.51 18.75 20.45 20.93 18.69 18.49 21.38 21.95

31.06 32.12 33.29 32.66 30.83 30.58 32.28 32.32 30.44 30.18

20.47 23.83 22.59 23.09 20.17 22.59 22.13 21.97 19.46 21.88

0.927 0.943 0.928 0.93 0.926 0.941 0.926 0.925 0.918 0.933

(c) Draw Rates and Purity before and after Reducing the Number of Coolers Side-Draw Stage


Draw Rate (kg/h)

Purity (mass %)

number of coolers

gasoline

diesel

gasoline

diesel

gasoline

diesel

wax

25 32 23 47 8 7 7 8 8 7

6 6 6 6 6 6 6 6 6 6

66 69 66 62 66 69 66 62 66 69

2000 1500 1600 1700 2000 1500 1600 1700 2000 1500

4800 4800 5000 4900 4800 4800 5000 4900 4800 4800

99.42 96.19 99.44 99.46 99.41 99.49 99.45 99.46 99.54 99.68

88.01 72.57 87.87 89.4 87.92 72.37 87.31 89.84 87.93 72.47

82.78 69.99 81.28 84.8 83.42 72.19 82.79 84.11 83.93 72.25

as 0.1 for different tray types. The range of the diameter calculated in all the cases is 3.8-5.5 m, which is a reasonable value. The liquid velocity on the trays varies between 0.15 cm/s to 1.63 cm/s, while the vapor velocities are in the range of 0.6-1.12 m/s. Figure 12 depicts the profiles for the column diameter and the molar flow rates of liquid and vapor along its height. All these values seem reasonably justified, compared to the results for a typical crude column (see the work of Seo et al.25). For the sake of completeness, it is noted that the selection of tray internals for FTS in RD will be a difficult task, because of the small size of the catalyst (which would make it tricky to be packed into catalyst bags, etc.) and its separation from the liquid stream. The tray type envisaged is possibly the design suggested by Liang et al.,26 while liquid (wax)-catalyst separation is discussed in a review on liquid-phase FTS reactors by Davis.27 Side Coolers. Table 9 shows the cooling duties and corresponding tray temperatures for four designs. The temperature varies between ∼200 °C and 240 °C. Maintaining this temperature on the stages is expected through the use of cooling coils to generate steam, preferably medium pressure steam. Approximate calculations performed showed that the number of turns for a coil-type heat exchanger (submerged in liquid on a tray) located inside the column is quite large and appears infeasible. Therefore, the other option is to use an external heat

exchanger to remove the exothermic heat. This is similar to a pump-around exchanger being used in the conventional crude columns, although for a different purpose. Since Aspen Plus permits the use of pump-arounds in a RadFrac module, simulations were performed using Design15RxnRateAlt to study the affect of pump-around on column performance. Each cooler already established in the configuration was removed, and a pump-around with an equivalent duty as that of the cooler was included. For the purpose of simulation, the withdrawal and return stages for the pump-around are specified to be the same. In practice, however, the return stage is two or three stages above the withdrawal stage and the trays between are designated as “heat-transfer stages”. The return temperature of the liquid to the column is specified to be 20 °C lower than the corresponding draw temperature in one case, and 180 °C for all the draws, irrespective of the draw temperature in another case. In both cases, it is observed that the column performance, in terms of conversion and product distribution, is not hampered. Only the internal liquid flows on the trays having pump-around draws get adjusted. As expected, when the draw return temperature is set to 180 °C in all the cases, the pump-around draw rate is low, because of a higher ∆T value. The design of shell-andtube exchangers for such purposes is widely practiced and would prove easier than coil-type exchangers. To conclude, one can


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Figure 10. Recommended final design algorithm.

have suitably designed pump-around exchangers to act as side coolers in the final column configuration. Conclusions and Remarks The proposed design algorithm has shown different rates of success in the examples considered. Illustration 1 proved to be the most successful; it demonstrated the flexibility and met the design targets, giving side draws considerably rich in each fraction. Illustration 2, which had a slightly different heater configuration, failed to meet the design target for conversion. Illustration 3, which was used for a high-conversion case, did not exactly match the design criteria. Furthermore, it proved difficult to increase the sidedraw rates, which affected the product purity. It may be inferred

that Step 3 of the proposed algorithm, where the heat exchangers are added, is an important decision variable that affects the latter steps. A few practical aspects, such as the product composition and their capability to be blended with refinery streams, a reduction in the number of coolers, the column diameters expected, etc. have also been considered and are commented upon. One may take note of the fact that the results of the simulation should be interpreted with caution, because of the possibility of multiple solutions, which were not encountered in this study, especially during the addition of side coolers. Although olefin readsorption is a part of the kinetic model used, the olefin-to-paraffin ratio is not considered to be one of the performance parameters in developing the algorithm. This is because of the partial-pressure-based model used, which

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Table 8. Typical (a) Gasoline Fraction Properties and (b) Diesel Fraction Properties (a) Typical Gasoline Fraction Properties

API flash point-API MABPa K-Watson gasoline purity gasoline composition C8 (within gasoline) C9 (within gasoline) C10 (within gasoline) C11 (within gasoline) C12 (within gasoline)

Case 1

Case 2

Case 3

Case 4

Case 5

59.69 59.93 °C 196.83 °C 12.78 98.20 mass %

61.44 43.55 °C 178.69 °C 12.73 98.14 mass %

62.30 45.26 °C 184.94 °C 12.84 97.84 mass %

61.28 48.52 °C 180.70 °C 12.74 98.36 mass %

60.68 50.60 °C 186.45 °C 12.75 98.36 mass %

0.33% 2.29% 17.28% 45.79% 34.30%

3.21% 17.39% 30.79% 31.02% 17.60%

4.94% 20.39% 24.61% 25.18% 24.89%

0.63% 7.80% 41.73% 47.18% 2.66%

0.59% 7.80% 33.13% 38.20% 20.28%

(b) Typical Diesel Fraction Properties

API cetane number flash point-API MABP K-Watson diesel purity diesel composition C13 (within diesel) C14 (within diesel) C15 (within diesel) C16 (within diesel) C17 (within diesel) C18 (within diesel) a

Case 1

Case 2

Case 3

Case 4

Case 5

53.16 85 90.76 °C 244.25 °C 12.75 93.10 mass %

50.60 105 100.57 °C 275.30 °C 12.81 84.46 mass %

51.32 98 99.45 °C 267.28 °C 12.80 92.23 mass %

52.04 90 91.75 °C 259.28 °C 12.79 75.99 mass %

50.38 107 100.54 °C 281.06 °C 12.84 75.04 mass %

53.56% 29.36% 9.72% 3.58% 2.13% 1.65%

16.64% 18.97% 18.55% 17.73% 16.00% 12.11%

16.71% 24.83% 20.98% 17.10% 12.82% 7.57%

29.91% 23.34% 17.44% 12.71% 9.49% 7.11%

20.02% 19.59% 18.13% 16.15% 14.12% 11.99%

Mean average boiling point.

Figure 11. ASTM D86 comparisons: (a) gasoline fraction and (b) diesel fraction.

creates a few anomalies, as reported earlier by Srinivas et al.,1 and is one of the limitations of the algorithm. Future work to consider the olefin-to-paraffin ratio as a performance parameter may consider the use of kinetic models based on concentration rather than partial pressure. The distillation effect on reaction

Figure 12. Typical (a) column diameter profile and (b) molar flow rate profile (liquid and vapor).

is also expected to be more pronounced in such a case. The average reactive stage temperature used here is the arithmetic

Ind. Eng. Chem. Res., Vol. 49, No. 20, 2010 Table 9. Representative Temperature and Cooling Duty Profiles Design16

Design16alt

Design16RxnRate Design16RxnRateAlt

T (°C) Q (MW) T (°C) Q (MW) T (°C) Q (MW) 203 219 228 233 234 238 239 241

-6 -6 -6 -5.5 -7 -6.5 -7.5 -5.5

199 214 207 226 232 237 235

-6 -8.5 -7 -6 -7 -8 -7.5

205 234 228 221 220 220 207

-4.8 -6.5 -7.6 -9.1 -10 -6.6 -5.3

T (°C)

Q (MW)

203 236 233 226 223 224 217 207

-5.8 -5.5 -6.3 -7.3 -8.4 -7.7 -5.7 -3.5

average. Similarly, the R value calculated for each case is the arithmetic average over the reactive stages. While this served the purpose, these can be improved further to take into consideration the catalyst mass (which was not needed here, because of the equal catalyst loading) and the reaction rates on each stage. The algorithm presented is indeed tedious to work on, because of the number of simulations needed in each step. This is attributed to the complexity of the problem and the difficulty in generating a good set of initial estimates. It is observed that providing proper initial estimates for temperature and compositions (both liquid and vapor) is necessary to obtain convergence. In doing repeated simulations, the solution from the earlier case is used as a guess after the next modification and, hence, makes the process simpler, although time-consuming and tedious. The proposed algorithm does not consider a design with stages that have unequal catalyst loading per stage and can be extended to consider such scenarios. While the reaction rate profile suggested the cooling duties as a first approximation, a better method of combining the coolers to reduce their number can be worked upon. Also, because of the importance of the average reactive stage temperature, aspects related to temperature controllability can be a future part of this work. Apart from the design illustrations described in this work, two more independent cases were attempted, using the proposed design methodology, whose details can be found elsewhere.28 Note that the following common points also have been observed in both of these examples: • Different column configurations are possible at the end of each step that could give similar results but with unfulfilled design criterion. • The step involving the addition or removal of heat-transfer devices is the most important step and consumes the maximum effort. • The iterative nature of the algorithm is also realized. Problems in configuring the column using the proposed methodology are experienced when the targeted average reactive stage temperature or conversion is higher, which indicates that designing for such cases is a difficult task and may require further attention. Often, a change in strategy midway through the execution of a single step of the proposed design methodology may be necessary, depending on the problem. It is also to be noted that combining two steps of the proposed design methodology (e.g., simultaneous addition of both side-draws and nonreactive stages) sometimes proves to be helpful. Thus, the revised algorithm recommended is only a first step toward designing a reactive distillation (RD) column for Fischer-Tropsch synthesis (FTS) and requires further improvements. Although this methodology does not guarantee an optimal RD design, the base case so obtained can be used for further optimization studies.

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Supporting Information Available: The following sections are available as part of the Supporting Information: CSTR sensitivity analysis, physical significance of the heater on stage 2, details of illustrations 2 and 3, and depiction of the reduction in the number of coolers. This information is available free of charge via the Internet at http://pubs.acs.org. Acknowledgment The authors wish to thank all the reviewers of this series of five papers (on Fischer-Tropsch synthesis (FTS) in reactive distillation (RD)) for their valuable suggestions, critical comments, and encouragement. Literature Cited (1) Srinivas, S.; Malik, R. K.; Mahajani, S. M. Feasibility of Reactive Distillation for Fischer-Tropsch Synthesis. 2. Ind. Eng. Chem. Res. 2009, 48, 4710–4718. (2) Srinivas, S.; Malik, R. K.; Mahajani, S. M. Feasibility of Reactive Distillation for Fischer-Tropsch Synthesis. 3. Ind. Eng. Chem. Res. 2009, 48, 4719–4730. (3) Barbosa, D.; Doherty, M. F. Design and Minimum Reflux Calculations for Single-Feed Multicomponent Reactive Distillation Columns. Chem. Eng. Sci. 1988, 43, 1523–1537. (4) Doherty, M. F.; Buzad, G. New tools for the design of kinetically controlled reactive distillation columns. Comput. Chem. Eng. 1994, 18, S1– S13. (5) Perez-Cisneros, E.; Schenk, M.; Gani, R.; Pilavachi, P. A. Aspects of Simulation, Design and Analysis of Reactive Distillation Operations. Comput. Chem. Eng. 1996, 20, S267–S272. (6) Mahajani, S. M. Design of Reactive Distillation Columns for Multicomponent Kinetically controlled Reactive systems. Chem. Eng. Sci. 1999, 54, 1425–1430. (7) Thery, R.; Meyer, X. M.; Joulia, X.; Meyer, M. Preliminary Design of Reactive Distillation Columns. Chem. Eng. Res. Des. 2005, 83, 379–400. (8) Sanchez-Daza, O.; Perez-Cisneros, E. S.; Bek-Pedersen, E.; Gani, R. Graphical and Stage-to-Stage Methods for Reactive Distillation Column Design. AIChE J. 2003, 49, 2822–2841. (9) Dragomir, M. R.; Jobson, M. Conceptual design of single-feed kinetically controlled reactive distillation columns. Chem. Eng. Sci. 2005, 60, 5049–5068. (10) Almeida-Rivera, C. P.; Swinkels, P. L. J.; Grievink, J. Designing reactive distillation processes: present and future. Comput. Chem. Eng. 2004, 28, 1997–2020. (11) Bessling, B.; Schembecker, G.; Simmrock, K. H. Design of Processes with Reactive Distillation Line Diagrams. Ind. Eng. Chem. Res. 1997, 36, 3032–3042. (12) Subawalla, H.; Fair, J. R. Design Guidelines for Solid-Catalyzed Reactive Distillation Systems. Ind. Eng. Chem. Res. 1999, 38, 3696–3709. (13) Melles, S.; Grievink, J.; Schrans, S. M. Optimisation of the conceptual design of reactive distillation columns. Chem. Eng. Sci. 2000, 55, 2089–2097. (14) Viveros-Garcia, T.; Ochoa-Tapia, J. A.; Lobo-Oehmichen, R.; Reyes-Heredia, J. A. De los; Perez-Cisneros, E. S. Conceptual Design of a reactive distillation process for ultra-low sulfur diesel production. Chem. Eng. J. 2005, 106, 119–131. (15) Cardenas-Guerra, J. C.; Lopez-Arenas, T.; Lobo-Oehmichen, R.; Perez-Cisneros, E. S. A reactive distillation process for deep hydrodesulfurization of diesel: Multiplicity and operation aspects. Comput. Chem. Eng. 2010, 34, 196–209. (16) Aspen Plus User Manual, Aspen Plus Version 2004.1; Aspen Technologies Inc.: Cambridge, MA, 2004. (17) Srinivas, S.; Malik, R. K.; Mahajani, S. M. Feasibility of Reactive Distillation for Fischer-Tropsch Synthesis. Ind. Eng. Chem. Res. 2008, 47, 889–899. (18) Wang, Y. N.; Ma, W. P.; Lu, Y. J.; Yang, J.; Xu, Y. Y.; Xiang, H. W.; Li, Y. N.; Zhao, Y. L.; Zhang, B. J. Kinetics modeling of FTS over an industrial Fe-Cu-K catalyst. Fuel 2003, 82, 195–213. (19) Marano, J. J.; Holder, G. D. Characterization of FT liquids for VLE calculations. Fluid Phase Equilib. 1997, 138, 1–21. (20) Dry, M. E. In Catalysis: Science and Technology, Vol. 1; Anderson, J. R., Boudart, M., Eds.; Springer Verlag: Berlin, 1981; pp 163. (21) Srinivas, S.; Mahajani, S. M.; Malik, R. K. Reactive Distillation for Fischer-Tropsch Synthesis: Feasible solution space, Ind. Eng. Chem. Res. 2010 , 49, 6350-6361.

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(22) de Klerk, A.; de Vaal, P. L. Alkylate Technology Selection for Fischer-Tropsch Syncrude Refining. Ind. Eng. Chem. Res. 2008, 47, 6870– 6877. (23) Leckel, D. Diesel Production from Fischer-Tropsch: The Past, the Present, and New Concepts. Energy Fuels 2009, 23, 2342–2358. (24) Kamara, B. I.; Coetzee, J. Overview of High-Temperature FischerTropsch Gasoline and Diesel Quality. Energy Fuels 2009, 23, 2242–2247. (25) Seo, J. W.; Oh, M.; Lee, T. H. Design Optimization of Crude Oil Distillation. Chem. Eng. Technol. 2000, 23, 157–164. (26) Liang, W.; Wang, Z.; Jin, Y.; Yu, Z.; Yang, S. Performance of a Three-Phase Fluidized Bed as a Reactive Distillation Device. Chem. Eng. Technol. 1996, 19, 456–461.

(27) Davis, B. H. Overview of reactors for liquid phase Fischer-Tropsch synthesis. Catal. Today 2002, 71, 249–300. (28) Srinivas, S. Reactive Distillation for Fischer-Tropsch Synthesis: Modeling and Simulation, Ph.D. Thesis, IIT-Bombay, Bombay, India, 2010.

ReceiVed for reView January 17, 2010 ReVised manuscript receiVed July 16, 2010 Accepted August 13, 2010 IE100108P

Reactive Distillation for FischerâTropsch Synthesis: Simulation-Based

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