Evaluation of Reactive Distillation and Side Reactor Configuration for

Jan 7, 2014 - With an iterative algorithm embedded into Aspen Plus and considering the effects of catalyst volume fraction and catalyst effectiveness,...
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Evaluation of Reactive Distillation and Side Reactor Configuration for Direct Hydration of Cyclohexene to Cyclohexanol Jianchu Ye, Jun Li, Yong Sha,* Handan Lin, and Daowei Zhou Department of Chemical Engineering and Biochemical Engineering, College of Chemistry and Chemical Engineering, Xiamen University, Xiamen 361005, Fujian China S Supporting Information *

ABSTRACT: With an iterative algorithm embedded into Aspen Plus and considering the effects of catalyst volume fraction and catalyst effectiveness, the reactive distillation (RD) process for direct hydration of cyclohexene to cyclohexanol is simulated and evaluated. Results indicate that the RD process, compared with the traditional reaction−separation process, is infeasible in light of total annual cost (TAC) though it is theoretically reasonable from a previous study. The infeasibility of the RD process is found to be mainly due to the low catalyst effectiveness, rather than the slow reaction rate or limitation of catalyst loading in the RD column. Accordingly, a side reactor configuration (SRC) for the direct hydration of cyclohexene is proposed and evaluated. Results show the advantages of the SRC with fewer reactors and reveal that the TAC of the SRC with a single side reactor as the best choice for the direct hydration process reduces 11.41% of that of the traditional reaction−separation process.

1. INTRODUCTION Cyclohexanol is an important chemical intermediate mainly used as the precursor of nylon.1 The direct hydration of cyclohexene to cyclohexanol process developed by Asahi Chemical Co. has been concerned with more interest because of the use of partial hydrogenation of benzene to cyclohexene, which can save the consumption of hydrogen and improve the operation safety.2 In the process, a slurry reactor is used with the discharge from the reactor being separated by a decanter into two phases: the organic phase which is further handled in distillation column to obtain cyclohexanol and the aqueous phase which is recycled back to the reactor with catalyst. Besides the normal reaction−separation procedure, a reactive distillation (RD) process for the direct hydration was proposed by Steyer et al.3 and Qi et al.4 as an alternative to the Asahi process with considering the chemical equilibrium limitation of the system. Although feasibility analysis has showed positive results such as the discharge of pure cyclohexanol, the one-step RD process is still considered to make no sense in developing a corresponding realistic process due to the slow hydration rate leading to a high requirement for the amount of catalyst.5 Accordingly, Steyer et al.2,5 proposed a new RD process with two steps: the esterification of cyclohexene with formic acid and the hydrolysis of the ester to cyclohexanol. Although the hydrolysis equilibrium constant for the second step is less than 1, the hydrolysis is still feasible after a careful design.2 A rigorous calculation further revealed the high yield of cyclohexanol in spite of the relatively slow hydrolysis rate.6 Meanwhile, the subsequent experiments carried out in vacuum showed the feasibility of the two-step RD process.7 However, the reported multiplicity in the hydrolysis column and the required recycle of the distillate from this column suggest a relatively narrow operation window for the whole process.6 Despite of many works which have mentioned the kinetic limitation in the one-step RD process with high feasibility, there is still no report stating its irrationality in the process cost. © 2014 American Chemical Society

Generally, lower process cost is possible for RD due to the absence of reactor and the enhanced equilibrium conversion. On the other hand, it is obvious that the direct hydration process has a natural superiority in configuration relative to the two-step process. So the one-step RD process is still considered as a good choice for the cyclohexanol production and evaluated in light of total annual cost (TAC) in this work. In order to conveniently design more reliable RD process, an iterative algorithm with considering the effects of catalyst volume fraction and catalyst effectiveness is embedded into Aspen Plus. Moreover, a side reactor configuration (SRC) for the direct hydration is proposed and evaluated to further decrease the TAC of the direct hydration process. The traditional reaction− separation process is also simulated as the reference of evaluation in this work.

2. SETUP FOR PROCESS SIMULATION AND EVALUATION 2.1. Kinetics and Reactor. Zhang et al.8 studied the direct cyclohexene hydration using a strong acid zeolite (ZSM-5) or an ion-exchange resin (Amberlyst 15) as the catalyst. They showed that the former catalyst performs better than the latter one. The reaction mechanism on the zeolite was proposed as follows. rol =

1 dnol aq = k(xene − xolaq /Ka) mcat dt

(1)

Equation 1 indicates that the reaction occurs in aqueous phase and the reaction rate depends on the aqueous phase concentrations of cyclohexene and cyclohexanol. Zhang et al.8 obtained the kinetic parameters in eq 1 by regression analysis of Received: Revised: Accepted: Published: 1461

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reactor can be determined on the basis of catalyst content and average density of reaction system. According to the Asahi process,12 the catalyst content is set to 30 wt % of the reaction mixture, and the volume ratio of the aqueous phase to the organic phase is 1.8:1, providing an average density of 890 kg/ m3 for the system. Note that the aqueous phase includes the catalyst. 2.2. Total Annual Cost. For process evaluation, the capacity of the hydration process is fixed at 60 000 t/a for cyclohexanol, and the total annual cost (TAC) is calculated with

their experimental data from a batch slurry reactor. In the regression, the concentration of dissolved cyclohexene in the aqueous phase was assumed to be constant, and a simplified nonlinear relationship between the aqueous and organic phase concentrations of cyclohexanol was utilized. Although the kinetic parameters obtained by them can describe the direct hydration process well, the simplified relation for the liquid− liquid equilibrium data is inconvenient to use in Aspen Plus. Therefore, the UNIFAC model with the group parameters from Skjold-Jørgensen et al.9 was applied to establish the relation between the aqueous and organic phase concentrations of cyclohexanol and further applied to recalculate the kinetic parameters in eq 1 with the same experimental kinetics data of Zhang et al. Finally, k = 6.4553 × 107 exp(−9259.1/T) kmol/ (kg cat·s) and Ka = 6.1101 × 10−4 exp(3524.8/T) were obtained by the new approach (refer to Supporting Information Figure R1). It is well-known that catalyst with fine particles and large specific surface area contributes to better industrial performance.6,12 This is why the multiphase slurry reactor has been used in the Asahi process. For modeling the multiphase slurry reactor, a sedimentation−diffusion model is utilized to consider the uneven distribution of catalyst in axial direction of reactor:10,11 ∂ccat ∂ 2c ∂c = Dcat cat + ucat cat 2 ∂t ∂z ∂z

TAC =

(2)

(3)

where l is the dimensionless length of the reactor, z/Lr, and Bocat is the dimensionless Bodenstein number of catalyst slurry, ucatLr/Dcat. Furthermore, two assumptions are made to neglect nonideal factors of slurry reactors: (1) no axial backmixing of the liquid phase and (2) even distribution of species in radial direction. Moreover, the kinetic equations have contained the influence of mass transfer resistance between liquid−liquid phases, so the mathematic model for the slurry reactor can be simplified as Finlet,ene dXene dl

Bocat mcat exp( −Bocat l) [1 − exp( −Bocat l)]

=

aq (1 − Bocat l)rol(T , xene , xolaq)

xaq ene

3. SIMULATION AND EVALUATION 3.1. Traditional Reaction−Separation Process. The traditional reaction−separation or reaction−distillation (R+D) process for direct cyclohexene hydration is schematically shown in Figure 1. The reaction temperature is 393.15 K from the trade-off between the reaction rate and the reaction equilibrium. Because the catalyst slurry that may flow into the column can run the risk of blocking a packing column, a plate column is chosen for the traditional process. Conventionally, the conversion per pass of cyclohexene in reactor is an important parameter. Improvement of the conversion can decrease the energy duty of the separation unit due to the less recycle of reactants. This could be achieved with more catalyst in the reactor, leading to higher cost of the reactor and catalyst. Figure 2 shows the effect of the conversion on TAC for the traditional process. For the case using one reactor, the optimal conversion is 0.11 with a minimum TAC of $4.02 × 106/a. However, when limiting the reactor diameter to less than 2 m, two reactors are necessary for loading large amount of catalyst with the minimum TAC increasing to $4.12 × 106/a under a slightly lower conversion of 0.10.

(4)

xaq ol

In eq 4, and are the aqueous phase concentrations of cyclohexene and cyclohexanol, respectively, which can be calculated on the basis of the phase equilibrium and the averaged composition of the liquid mixture under the conversion Xene. However, it is difficult to acquire accurately the Bodenstein number that is affected by the size of catalyst, structure of reactor, fluid velocity, etc. When the particle size of catalyst is relatively small, for example, the order of 10 μm in common slurry reactors, the Bodenstein number could be small enough by optimizing the structure and operation conditions of reactor. If the number is less than 10−2, the uneven distribution of catalyst slurry in axial direction could be neglected for the hydration process, and eq 5 can be used for calculation. Finlet,ene dXene dl

aq = mcat rol(T , xene , xolaq)

(6)

where the capital cost covers the costs of column, tray or packings, reactor, and heat exchanger; the operating cost includes the costs of steam, cooling water, and catalyst; 3 payback years are used. The heat transfer coefficient is fixed at 560 and 790 W/m2·K for condenser and reboiler, respectively. A height/diameter ratio of 10 is specified to determine the detailed size of the reactor with its volume; the reactor diameter is limited to 2 m to eliminate the uneven radial distribution of species as much as possible (but this limitation may lead to higher capacity cost of reactor since more parallel reactors are required). Moreover, an extra of 20% greater than the calculated is accounted for the final reactor cost since accessories should be included. More details can be found in the work of Tang et al.13 HETP = 0.35 m for packing columns in this work is fixed (a moderate value) with a price of $3,000/m3 for packings. This price is slightly higher than that of market since no column internals, e.g., liquid distributor, are considered. Two types of steams are used as the heat resources: one at 1.2 MPa costs $0.006/kg; the other at 4.0 MP has a 25% increase in price. The cost of cooling water is $5 × 10−6/kg. Moreover, the catalyst costs $7/kg with a lifetime of 3 months. The recovery of the fine catalyst from the slurry reactor pays an extra of 10% increase to the catalyst cost (yet, this should not be considered in the RD process). These prices are based on the work of Tang et al.13

The analytic solution of eq 2 in steady state is ccat = ccat,inlet exp( −Bocat l)

capital cost + operating cost payback year

(5)

With eq 5, the total catalyst loading mcat to achieve a certain cyclohexene conversion can be calculated, then the volume of 1462

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under the determined column construction and operation pressure can be carried out by the six steps as follow: (1) Determine the minimum reflux ratio using the Underwood method, and then the initial reflux ratio (20% greater than the minimum reflux ratio). (2) Estimate the column diameter dc from the maximum allowable vapor capacity factor of packing and determine the catalyst mass mcat,n on the nth reactive tray from the estimated dc, the density of catalyst (mass of catalyst in per volume of column) and the HETP. (3) Simulate the reactive column and obtain the concentration of product xB at bottom. (4) Increase the reflux ratio by 10% and repeat steps 2 and 3. (5) Determine the new reflux ratio according to an iterative equation of reflux ratio. (6) Repeat steps 2, 3, and 5 until the aimed purity of product at bottom is achieved. where, the third step can be achieved either by the equilibrium (EQ) or by the nonequilibrium (NEQ) model in Aspen Plus. As concluded, the difference between the results of the two models can be neglected for processes with relatively slow reactions, for example, the production of TAME, but not for the production of MTBE.16,17 It is because the slow reaction rate affects more obviously on the formation rate of product relative to the mass transfer insistence between the vapor and liquid phases. So the EQ model is chosen in this work to simulate the RD process with a slow reaction. Obviously, Aspen Plus can not automatically complete the whole iteration, suggesting that the application of Aspen Plus in design be limited although it involves with abundant databases and various thermodynamic methods. So embedding the iterative algorithm into the software is anticipated to bring convenience to the steady-state design of RD process. As an important reference, Vadapalli and Seader18 embedded an iterative algorithm for bifurcation analysis into Aspen Plus via the inline FORTRAN option available in the software. Figure 3a shows the flowchart of a relatively simple iterative algorithm for the design of RD column, where “Current sim” and “Next sim” are implemented by the RadFrac module in Aspen Plus. The reflux ratio iteration in the fifth step is carried out by the module F-1 with

Figure 1. Flowsheet of the traditional reaction−separation process.

Figure 2. Effect of the conversion per pass in reactor on TAC of the traditional process.

The evaluated results for the traditional reaction−separation process are listed in Table 1 (marked with “R+D”), where the cost of the decanter is not included. As indicated,12 the residence time of the decanter can be several seconds, suggesting that its contribution to the TAC is less than $0.01 × 106/a. From Table 1, it can be found that the energy cost shares the maximum in the TAC. The high energy cost is attributed to the high circulating flux of the unreacted cyclohexene caused by the low conversion per pass. 3.2. RD Process. Previously, Ye et al.14 analyzed in detail the kinetic−thermodynamic feasibility of the RD process for direct cyclohexene hydration. The study showed that higher pressure is favorable to the process; therefore, in this evaluation 0.5 MPa is selected after considering the cost increase with pressure. On the other hand, the study also showed that the reactive section should locate at the upper part of the column, and an unreactive section is necessary above the reactive section to separate the entrained cyclohexanol. With the help of the information from the feasibility analysis, the RD process will be further designed by an iterative algorithm. 3.2.1. Iterative Algorithm Embedded in Aspen Plus. In the literature, an iterative algorithm has been widely used for the steady-state design of RD column due to the coupling of reflux ratio and catalyst loading. On the basis of the iterative steps proposed by Subawalla and Fair,15 the design of the RD column

⎛ x − x1 ⎞ RR 2 = ⎜1 + ΔR B ⎟RR RR − RR1 ⎠ ⎝

(7)

where, ΔR is the set iteration step size of reflux ratio; RR1, RR, and RR2 are the reflux ratios used in the previous, current, and next simulation, respectively; x1 and xB are the previous and current cyclohexanol contents at the bottom, respectively. In implementation of the iterative algorithm, the values of RR and xB are assigned to RR1 and x1, respectively, before starting the “Current sim” module. After this, the value of RR2 from the module F-1 is assigned to the reflux ratio RR of the RD column. Finally, simulation for the RD column is carried out to obtain a new xB. It is worth noting that the module F-3 in Figure 3a is used to calculate the bulk density ρcat,col of catalyst in column with ρcat,col = ψcat(1 − εcat)ρcat,p and the mass of catalyst at the nth tray with mcat,n = πdc2ρcat,col × HETP/4. Obviously, these important variables depend on the catalyst volume fraction ψcat of catalytic packings, which significantly changes with the diameter of packings. Hoffmann et al.19 reported the relationship between the catalyst volume fraction and the diameter of a 1463

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Table 1. Comparison of the Information of Equipment and Energy Duty of the Three Processes items column

reboiler

condenser intermediate heat exchanger

reactor

cyclohexene conversion per pass capacity cost, $106/a

operating cost, $106/a

pressure, MPa reboiler temperature, K condenser temperature, K reboiler duty, MW condenser duty, MW diameter, m reflux ratio tray number (including condenser) reactive tray catalyst volume fraction total catalyst loading, kg effectiveness of catalyst steam flow rate, kg/s steam pressure, MPa area, m2 cooling water flow rate, kg/s area, m2 duty, MW steam flow rate, kg/s steam pressure, MPa area, m2 temperature, K number total catalyst loading, kg volume, m3 column reactor heat exchanger energy catalyst

TAC, $106/a

R+D

RD

RD-eff

SRC

0.1 432.89 343.93 17.69 −18.89 3.8 1.0 20

0.5 498.80 395.17 17.32 −17.13 3.7 7.2 25 2−21 0.545 32780 1.00 9.77 4.0 856.54 117.08 391.05

0.5 498.35 395.16 29.17 −28.93 5.1 12.9 25 2−21 0.550 62880 0.350 16.45 4.0 1442.5 197.76 660.5

0.1 432.81 343.72 11.68 −14.95 3.4 7.0 20

0.36 0.26

0.36 0.41

0.55 2.17 0.92 3.90

0.78 3.65 1.76 6.60

8.88 1.2 736.65 282.52 907.01

393.15 2 34930 65.41 0.10 0.35 0.44 0.65 1.60 1.08 4.12

5.87 1.2 481.68 223.51 717.58 2.63 1.32 1.2 36.21 393.15 2 34200 65.16 0.36 0.30 0.43 0.57 1.30 1.05 3.65

As another important influence factor, the effectiveness factor of catalyst in column also needs to be considered in the iterative algorithm. Although fine-grained catalyst slurry can provide high effectiveness, they are difficult to locate appropriately in the column and may block the column. Generally, the size of catalyst particle in commercial catalytic packings is in the range of 0.5−1.0 mm. In this situation, the effectiveness factor of catalyst could be lower than 1 even for slow reactions (e.g., the direct hydration of cyclohexene) due to the low effective diffusivity of liquid species inside catalyst particles. It suggests that the influence factor can not be neglected in the steady-state design of the RD process. Sundmacher and Hoffmann20 used an effectiveness factor in a pseudohomogeneous reaction rate expression to consider the influences of internal and external diffusion resistance. Although the “dusty fluid model” proposed by Higler et al.17 is more rigorous, some parameters in the model are difficult to determine for the reactive system of cyclohexanol. So in this work, a method similar to that of Sundmacher and Hoffmann20 is utilized. Accordingly, the internal and overall effectiveness factors of catalyst on the nth reactive tray are calculated by eqs 8 and 9, respectively. Furthermore, the effect of the catalyst effectiveness on process can be reflected by the effective catalyst mass described by eq 10 since there is only a reaction.

Figure 3. Iterative algorithm to steady-state design of RD column: (a) without and (b) with the consideration of catalyst effectiveness.

specific kind of catalytic packing. For example, the catalyst volume fraction is 0.398 for the catalytic packing MULTIPAKII with diameter of 0.1 m, while it is 0.550 for the diameter of 4.0 m. The column diameter dc required in the module F-3 is calculated in the module F-2 with the hydraulic characteristics of the packing, the mass flow rate FMV,n and the vapor density ρV,n from the “Current sim” module.

⎛ dc ⎞ ηin, n = 3De,1⎜ 1 ⎟ /[rol(T , xene , xol)r = R ρcat,p R ] ⎝ dr ⎠ r = R 1464

(8)

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Industrial & Engineering Chemistry Research ⎛ dc ⎞ ηall, n = 3De,1⎜ 1 ⎟ /[rol(Tn , xnaq,ene , xnaq,ol)ρcat,p R ] ⎝ dr ⎠ r = R mcat,eff, n = πηall, nρcat,col dc 2 HETP/4

Article

is 0.370, not 0.550 for MULTIPAK-II. As Figure 4 shows, the TAC increases much when the catalytic packing with a lower

(9) (10)

The concentration gradient in eqs 8 and 9 can be obtained by calculating the concentration and temperature profiles in catalyst particle, which are achieved by solving eqs 11 and 12 with their boundary conditions eqs 13 and 14. De , i d ⎛ 2 dci ⎞ ⎜r ⎟ + υiρcat, p rol = 0 r 2 dr ⎝ dr ⎠

(11)

λcat d ⎛ 2 dT ⎞ ⎜r ⎟ + ρ r ( −Hr ) = 0 cat, p ol r 2 dr ⎝ dr ⎠

(12)

dci dT = =0 dr dr

r = 0, De, i r = R, λcat

dci = kg, i(cnaq, i − ci) dr dT = hg (Tn − T ) dr

(13) Figure 4. Required reflux ratio and TAC of RD process when using different catalytic packing.

catalyst volume fraction is utilized. For example, when either MULTIPAK-I or the catalytic packing Bale23 is selected, the RD process shows no any superiority in light of the TAC compared to the traditional process. Figure 4 also shows that higher catalyst volume fraction reduces the required reflux ratio, suggesting that more catalyst can reduce the energy duty of separation in the RD process. This is reasonable because the formation rate of product in RD column depends on both of the catalyst loading and the removing rate of the product. When MULTIPAK-II is used as the catalytic packing, the diameter of catalyst particle in the packing is 0.85 mm and the effectiveness of catalyst is considered, the RD process is recalculated. Figure 5 shows the concentration profiles of

(14)

where the bulk concentration of component i corresponds to that in the aqueous phase since the hydrophilic zeolite is immersed in the phase according to Steyer et al.’s work.2 The mass and heat transfer coefficients in eq 14 are estimated by eq 15 with Re = 8FMLR/[πdc2(1 − εcat)μL].21 kg, i uL

Sci 2/3 =

hg c puL

Pr 2/3 =

0.765 0.365 + 0.335 Re 0.82 Re

(15)

By solving eqs 11 and 12, the effectiveness factor ηall,n could be obtained. The computation is carried out in the module F-4 as shown in Figure 3b. The calculated result of ηall,n is transferred to the module F-3 to calculate the effective catalyst loading mcat,eff,n by eq 10. The physical properties required in calculation are directly from Aspen Plus, while the effective diffusivity De inside catalyst particles is obtained by De = Dε/τ with D being calculated by the well-known Wilke-Chang equation22 and ε/τ = 0.08 being selected as the moderate value (ε is the porosity and τ is the tortuosity). 3.2.2. Performance of RD Process. According to the iterative algorithm shown in Figure 3a without considering the catalyst effectiveness, the performance of the RD process is shown in Table 1 (the “RD” column) when MULTIPAK-II with a high catalyst volume fraction is used as packing. Obviously, the TAC is lower than that of the traditional reaction−separation process. This decrease is not only due to the absence of the slurry reactor, but also the less catalyst loading (3.28 × 104 kg) compared to the traditional one (3.49 × 104 kg). On the other hand, the very close energy duties of the two processes or the circulating fluxes of the unreacted reactants mean the higher utilization of catalyst in the RD process, which is considered as the potential reason of the decline of TAC. Furthermore, the required reflux ratio and TAC of the RD process when using different type of catalytic packing are calculated. Note that the relationship between the catalyst volume fraction and the packing diameter is different for the different type of catalytic packing. For example, the volume fraction of catalyst in MULTIPAK-I with the diameter of 4.0 m

Figure 5. Concentration profiles of cyclohexene and cyclohexanol in catalyst particle.

cyclohexene and cyclohexanol in catalyst particles at the last reactive tray. From concentration profiles, the internal and overall effectiveness factors at each reactive tray can be calculated and are shown in Figure 6. The two independent points in Figure 5 represent the aqueous phase concentrations of cyclohexene and cyclohexanol on the last reactive tray. From Figure 5, it can be known that there is no large difference between the concentration at the catalyst surface and the bulk 1465

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necessary for the SRC having a low TAC. But based on several aspects of reasons, this low TAC is not possible in the SRC with that setup in fact. First, the large amount of catalyst brings the SRC to high catalyst and reactor costs. In addition, high reactant concentration and low product concentration near the distillation column top can help to achieve faster reaction rate; accordingly, except the side reactors near the column top, other side reactors show no remarkable performance but increase the catalyst cost since all the side reactors have the same catalyst loading. Moreover, the axial backmixing of liquid phase in the side reactors has been neglected in the calculation, suggesting the conversion is hardly limited by chemical equilibrium. So the SRC with a single side reactor may be optimal in theory. In addition, from the viewpoint of process operation and control, it is also obviously better to decrease the number of side reactors to one. Figure 7 shows the simplest SRC with a single side reactor approaching to the column top. In the SRC, the feed for the

Figure 6. Internal and overall effectiveness factors of catalyst at different reactive tray.

concentration, revealing that the resistance of the external diffusion is low and the catalytic process is controlled by the internal diffusion. This is also confirmed by the results in Figure 6: the average ratio of the overall effectiveness factor to the internal one is about 0.960. The final results with considering the effect of catalyst effectiveness are listed in Table 1 marked with “RD-eff”. The TAC is much higher than that of the traditional process, indicating that the RD process is not a reasonable choice. From the comparison between the cases with and without considering the catalyst effectiveness, the poor performance of RD process is found to be mainly contributed to the low catalyst effectiveness or large size of catalyst particle, not the RD technology itself or the slow hydration rate. However, for most RD processes with catalytic packings, complete elimination of the internal diffusion resistance of a catalyst is almost impossible. Therefore, the consideration of catalyst effectiveness is very important for determining if a more detailed process design is necessary, for example, the more rigorous calculation by the NEQ model. 3.2.3. Further Process Design. A case is studied in this section, where the catalyst volume fraction is assumed to 0.95 and the diffusion resistance of catalyst is neglected. The calculated results are shown as the two independent points (marked with “Assumed RD”) in Figure 4, which shows that the TAC can further decrease to $3.76 × 106/a though this decrease has been less dramatic due to higher catalyst cost (4.56 × 104 kg catalyst is handled). However, for a common reactive distillation column, the catalyst volume fraction of over 0.550 or the catalyst effectiveness factor being 1 is almost impossible. In order to break through the limitation of catalyst loading and achieve high catalyst effectiveness, a side reactor configuration (SRC) of the direct hydration process is proposed for the timely separation of product and to limit the fine-grained catalyst slurry in the appropriate location by side reactors. In addition, changing over catalyst is easier in the SRC. This kind of process configuration has been reported by previous works for intensification of other chemical processes, for example, the indirect hydration process of cyclohexene to cyclohexanol.24 3.3. Side Reactor Configuration. 3.3.1. Process Design. From the calculation for the RD process in section 3.2.3, 4.56 × 104 kg catalyst in 20 side reactors (not in parallel) seems to be

Figure 7. Flowsheet of the SRC with a single side reactor.

reactor is from the second tray of the column with the condenser as the first tray; the reactor discharge is fed into a decanter to obtain the organic and aqueous phases. Part of the aqueous phase is recycled into the reactor for maintaining the volume ratio 1.8:1 of aqueous to organic phase, and the other part is charged to the third tray of the column to meet with the organic phase from the decanter. Moreover, it is worth stating that the definition of reactor number in a SRC is based on the number of tray that has a side feed from reactor; so the process in Figure 7 is still named as the SRC with a single side reactor although there are two parallel side reactors. The process in Figure 7 is similar to the traditional reaction− separation process but still with obvious differences: there is almost the same equipment and the temperature difference between the distillation column and the slurry reactor is also allowable because the temperature of the side reactor can be adjusted at any expected value with the heat exchanger. However, in the SRC, the reactants are fed directly into the distillation column, not the reactor in the traditional process. Because of the allowable temperature difference, the column pressure can be at 0.1 MPa instead of 0.5 MPa in the RD process. The lower operating pressure can help to decrease the energy cost and use low grade steam as the heat source. 3.3.2. Process Evaluation. For the combination of a side reactor and the tray with a side feed from the side reactor, for example, that of the only side reactor and the third tray in 1466

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Figure 7, the component balance in steady state can be written as Vn + 1, iyn + 1, i + Ln − 1, ixn − 1, i − Vn , iyn , i − Ln , ixn , i + υi mcat , n =0

∫0

1

rol(Tr , n , xraq, n ,ene , xraq, n ,ol) dl (18)

According eq 18, different numbers of side reactors are examined to study whether the low TAC of the RD process in section 3.2.3 can be achieved by a SRC and further determine whether the SRC with a single side reactor is optimal in reality. For a better evaluation, the energy duty is fixed at 14.02 MW (the same to the RD process in section 3.2.3) in all of the calculations. According to calculation, the results of the TAC and the total catalyst loading in the different process are shown in Figure 8 (the RD process marked with “Assumed RD”).

Figure 9. Effect of the catalyst loading on TAC for the SRC with N = 1.

evaluation information for the optimal SRC is shown in Table 1 (marked with “SRC”). The information confirms that the TAC of the SRC finally decreases to $3.65 × 106/a and is 11.41% lower than that of the traditional reaction−separation process. Moreover, it is worthwhile to mention that the SRC with a single side reactor can be built easily on the basis of the traditional process due to the similarity between these two processes, suggesting very high feasibility of the SRC for direct hydration of cyclohexene.

4. DISCUSSION In order to systematically analyze the above different processes, Figure 10 shows the relationship between the total catalyst loading and the energy duty for different processes (the RD process has no consideration of the catalyst effectiveness).

Figure 8. Effect of number of side reactors in SRC on the required catalyst loading and TAC under the energy duty of 14.02 MW.

Figure 8 shows that more side reactors such as N = 4 lead to larger catalyst loading and higher TAC. Furthermore, as shown in Figure 8, the optimal number of side reactors is two, not one because the hydration rate of cyclohexene near the outlet of a single reactor with 3.78 × 104 kg catalyst is obviously limited by the chemical equilibrium. However, the SRC with a single side reactor is still ultimately recommended for the direct hydration of cyclohexene due to its simplicity and its only slightly higher TAC than the SRC with two side reactors. In addition, it is worth noting that the results presented in Figure 8 are not from the optimal SRC with a single side reactor. Obviously, the optimal catalyst loading should be less to avoid the limitation of chemical equilibrium and decrease the costs of catalyst and reactor as much as possible. Meanwhile, a higher reflux ratio should be necessary to compensate the decline of conversion caused by decreasing catalyst. For the trade-off, the effect of the catalyst loading on the TAC of the SRC with a single side reactor is studied and is shown in Figure 9. As seen, the optimal catalyst loading is 3.42 × 104 kg when N = 1 and less than that in Figure 8. Moreover, it needs to state that the abnormal change in Figure 9 corresponding to the catalyst loading of 3.78 × 104 kg is because the number of parallel side reactors increases to 3 from 2. The composition profiles of the liquid streams across side reactor and distillation column in this optimal SRC are respectively shown in Figure S1 and S2 in the Supporting Information. Moreover, the detailed

Figure 10. Relationship between the catalyst loading and the energy duty in different processes.

In the figure, the x-coordinate represents energy duty and further reflects the column cost, heat exchanger cost, cooling water, and steam cost in the TAC, while the y-coordinate reflects the rest of the parts of the TAC, the catalyst cost, and the reactor cost (if any). So as a rule, the minimum process TAC always corresponds to the point where both of the values of x- and y-coordinate are relatively less. From the rule, it is 1467

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easy to be known that the SRC with a single side reactor can have the lower TAC than other processes since the representing line is bellow other lines. Therefore, Figure 10 can be helpful to avoid the tedious calculation of TAC and directly choose the appropriate process with low TAC. Moreover, the rule can also offer help to determine the optimal point in a process as shown with a star in Figure 10. From the rule, it is undoubted that the point with a shorter distance from that origin can have a low TAC. For the traditional process and the SRC, the correctness of the rule is easy to be found. However, the rule seems unreasonable for the RD process. This is mainly because there is no reactor used and the value of the y-coordinate only representing the catalyst cost is not significant to the TAC. Moreover, the information in Figure 10 further indicates that the optimal SRC should possess fewer reactors that locate at the column top. However, the conclusion on the optimal number of side reactors is based on the premise that the axial backmixing of liquid phase in slurry reactor is neglected. For the infinite backmixing, the reactant conversion in a large reactor could be more easily limited by the chemical equilibrium. In this case, the catalyst in the optimal SRC tends to be assigned into more side reactors (like a reactor cascade), and the evaluation without the consideration of the declining operability shows that the optimal number of side reactors is 4.

λ = thermal conductive, W/m·K υ = stoichiometric coefficient ε = void fraction ρ = density, kg/m3 η = effectiveness factor μ = viscosity, Pa·s Supscripts

aq = aqueous phase Subscripts

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ASSOCIATED CONTENT

S Supporting Information *

cat = catalyst col = column ene = cyclohexene eff = effective i = component n = tray number ol = cyclohexanol p = particle r = reactor V, L = vapor and liquid, respectively

REFERENCES

(1) Musser, T. M. Cyclohexanol and cyclohexanone. In Industrial Organic Chemicals: Starting Materials and Intermediates; an Ullmann’s Encyclopedia; Wiley-VCH: Weinheim, 2003; Vol. 10, pp 279−290. (2) Steyer, F.; Freund, H.; Sundmacher, K. A Novel Reactive Distillation Process for the Indirect Hydration of Cyclohexene to Cyclohexanol Using a Reactive Entrainer. Ind. Eng. Chem. Res. 2008, 47, 9581−9587. (3) Steyer, F.; Qi, Z.; Sundemacher, K. Synthesis of cyclohexanol by three-phase reactive distillation influence of kinetics on phase equilibria. Chem. Eng. Sci. 2002, 57, 1511−1520. (4) Qi, Z.; Kolah, A.; Sundmacher, K. Residue curve maps for reactive distillation systems with liquid-phase splitting. Chem. Eng. Sci. 2002, 57, 163−178. (5) Steyer, F.; Sundmacher, K. Cyclohexanol Production via Esterification of Cyclohexene with Formic Acid and Subsequent

Composition profiles of the liquid streams across side reactors and distillation column, shown in Figures S1 and S2. They are from the optimal cases with one side reactor and with two side reactors. This material is available free of charge via the Internet at http://pubs.acs.org.

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NOTATION Bo = Bodenstein number c = concentration, kmol/m3 dc = diameter of column, m Dcat = diffusion coefficient of catalyst in reactor, m2/s De = effective diffusion coefficient in catalyst, m2/s Finlet = mole flow rate at reactor inlet, kmol/s FM = mass flow rate, kg/s hg = heat transfer coefficient, W/m2·K k = forward reaction rate constant, kmol/kg cat·s kg = mass transfer coefficient, m/s Ka = chemical equilibrium constant l = dimensionless length of reactor Lr = length of reactor, m mcat = catalyst loading amount, kg N = number of side reactors rol = formation rate of cyclohexanol, kmol/kg cat·s R = radius of catalyst, m Re, Pr, Sc = Reynolds number, Prandtl number, and Schmidt number RR = reflux ratio T = temperature, K V, L = vapor and liquid mole flow rate, respectively, kmol/s X = conversion y, x = mole fraction of vapor and liquid phase, respectively u = velocity, m/s

Greek Letters

5. CONCLUSIONS The RD process for direct hydration of cyclohexene to cyclohexanol is studied and compared with the traditional reaction−separation process. An iterative algorithm, which considers the effects of catalyst volume fraction and catalyst effectiveness, is embedded into Aspen Plus via inline FORTRAN to conveniently design more reliable RD process. Evaluation indicates that the RD process performs poorly with a relatively large TAC due to the low catalyst effectiveness, but not to the limited catalyst loading or the slow hydration rate. Accordingly, a side reactor configuration (SRC) is proposed for the direct hydration. It is found by optimization that the catalyst tends to be assigned into fewer reactors approaching to the column top for the optimal SRC. The subsequent evaluation further reveals that the SRC with a single reactor as the best choice for the direct hydration process has very high feasibility and obvious superiority in TAC compared with the traditional reaction−separation process. Finally, the TAC of the SRC can decrease by 11.41% of the traditional process.

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Article

AUTHOR INFORMATION

Corresponding Author

*Tel.: +86 13606052265. Fax: +86 592 2183054. E-mail: [email protected]. Notes

The authors declare no competing financial interest. 1468

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