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Feed-splitting operating strategy of a reactive distillation column for energy-saving production of butyl propionate
Journal of the Taiwan Institute of Chemical Engineers 41 (2010) 403–413
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Feed-splitting operating strategy of a reactive distillation column for energy-saving production of butyl propionate Hao-Yeh Lee a, Cheng-Hsun Jan a, I-Lung Chien b, Hsiao-Ping Huang a,* a b
Department of Chemical Engineering, National Taiwan University, Taipei 10617, Taiwan Department of Chemical Engineering, National Taiwan University of Science and Technology, Taipei 10672, Taiwan
A R T I C L E I N F O
A B S T R A C T
Article history: Received 31 December 2009 Received in revised form 19 February 2010 Accepted 2 March 2010
In the study of designing a reactive distillation process for producing heavy esters, it is found that splitting one of the rate-determining feed to the different location of a reactive distillation column can lead to a significant saving in the total annual cost and a reduction of the operating cost. Using esterification reaction to produce butyl propionate as an illustration, the saving of total annual cost can be up to 10.4% and the reduction of operating cost can be up to 11.9% in comparison with the design without feed-splitting. For the feasibility of the feed split, overall control strategy of this design alternation has also been studied. It is found that both design and control are feasible in the case of this feed-splitting operating strategy. This result is useful in prior to undertake optimal design of process and control using the mathematical programming approaches. ß 2010 Taiwan Institute of Chemical Engineers. Published by Elsevier B.V. All rights reserved.
Keywords: Butyl propionate Reactive distillation Feed-splitting Design and control
1. Introduction Reactive distillation received much attention in recent year due to its potential for capital productivity improvements, selectivity improvements, reduced energy use, and the reduction or elimination of solvents in the process (Malone and Doherty, 2000). In a book presenting the status and future directions of reactive disitllation (Sundmacher and Kienle, 2003), over one hundred industrially or potentially important reactions for reactive distillation applications are given. This illustrates the importance of this technology in industrial applications. Butyl proprionate (BuOPr), replacing light solvents such as methyl acetate or ethyl acetate, has increasingly been used as a cleaning agent for processing polymer because of its relatively low volatility. This important less-volatile solvent can be produced directly from n-butanol (BuOH) and propionic acid (HOPr) via esterification reaction with the presence of diluted sulfuric acid or cation exchange resin as catalyst. The design or control of the reactive distillation system producing butyl propionate is relatively scarce. The only paper in the literature is by Huang et al. (2004) studying the temperature control of such reactive distillation system. In that paper, no splitting of the feed stream was considered. In this paper, possible distribution of the feeds through feedsplitting will be explored to see if any benefit can be procured.
* Corresponding author. Tel.: +886 2 2363 8999; fax: +886 2 2362 3935. E-mail address: [email protected] (H.-P. Huang).
Distribution of feed in a tubular reactor was commonly used in the open literature to increase the yield of major product. However, no paper discusses feed-splitting operating strategy in a reactive distillation column system. The only exception is in Chapter 10 of the book by Luyben (2002) studying an ethylene glycol reactive distillation column. Splitting of ethylene oxide fresh feed into four equal streams feeding into different trays was studied to retard the parallel undesirable reaction. However, the impact on yield of distributing the feed is small. The organization of this paper is as follows. In Section 2, kinetic and thermodynamic models of the four-component system for the production of butyl proprionate will be given. The optimized design flowsheets with or without feed-splitting will be given in Section 3 for comparison purpose. The benefit of feed-splitting is also confirmed from the study of other heavy ester processes such as butyl acetate and amyl acetate. The overall control strategy of the optimized design with feed-splitting for the butyl proprionate process will be given in Section 4 with some concluding remarks given in Section 5. 2. Kinetic and thermodynamic models 2.1. Kinetic model For the kinetics of this reaction, Sharma et al. (1973) presented results using gel-type Dowex 50W-X8 as solid catalyst. Liu and Tan (2001) compared results from several cation exchange resins, and Lee et al. (2002) presented kinetic model using Amberlyst 35 as
1876-1070/$ – see front matter ß 2010 Taiwan Institute of Chemical Engineers. Published by Elsevier B.V. All rights reserved. doi:10.1016/j.jtice.2010.03.003
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2.2. Thermodynamic model
Nomenclature AmOAc BuOAc BuOH BuOPr HOPr k1 k1 Keq NR NRXN NS r rA TAC
amyl acetate butyl acetate n-butanol butyl proprionate propionic acid forward reaction rate constant reverse reaction rate constant equilibrium constant rectifying section reactive section stripping section reaction rate (kmol/s) reaction rate per unit catalyst weight [mol/(min kg cat)] total annual cost
catalyst. This slightly endothermic reversible reaction can be seen as below: C4 H9 OH þ C2 H5 COOH $ C2 H5 COOC4 H5 þ H2 O ðBuOHÞ
ðHOPrÞ
ðBuOPrÞ
We use the same kinetic model as used in Huang et al. (2004) in our study which is a quasi-homogeneous model with nonidealsolution assumption as appeared in Lee et al. (2002). The kinetic equation is: r A ¼ k1 aHOPr aBuOH k1 aBuOPr aH2 O
(1)
where rA is the reaction rate per unit catalyst weight [mol/ (min kg cat)] and a stands for the activity of corresponding components. In the above equation, k1 is the forward reaction rate constant [mol/(min kg cat)] as: k1 ¼ 1:6786 1010 ef7954=½TðKÞg with T in Kelvin. The rate constant for the reverse reaction, k1, is: k1 ¼ 3:1085 108 ef7135=½TðKÞg In the reactive distillation simulation below, the density of the catalyst is assumed to be 800 kg/m3 and also assuming that the catalyst occupy 50% of the tray liquid holdup.
The experimental vapor–liquid equilibrium (VLE) data of the most binary pairs of this system were given by Gmehling et al. (1977). For the propionic acid–butyl propionate pair and the nbutanol–butyl propionate pair, their experimental data can be found from the papers by Liu and Tan (2001), and Gonzales and Ortega (1996), respectively. For the experimental liquid–liquid equilibrium (LLE) data, the n-butanol–water pair and the butyl propionate–water pair can be found from the book by Sørensen and Arlt (1979) and ternary data of n-butanol–butyl propionate– water can be found from the paper by Lee et al. (2004). NRTL model parameters were obtained to fit the above equilibrium data. Because carboxylic acids have been reported to associate in the vapor phase, the fugacity coefficients in the vapor phase were thus calculated by employing the Hayden and O’Connell (1975) model. The Aspen Plus1 built-in association parameters are used in the calculations. The complete set of NRTL model parameters used in this paper are given in Table 1. For model validation, the computed VLE x–y and T–x–y curves of the complete system and the computed LLE curve of the n-butanol–butyl propionate–water ternary system are compared with experimental data in Figs. 1 and 2, respectively. From these two figures, the predicted VLE and LLE fit the experimental data pretty well. The predicted azeotropic compositions and temperatures of the system in comparison with the experimental data from Gmehling et al. (2004) are also given in Table 2. Based on these predictions, the temperatures of the normal boiling point of the pure components can be sorted as follows: H2 O ð100:02 CÞ
<
BuOH
ð117:68 CÞ
<
HOPr
ð141:14 CÞ
<
BuOPr
ð146:8 CÞ
In comparison with the above boiling points and the azeotropic temperatures in Table 2, the main product BuOPr has the maximum temperature and the ternary azeotrope of nbutanol–butyl propionate–water at 92.16 8C has the minimum temperature. A residue curve maps of quaternary mixtures that put four ternary residue curve maps together into a composite diagram are shown in Fig. 3. It is easily seen that the ternary azeotrope of n-butanol–butyl propionate–water is inside the LLE region that has two liquid phases. By making use of this LLE, very pure aqueous water can be obtained by a decanter. The organic phase in the decanter can be recycled back to the reactive distillation column. Because of having the highest boiling point,
Table 1 NRTL model parameters of this system. Comp,i
HOPr
HOPr
HOPr
BuOH
BuOH
Comp,j
BuOH
BuOPr
H2O
BuOPr
H2O
H2O
0 0 224.635 1104.285 0.3
0 0 384.9826 157.666 0.3
0.45921 4.378206 805.559 1064.285 0.1
0.47935 11.93514 638.1161 1415.35 0.203312
0 0 0 0 276.983 157.845 254.9331 403.3135 0.3 0.3 P P x j t ji G ji P xG x t G m m mj mj t i j P Aspen Plus NRTL: ln g i ¼ Pj x G þ j P jx i Gj x G aij aji bij bji cij
k k ki
k k kj
k k kj
BuOPr
where : Gi j ¼ expðai j t i j Þ; t i j ¼ ai j þ ðbi j =TÞ; ai j ¼ ci j ; t ii ¼ 0; Gii ¼ 1.
Table 2 Experimental and predicted values for azeotropes at pressure of 1 atm. Components
Experimental composition (mol%)
Experimental temperature (8C)
Computed composition (mol%)
Computed temperature (8C)
HOPr–H2O BuOH–H2O BuOPr–H2O BuOH–BuOPr–H2O
(4.81, 95.19) (24.76, 75.24) (16.62, 83.38) –
99.1 92.7 94.8 –
(4.14, 95.86) (24.70, 75.30) (17.19, 82.81) (20.23, 5.52, 74.25)
99.88 92.58 94.88 92.16
H.-Y. Lee et al. / Journal of the Taiwan Institute of Chemical Engineers 41 (2010) 403–413
Fig. 1. Binary VLE x–y and T–x–y diagrams at 1 atm for (a) HOPr-BuOH, (b) HOPr-BuOPr, (c) HOPr-H2O, (d) BuOH-BuOPr, (e) BuOH-H2O, and (f) BuOPr-H2O.
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Fig. 1. (Continued ).
3. Process design 3.1. Design flowsheet with no feed-splitting
Fig. 2. LLE experiment and predicted diagram for the ternary system of n-butanol– butyl propionate–water.
BuOPr will be rich in the bottom once distilled. The conceptual design of the overall system thus includes a reactive distillation column and a decanter as shown in Fig. 4, where it shows that the bottom product is BuOPr and the top product is aqueous water. In the study, pure HOPr and BuOH feed streams with both flow rates of 50 kmol/h are assumed.
Fig. 3. Composite RCMs for the four-component system.
The design variables in the system of Fig. 4 include: total number of stages of reactive section (NRXN), total number of stages of rectifying section (NR), total number of stages of stripping section (NS), the locations of the propionic acid feed and n-butanol feed. From Fig. 4, the propionic acid feed and n-butanol feed are placed at top and bottom stages of the reactive section, respectively. However, from the result of a study by Cheng and Yu (2005), the feed tray location is very important in the design of the reactive distillation column system; thus, in this study, the two feed tray locations can be varied. The reboiler duty of each simulation run is set to meet the product purity specifications: BuOPr product purity 99.5 wt% and propionic acid impurity in this product stream 100 ppm. A sequential iterative optimization search procedure is performed to minimize the Total Annual Cost (TAC) of the system. TAC includes: annualized capital cost, operating cost, and catalyst replacement cost. The annualized capital cost is calculated with the payback period of 3 years and using the cost data in Appendix E of Douglas (1988) book. The annualized capital cost includes: column shell, column trays, reboiler, and condenser. The operating cost includes steam and cooling water costs to operate reboiler and
Fig. 4. Proposed overall process design with no feed-splitting.
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Fig. 7. Summarized optimization plot at various NRXN. Fig. 5. Optimization plot with NRXN = 21.
condenser. The unit catalyst cost is assumed to be 7.7 $/kg and the catalyst life is assumed to be 3 months. The optimization procedure is as follows: (1) Set the feed tray locations of the propionic acid and n-butanol at the top and bottom stages of the reactive section, respectively. (2) Set a total number of stages of reactive section (NRXN). (3) Varying total number of stages of rectifying section (NR) and total number of stages of stripping section (NS) until TAC is minimized. (4) Go back to Step (2) to set another NRXN. (5) Go back to Step (1) to change the feed tray locations by moving propionic acid feed downward and n-butanol feed upward until TAC is minimized. An illustration plot for Step (3) can be seen in Fig. 5. In this figure, NRXN is set at 21 and the feed tray locations of the propionic acid and n-butanol are at the top and bottom stages of the reactive section, respectively. From this figure, one would observed that NR = 1 and NS = 6 are best in this situation. Another illustration for Step (5) can be seen in Fig. 6. From this figure, HOPr feed tray at 4th stage and BuOH feed tray at 6th stage are best in this situation. Notice that this altering of the feed tray locations represent significant reducing of TAC from the original
Fig. 6. Optimization plot with NR = 1, NRXN = 21, and NS = 6.
assumption of HOPr feed tray at the top of the reactive section (2nd stage) and BuOH feed tray at the bottom of the reactive section (22nd stage). The feed tray locations are counting from top to bottom with top tray to be the first stage and the reboiler to be the last stage. The above two illustration plots show that with NRXN = 21 the best combination of the other design variables are: NR = 1, NS = 6, HOPr feed tray at 4th stage, and BuOH feed tray at 6th stage. Collecting results from other optimized runs, Fig. 7 shows a summarized optimization plot at various NRXN. From this figure, the NRXN = 21 gave the best result with NR = 1, NS = 6, HOPr feed tray at 4th stage, and BuOH feed tray at 6th stage. The minimized TAC is $358,700 with annualized capital cost of $207,400, operating cost of $128,100, and catalyst cost of $23,200. The optimized design flowsheet is thus shown in Fig. 8. The liquid composition profile and reaction rate of the optimized process design can be seen in Fig. 9. The liquid compositions of the HOPr and BuOH are all at their highest values
Fig. 8. Optimized design flowsheet with no feed-splitting.
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Fig. 10. Summarized optimization plot at various split HOPr feed flow rate. Fig. 9. Liquid composition profile and reaction rate of the optimized design with no feed-splitting.
at their feed tray locations. The highest reaction rate is at 5th stage between the two feed tray locations. Notice that the BuOH composition maintained at high value from its feed tray location (6th tray) downward. It is speculated that if any benefit can be gained by splitting HOPr feed into two or more streams with partial HOPr feed entering into the lower part of the reactive section where BuOH composition are at its higher value. For illustration and process simplification purpose, we will only consider splitting the HOPr feed into two sub-feed streams in the following. 3.2. Design flowsheet with HOPr feed-splitting 3.2.1. Retain main HOPr and BuOH feed tray locations Assuming that we retain the main HOPr feed tray location at 4th stage and BuOH feed tray location at 6th stage as in the optimized design flowsheet in previous Fig. 8. Also, from all the previous optimization runs, the stages of rectifying section are all found to be NR = 1, thus this is also retained in the following optimization search. With the allowing of HOPr feed to split, there will be two other design variables to be determined: the split HOPr feed tray location and the split HOPr feed flow rate. The sequential iterative optimization procedure is as follows: (1) Set the split HOPr feed flow rate. (2) Set the split HOPr feed tray location. (3) Varying total number of stages of reactive section (NRXN) and total number of stages of stripping section (NS) until TAC is minimized. (4) Go back to Step (2) to set another split HOPr feed tray location to minimize TAC. (5) Go back to Step (1) to set another split HOPr feed flow rate to minimize TAC. Fig. 10 shows the summarized result of this optimization search at various split HOPr feed flow rate. The TAC is minimized at split feed flow rate of 10 kmol/h. At this design, split HOPr feed tray location is at 6th stage, NRXN = 24, and NS = 1. Notice that the total number of stages of the reactive distillation column (including reboiler) is reduced from 28 to 26. The TAC is reduced from $358,700 to $331,400 (a reduction of 7.6%) with annualized capital cost from $207,400 to $190,500 (a reduction of 8.1%) and operating cost from $128,100 to $116,700 (a reduction of 8.8%). However, because the total number of stages of reactive section is increased, the catalyst cost is increased from $23,200 to $24,200 (an increase of 4.3%).
3.2.2. Allow the main HOPr feed tray location to be varied With the main HOPr feed tray location as an additional design variable, the search for the minimized TAC becomes more timeconsuming. The results in Section 3.2.1 can be considered as a set of optimized results with the main HOPr feed tray location set at 4th stage. By varying the main HOPr feed tray location, other sets of optimized results can be obtained. Table 3 shows the summarized optimization results at various main HOPr feed tray locations. Notice that with the main HOPr feed tray locations set at 1st, 2nd, or 3rd stages, the flow rate at the main HOPr feed tray location is actually smaller than the split HOPr feed flow rate. Fig. 11 shows the summarized optimization results of TAC at various main HOPr feed tray locations. The TAC is minimized with the main HOPr feed tray location at 2nd stage. At this design, split HOPr feed tray location is at 6th stage with flow rate of 30 kmol/h, NRXN = 23, NS = 2, and NR = 1. The main portion of the HOPr feed is entering the column at the same location as the BuOH feed does. The remaining HOPr feed enters the column at the top of the reactive section. The TAC, annualized capital cost, and operating cost are all further reduced from the result in Section 3.2.1. Also notice that the catalyst cost is reduced from the base case in Section 3.1. The reason is that although the number of stages of the reactive section is increased from 21 to 23, the column diameter is also decreased from 1.117 m to 1.048 m, thus resulting in the reduction of catalyst cost from $23,200 to $22,400. With this simple process design alternation by considering feed-splitting, the TAC can be reduced from $358,700 to $321,400 (a reduction of 10.4%) with annualized capital cost from $207,400 to $186,200 (a reduction of 10.2%) and operating cost from $128,100 to $112,800 (a reduction of 11.9%). The optimized design flowsheet is summarized in Fig. 12. The comparison of the specific reaction rate for the original base case with no feed-splitting and the optimized results from Sections 3.2.1 and 3.2.2 can be seen in Fig. 13. Notice that without feedsplitting, the maximum value occurs at 5th stage (between the two feed tray locations) and decreases on both directions upward and downward of the reactive distillation column. However, in the two feed-splitting cases, the profile of specific reaction rate has two peaks when a portion of HOPr feed is moved downward. And according to the composition profiles of optimized design flowsheet with feed-splitting shown in Fig. 14, the HOPr composition is still higher below the BuOH feed stage in reactive zone than the case without feed-splitting shown in Fig. 9. It means that high concentration of both reactants would help the specific
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Table 3 Summarized optimization results at various main HOPr feed tray locations.
Total stages NR NRXN NS HOPr feed flow (kmol/h) Split HOPr feed flow (kmol/h) HOPr feed tray Split HOPr feed tray BuOH feed tray Column diameter (m) Reboiler duty (kW) TAC ($1000/yr) Capital cost ($1000/yr) Operating cost ($1000/yr) Catalyst cost ($1000/yr)
Base case
HOPr at 1st stage
HOPr at 2nd stage
HOPr at 3rd stage
HOPr at 4th stage
28 1 21 6 50 – 4 – 6 1.117 1131 358.7 207.4 128.1 23.2
27 1 25 1 15 35 1 6 6 1.052 1002 327.2 189.2 113.5 24.5
26 1 23 2 20 30 2 6 6 1.048 995.9 321.4 186.2 112.8 22.4
27 1 24 2 25 25 3 6 6 1.046 990.0 324.3 188.9 112.2 23.3
26 1 24 1 40 10 4 6 6 1.067 1030 331.4 190.5 116.7 24.2
Fig. 11. Summarized optimization plot at various main HOPr feed tray locations.
Fig. 12. Optimized design flowsheet with feed-splitting.
Fig. 13. Comparison of the specific reaction rate for three design cases.
Fig. 14. Liquid composition profile and reaction rate of the optimized design with feed-splitting.
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Table 4 Summary of kinetic models, k1, and Keq used in the study. System
Kinetic model (catalyst)
k1 (T = 363 K)
BuOPr
Quasi-homogeneous model (Amberlyst 35) r ¼ mcat ðk1 aHOPr aBuOH k1 aBuOPr aH2 O Þ 1 1 7954 1:6786 1010 exp k1 ¼ 1000 60 T 1 1 7135 3:1085 108 exp k1 ¼ 1000 60 T
8.52 105 [kmol/(kgcat s)]
Keq (T = 363 K)
5.66
Pseudo-homogeneous model (Amberlyst 15) r ¼ mcat ðk1 aHAc aBuOH k1 aBuOAc aH2 O Þ BuOAc
70; 660 RT 74241:7 ¼ 1:0135 106 exp p RT
k1 ¼ 3:3856 106 exp k1
2.32 104 [kmol/(kgcat s)]
10.9
Quasi-homogeneous model (Amberlyst 15) r ¼ mcat ðk1 C HAc C AmOH k1 C AmOAc C H2 O Þ AmOAc
51; 740 RT 45; 280 ¼ 2:2533 exp RT
k1 ¼ 31:1667 exp k1
reaction rate increasing for this 2nd order reaction system. By calculating the summation of the specific reaction rate at each stage, it is found that the design flowsheet in Fig. 12 gave the largest value representing most efficient usage of the reactive volume. 3.3. Extension of this feed-splitting strategy to other heavy ester processes The design flowsheet for the production of butyl proprionate belongs to a general class of the reactive distillation flowsheet called Type III as in the paper by Tang et al. (2005). In that paper, the design flowsheets of butyl acetate (BuOAc) and amyl acetate (AmOAc) have been studied. We have also preliminary investigated the alternative feed-splitting strategy of these two heavy ester flowsheets. Same kinetic and thermodynamic models as in Tang et al. (2005) are used in this study and the design flowsheet of the two processes with no feed-splitting are established. These two base cases are compared with the results from feed-splitting operating strategy. For the feed-splitting design, the main portion of the acid feed is designed to enter the column at the same location as the alcohol feed does. The remaining acid feed is designed to enter the column at the top of the reactive section. This is the design experience from the result of Section 3.2. With no detailed optimization search, the above feed-splitting design has already seen the saving of TAC in comparison with the no feed-splitting design. For the case of butyl acetate, significant saving of TAC over 30% can be realized with this simple design alternation. For the case of amyl acetate, however, the TAC saving is much less (only 2.4%). The reason for the discrepancy in TAC savings on butyl proprionate, butyl acetate, and amyl acetate can be explained by their equilibrium rate constants. Table 4 shows the kinetic models and their model parameters used in the simulation. The table also listed their forward reaction rate constants and the equilibrium rate constants at 363 K. Notice that the equilibrium rate constant of the BuOAc system are the greatest and that of the AmOAc system are the smallest. Thus, the BuOAc system has the most potential for improving the reaction performance while the AmOAc system has little room for any further improvement.
1.13 106 [m6/(kmol kgcat s)]
1.6
4. Overall control strategy The inventory control loops of the design flowsheet in Fig. 12 will be determined first. The three level control loops are: bottom level controlled by the bottom product flow; organic level controlled by the organic reflux flow; and the aqueous level controlled by the aqueous product flow. Other control loops include the top column pressure controlled at 1.1 atm by manipulating the top vapor flow and the decanter temperature controlled at 40 8C by manipulating the condenser duty. There are also two ratio-control schemes implemented in the overall control strategy. One of it is to maintain the ratio of flow rates between the splitting HOPr feed (at 6th stage) and the HOPr feed (at 2nd stage) at 1.5. The other ratio scheme is to keep the ratio of flow rates between the total HOPr feed and the BuOH feed at 1.0. Notice that the above ratio scheme of keeping the ratio of flow rates between the total HOPr feed and the BuOH feed at 1.0 is not suitable when either of the feed compositions change. For example, if BuOH feed composition changes from pure to include some water in this feed stream, this ratio should be decreased to allow less HOPr feed from entering the column. This set value of the ratio should be adjusted by a product quality control loop. In this paper, we assumed that the on-line composition measurement is not readily available, thus, the product quality control loop will be replaced by some tray temperature control loop. Beside the ratio mentioned above can be adjusted by a tray temperature control loop, another temperature control loop is needed to hold the temperature profile inside the column by manipulating the reboiler duty. Therefore, a dual-temperature control strategy commonly used in reactive distillation column will be investigated in this system. The two temperature control points can be determined by open-loop sensitivity analysis. Figs. 15 and 16 show the open-loop sensitivity test by changing either of the ratio of the two feeds or the reboiler duty for 0.1% and plotting the tray temperature differences versus the base case condition. From these two figures, it is determined that the 10th stage temperature should be controlled by manipulating the reboiler duty and the 22nd stage temperature should be controlled by manipulating the ratio of the two feeds.
H.-Y. Lee et al. / Journal of the Taiwan Institute of Chemical Engineers 41 (2010) 403–413
Fig. 15. Open-loop sensitivity test with 0.1% changes in the feed ratio.
411
Fig. 16. Open-loop sensitivity test with 0.1% changes in the reboiler duty.
Fig. 17. Closed-loop simulation with 10% changes in the BuOH or HOPr feed composition.
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Fig. 18. Closed-loop simulation results with and without feedforward compensation for +10% throughput changes.
Fig. 17 shows 10% changes in the BuOH and HOPr feed compositions. In this two simulation runs, the feed compositions were changed to include 10% water in the original pure feed stream. It is noticed that the dual-temperature control strategy works well despite these feed composition changes. The purity of the BuOPr product stream was closely maintained (99.5 wt% BuOPr and 100 ppm HOPr impurity). For the throughput changes, the above dual-temperature control strategy was not able to hold the high-purity product specifications. However, throughput change is considered as a measurable disturbance, thus some feedforward compensation as in Huang et al. (2004) can be used to circumvent the above problem. The set-point of the 22nd stage temperature was determined as a linear function of the throughput (nominal case at 50 kmol/h BuOH feed flow rate). For a 10% BuOH feed flow rate change, the set-point of this temperature loop was adjusted
from 121.96 8C to 125.72 8C in order to maintain the high-purity product specifications. The set-point of the other temperature control loop (10th stage temperature) does not need to be changed. Fig. 18 shows such simulation runs with and without this feedforward compensation for 10% throughput changes. It is clearly shown that the HOPr impurity is reduced from 155 ppm to 93 ppm with the implementation of this feedforward compensation. The final proposed overall control strategy can be seen in Fig. 19. With HOPr feed being splitting into two sub-streams, an additional degree of freedom can be utilized by the overall control strategy. However, simply holding the two sub-stream flow rates at a constant ratio, the abovementioned overall control strategy can already properly handle feed composition and throughput disturbance. Thus, there is no need to develop other more complicated control strategy to use this extra degree of freedom.
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Fig. 19. Proposed overall control strategy.
5. Conclusions In this study, the design and control of a reactive distillation column system for the production of butyl propionate has been investigated. Significant TAC and operating costs can be reduced by properly locate the two feed streams inside the reactive section. By simply splitting the HOPr feed into two sub-streams and entering into the reactive distillation column at different appropriate locations can further save TAC by 10.4% and cut the required operating cost by 11.9%. This study has been extended to other heavy ester systems and saving of TAC has also been shown. The percentage of the saving can be qualitatively approximated by the value of the equilibrium rate constant. Bigger saving can be realized with this simple feed-splitting design if the value of the equilibrium rate constant is large. A simple overall control strategy for this process with dual-temperature control has also been proposed. The high-purity product specifications can be maintained in the face of feed compositions and throughput changes. References Cheng, Y. C. and C. C. Yu, ‘‘Effects of Feed Tray Locations to the Design of Reactive Distillation and its Implication to Control,’’ Chem. Eng. Sci., 60, 4661 (2005). Douglas, J. M., Conceptual Design of Chemical Processes, McGraw-Hill Inc., New York, USA (1988).
Gmehling, J., J. Menke, J. Krafczyk, and K. Fischer, Azeotropic Data,’’ 2nd ed., Wiley–VCH, Weinheim, Germany (2004). Gmehling, J., U. Onken, and W. Arlt, Vapor–Liquid Equilibrium Data Collelection, Chemistry Data Series,’’ vol. 1. DECHEMA, Frankfurt/Main, Germany (1977). Gonzales, E. and J. Ortega, ‘‘Densities and Isobaric Vapor–Liquid Equilibria for the Mixtures Formed by Four Butyl Esters and 1-Butanol,’’ J. Chem. Eng. Data, 41, 53 (1996). Hayden, J. G. and J. P. O’Connell, ‘‘A Generalized Method for Predicting Second Virial Coefficients,’’ Ind. Eng. Chem. Proc. Des. Dev., 14, 209 (1975). Huang, S. G., C. L. Kuo, S. B. Hung, Y. W. Chen, and C. C. Yu, ‘‘Temperature Control of Heterogeneous Reactive Distillation,’’ AIChE J., 50, 2203 (2004). Lee, M. J., J. Y. Chiu, and H. M. Lin, ‘‘Kinetics of Catalytic Esterification of Propionic Acid and n-Butanol over Amberlyst 35,’’ Ind. Eng. Chem. Res., 41, 2882 (2002). Lee, M. J., L. H. Tsai, G. B. Hong, and H. M. Lin, ‘‘Multiphase Equilibria for Binary and Ternary Mixtures Containing Propionic Acid, n-Butanol, Butyl Propionate, and Water,’’ Fluid Phase Equilibria, 216, 219 (2004). Liu, W. T. and C. S. Tan, ‘‘Liquid-Phase Esterification of Propionic Acid with n-Butanol,’’ Ind. Eng. Chem. Res., 40, 3281 (2001). Luyben, W. L., Plantwide Dynamic Simulators in Chemical Processing and Control, Marcel Dekker, New York, USA (2002). Malone, M. F. and M. F. Doherty, ‘‘Reactive Distillation,’’ Ind. Eng. Chem. Res., 39, 3953 (2000). Sharma, O. N., G. D. Nageshwar, and P. S. Mene, ‘‘Esterification of Butanol with Propionic Acid Catalysed by a Cation Exchange Resin,’’ Indian J. Technol., 11, 360 (1973). Sørensen, J. M. and W. Arlt, Liquid–Liquid Equilibrium Data Collelection, Chemistry Data Series,’’ vol. 5. DECHEMA, Frankfurt/Main, Germany (1979). Sundmacher, K. and A. Kienle, Reactive Distillation: Status and Future Directions, Wiley– VCH, Weinheim, Germany (2003). Tang, Y. T., Y. W. Chen, H. P. Huang, C. C. Yu, S. B. Hung, and M. J. Lee, ‘‘Design of Reactive Distillations for Acetic Acid Esterification,’’ AIChE J., 51, 1683 (2005).
Contents lists available at ScienceDirect
Journal of the Taiwan Institute of Chemical Engineers journal homepage: www.elsevier.com/locate/jtice
Feed-splitting operating strategy of a reactive distillation column for energy-saving production of butyl propionate Hao-Yeh Lee a, Cheng-Hsun Jan a, I-Lung Chien b, Hsiao-Ping Huang a,* a b
Department of Chemical Engineering, National Taiwan University, Taipei 10617, Taiwan Department of Chemical Engineering, National Taiwan University of Science and Technology, Taipei 10672, Taiwan
A R T I C L E I N F O
A B S T R A C T
Article history: Received 31 December 2009 Received in revised form 19 February 2010 Accepted 2 March 2010
In the study of designing a reactive distillation process for producing heavy esters, it is found that splitting one of the rate-determining feed to the different location of a reactive distillation column can lead to a significant saving in the total annual cost and a reduction of the operating cost. Using esterification reaction to produce butyl propionate as an illustration, the saving of total annual cost can be up to 10.4% and the reduction of operating cost can be up to 11.9% in comparison with the design without feed-splitting. For the feasibility of the feed split, overall control strategy of this design alternation has also been studied. It is found that both design and control are feasible in the case of this feed-splitting operating strategy. This result is useful in prior to undertake optimal design of process and control using the mathematical programming approaches. ß 2010 Taiwan Institute of Chemical Engineers. Published by Elsevier B.V. All rights reserved.
Keywords: Butyl propionate Reactive distillation Feed-splitting Design and control
1. Introduction Reactive distillation received much attention in recent year due to its potential for capital productivity improvements, selectivity improvements, reduced energy use, and the reduction or elimination of solvents in the process (Malone and Doherty, 2000). In a book presenting the status and future directions of reactive disitllation (Sundmacher and Kienle, 2003), over one hundred industrially or potentially important reactions for reactive distillation applications are given. This illustrates the importance of this technology in industrial applications. Butyl proprionate (BuOPr), replacing light solvents such as methyl acetate or ethyl acetate, has increasingly been used as a cleaning agent for processing polymer because of its relatively low volatility. This important less-volatile solvent can be produced directly from n-butanol (BuOH) and propionic acid (HOPr) via esterification reaction with the presence of diluted sulfuric acid or cation exchange resin as catalyst. The design or control of the reactive distillation system producing butyl propionate is relatively scarce. The only paper in the literature is by Huang et al. (2004) studying the temperature control of such reactive distillation system. In that paper, no splitting of the feed stream was considered. In this paper, possible distribution of the feeds through feedsplitting will be explored to see if any benefit can be procured.
* Corresponding author. Tel.: +886 2 2363 8999; fax: +886 2 2362 3935. E-mail address: [email protected] (H.-P. Huang).
Distribution of feed in a tubular reactor was commonly used in the open literature to increase the yield of major product. However, no paper discusses feed-splitting operating strategy in a reactive distillation column system. The only exception is in Chapter 10 of the book by Luyben (2002) studying an ethylene glycol reactive distillation column. Splitting of ethylene oxide fresh feed into four equal streams feeding into different trays was studied to retard the parallel undesirable reaction. However, the impact on yield of distributing the feed is small. The organization of this paper is as follows. In Section 2, kinetic and thermodynamic models of the four-component system for the production of butyl proprionate will be given. The optimized design flowsheets with or without feed-splitting will be given in Section 3 for comparison purpose. The benefit of feed-splitting is also confirmed from the study of other heavy ester processes such as butyl acetate and amyl acetate. The overall control strategy of the optimized design with feed-splitting for the butyl proprionate process will be given in Section 4 with some concluding remarks given in Section 5. 2. Kinetic and thermodynamic models 2.1. Kinetic model For the kinetics of this reaction, Sharma et al. (1973) presented results using gel-type Dowex 50W-X8 as solid catalyst. Liu and Tan (2001) compared results from several cation exchange resins, and Lee et al. (2002) presented kinetic model using Amberlyst 35 as
1876-1070/$ – see front matter ß 2010 Taiwan Institute of Chemical Engineers. Published by Elsevier B.V. All rights reserved. doi:10.1016/j.jtice.2010.03.003
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2.2. Thermodynamic model
Nomenclature AmOAc BuOAc BuOH BuOPr HOPr k1 k1 Keq NR NRXN NS r rA TAC
amyl acetate butyl acetate n-butanol butyl proprionate propionic acid forward reaction rate constant reverse reaction rate constant equilibrium constant rectifying section reactive section stripping section reaction rate (kmol/s) reaction rate per unit catalyst weight [mol/(min kg cat)] total annual cost
catalyst. This slightly endothermic reversible reaction can be seen as below: C4 H9 OH þ C2 H5 COOH $ C2 H5 COOC4 H5 þ H2 O ðBuOHÞ
ðHOPrÞ
ðBuOPrÞ
We use the same kinetic model as used in Huang et al. (2004) in our study which is a quasi-homogeneous model with nonidealsolution assumption as appeared in Lee et al. (2002). The kinetic equation is: r A ¼ k1 aHOPr aBuOH k1 aBuOPr aH2 O
(1)
where rA is the reaction rate per unit catalyst weight [mol/ (min kg cat)] and a stands for the activity of corresponding components. In the above equation, k1 is the forward reaction rate constant [mol/(min kg cat)] as: k1 ¼ 1:6786 1010 ef7954=½TðKÞg with T in Kelvin. The rate constant for the reverse reaction, k1, is: k1 ¼ 3:1085 108 ef7135=½TðKÞg In the reactive distillation simulation below, the density of the catalyst is assumed to be 800 kg/m3 and also assuming that the catalyst occupy 50% of the tray liquid holdup.
The experimental vapor–liquid equilibrium (VLE) data of the most binary pairs of this system were given by Gmehling et al. (1977). For the propionic acid–butyl propionate pair and the nbutanol–butyl propionate pair, their experimental data can be found from the papers by Liu and Tan (2001), and Gonzales and Ortega (1996), respectively. For the experimental liquid–liquid equilibrium (LLE) data, the n-butanol–water pair and the butyl propionate–water pair can be found from the book by Sørensen and Arlt (1979) and ternary data of n-butanol–butyl propionate– water can be found from the paper by Lee et al. (2004). NRTL model parameters were obtained to fit the above equilibrium data. Because carboxylic acids have been reported to associate in the vapor phase, the fugacity coefficients in the vapor phase were thus calculated by employing the Hayden and O’Connell (1975) model. The Aspen Plus1 built-in association parameters are used in the calculations. The complete set of NRTL model parameters used in this paper are given in Table 1. For model validation, the computed VLE x–y and T–x–y curves of the complete system and the computed LLE curve of the n-butanol–butyl propionate–water ternary system are compared with experimental data in Figs. 1 and 2, respectively. From these two figures, the predicted VLE and LLE fit the experimental data pretty well. The predicted azeotropic compositions and temperatures of the system in comparison with the experimental data from Gmehling et al. (2004) are also given in Table 2. Based on these predictions, the temperatures of the normal boiling point of the pure components can be sorted as follows: H2 O ð100:02 CÞ
<
BuOH
ð117:68 CÞ
<
HOPr
ð141:14 CÞ
<
BuOPr
ð146:8 CÞ
In comparison with the above boiling points and the azeotropic temperatures in Table 2, the main product BuOPr has the maximum temperature and the ternary azeotrope of nbutanol–butyl propionate–water at 92.16 8C has the minimum temperature. A residue curve maps of quaternary mixtures that put four ternary residue curve maps together into a composite diagram are shown in Fig. 3. It is easily seen that the ternary azeotrope of n-butanol–butyl propionate–water is inside the LLE region that has two liquid phases. By making use of this LLE, very pure aqueous water can be obtained by a decanter. The organic phase in the decanter can be recycled back to the reactive distillation column. Because of having the highest boiling point,
Table 1 NRTL model parameters of this system. Comp,i
HOPr
HOPr
HOPr
BuOH
BuOH
Comp,j
BuOH
BuOPr
H2O
BuOPr
H2O
H2O
0 0 224.635 1104.285 0.3
0 0 384.9826 157.666 0.3
0.45921 4.378206 805.559 1064.285 0.1
0.47935 11.93514 638.1161 1415.35 0.203312
0 0 0 0 276.983 157.845 254.9331 403.3135 0.3 0.3 P P x j t ji G ji P xG x t G m m mj mj t i j P Aspen Plus NRTL: ln g i ¼ Pj x G þ j P jx i Gj x G aij aji bij bji cij
k k ki
k k kj
k k kj
BuOPr
where : Gi j ¼ expðai j t i j Þ; t i j ¼ ai j þ ðbi j =TÞ; ai j ¼ ci j ; t ii ¼ 0; Gii ¼ 1.
Table 2 Experimental and predicted values for azeotropes at pressure of 1 atm. Components
Experimental composition (mol%)
Experimental temperature (8C)
Computed composition (mol%)
Computed temperature (8C)
HOPr–H2O BuOH–H2O BuOPr–H2O BuOH–BuOPr–H2O
(4.81, 95.19) (24.76, 75.24) (16.62, 83.38) –
99.1 92.7 94.8 –
(4.14, 95.86) (24.70, 75.30) (17.19, 82.81) (20.23, 5.52, 74.25)
99.88 92.58 94.88 92.16
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Fig. 1. Binary VLE x–y and T–x–y diagrams at 1 atm for (a) HOPr-BuOH, (b) HOPr-BuOPr, (c) HOPr-H2O, (d) BuOH-BuOPr, (e) BuOH-H2O, and (f) BuOPr-H2O.
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Fig. 1. (Continued ).
3. Process design 3.1. Design flowsheet with no feed-splitting
Fig. 2. LLE experiment and predicted diagram for the ternary system of n-butanol– butyl propionate–water.
BuOPr will be rich in the bottom once distilled. The conceptual design of the overall system thus includes a reactive distillation column and a decanter as shown in Fig. 4, where it shows that the bottom product is BuOPr and the top product is aqueous water. In the study, pure HOPr and BuOH feed streams with both flow rates of 50 kmol/h are assumed.
Fig. 3. Composite RCMs for the four-component system.
The design variables in the system of Fig. 4 include: total number of stages of reactive section (NRXN), total number of stages of rectifying section (NR), total number of stages of stripping section (NS), the locations of the propionic acid feed and n-butanol feed. From Fig. 4, the propionic acid feed and n-butanol feed are placed at top and bottom stages of the reactive section, respectively. However, from the result of a study by Cheng and Yu (2005), the feed tray location is very important in the design of the reactive distillation column system; thus, in this study, the two feed tray locations can be varied. The reboiler duty of each simulation run is set to meet the product purity specifications: BuOPr product purity 99.5 wt% and propionic acid impurity in this product stream 100 ppm. A sequential iterative optimization search procedure is performed to minimize the Total Annual Cost (TAC) of the system. TAC includes: annualized capital cost, operating cost, and catalyst replacement cost. The annualized capital cost is calculated with the payback period of 3 years and using the cost data in Appendix E of Douglas (1988) book. The annualized capital cost includes: column shell, column trays, reboiler, and condenser. The operating cost includes steam and cooling water costs to operate reboiler and
Fig. 4. Proposed overall process design with no feed-splitting.
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Fig. 7. Summarized optimization plot at various NRXN. Fig. 5. Optimization plot with NRXN = 21.
condenser. The unit catalyst cost is assumed to be 7.7 $/kg and the catalyst life is assumed to be 3 months. The optimization procedure is as follows: (1) Set the feed tray locations of the propionic acid and n-butanol at the top and bottom stages of the reactive section, respectively. (2) Set a total number of stages of reactive section (NRXN). (3) Varying total number of stages of rectifying section (NR) and total number of stages of stripping section (NS) until TAC is minimized. (4) Go back to Step (2) to set another NRXN. (5) Go back to Step (1) to change the feed tray locations by moving propionic acid feed downward and n-butanol feed upward until TAC is minimized. An illustration plot for Step (3) can be seen in Fig. 5. In this figure, NRXN is set at 21 and the feed tray locations of the propionic acid and n-butanol are at the top and bottom stages of the reactive section, respectively. From this figure, one would observed that NR = 1 and NS = 6 are best in this situation. Another illustration for Step (5) can be seen in Fig. 6. From this figure, HOPr feed tray at 4th stage and BuOH feed tray at 6th stage are best in this situation. Notice that this altering of the feed tray locations represent significant reducing of TAC from the original
Fig. 6. Optimization plot with NR = 1, NRXN = 21, and NS = 6.
assumption of HOPr feed tray at the top of the reactive section (2nd stage) and BuOH feed tray at the bottom of the reactive section (22nd stage). The feed tray locations are counting from top to bottom with top tray to be the first stage and the reboiler to be the last stage. The above two illustration plots show that with NRXN = 21 the best combination of the other design variables are: NR = 1, NS = 6, HOPr feed tray at 4th stage, and BuOH feed tray at 6th stage. Collecting results from other optimized runs, Fig. 7 shows a summarized optimization plot at various NRXN. From this figure, the NRXN = 21 gave the best result with NR = 1, NS = 6, HOPr feed tray at 4th stage, and BuOH feed tray at 6th stage. The minimized TAC is $358,700 with annualized capital cost of $207,400, operating cost of $128,100, and catalyst cost of $23,200. The optimized design flowsheet is thus shown in Fig. 8. The liquid composition profile and reaction rate of the optimized process design can be seen in Fig. 9. The liquid compositions of the HOPr and BuOH are all at their highest values
Fig. 8. Optimized design flowsheet with no feed-splitting.
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Fig. 10. Summarized optimization plot at various split HOPr feed flow rate. Fig. 9. Liquid composition profile and reaction rate of the optimized design with no feed-splitting.
at their feed tray locations. The highest reaction rate is at 5th stage between the two feed tray locations. Notice that the BuOH composition maintained at high value from its feed tray location (6th tray) downward. It is speculated that if any benefit can be gained by splitting HOPr feed into two or more streams with partial HOPr feed entering into the lower part of the reactive section where BuOH composition are at its higher value. For illustration and process simplification purpose, we will only consider splitting the HOPr feed into two sub-feed streams in the following. 3.2. Design flowsheet with HOPr feed-splitting 3.2.1. Retain main HOPr and BuOH feed tray locations Assuming that we retain the main HOPr feed tray location at 4th stage and BuOH feed tray location at 6th stage as in the optimized design flowsheet in previous Fig. 8. Also, from all the previous optimization runs, the stages of rectifying section are all found to be NR = 1, thus this is also retained in the following optimization search. With the allowing of HOPr feed to split, there will be two other design variables to be determined: the split HOPr feed tray location and the split HOPr feed flow rate. The sequential iterative optimization procedure is as follows: (1) Set the split HOPr feed flow rate. (2) Set the split HOPr feed tray location. (3) Varying total number of stages of reactive section (NRXN) and total number of stages of stripping section (NS) until TAC is minimized. (4) Go back to Step (2) to set another split HOPr feed tray location to minimize TAC. (5) Go back to Step (1) to set another split HOPr feed flow rate to minimize TAC. Fig. 10 shows the summarized result of this optimization search at various split HOPr feed flow rate. The TAC is minimized at split feed flow rate of 10 kmol/h. At this design, split HOPr feed tray location is at 6th stage, NRXN = 24, and NS = 1. Notice that the total number of stages of the reactive distillation column (including reboiler) is reduced from 28 to 26. The TAC is reduced from $358,700 to $331,400 (a reduction of 7.6%) with annualized capital cost from $207,400 to $190,500 (a reduction of 8.1%) and operating cost from $128,100 to $116,700 (a reduction of 8.8%). However, because the total number of stages of reactive section is increased, the catalyst cost is increased from $23,200 to $24,200 (an increase of 4.3%).
3.2.2. Allow the main HOPr feed tray location to be varied With the main HOPr feed tray location as an additional design variable, the search for the minimized TAC becomes more timeconsuming. The results in Section 3.2.1 can be considered as a set of optimized results with the main HOPr feed tray location set at 4th stage. By varying the main HOPr feed tray location, other sets of optimized results can be obtained. Table 3 shows the summarized optimization results at various main HOPr feed tray locations. Notice that with the main HOPr feed tray locations set at 1st, 2nd, or 3rd stages, the flow rate at the main HOPr feed tray location is actually smaller than the split HOPr feed flow rate. Fig. 11 shows the summarized optimization results of TAC at various main HOPr feed tray locations. The TAC is minimized with the main HOPr feed tray location at 2nd stage. At this design, split HOPr feed tray location is at 6th stage with flow rate of 30 kmol/h, NRXN = 23, NS = 2, and NR = 1. The main portion of the HOPr feed is entering the column at the same location as the BuOH feed does. The remaining HOPr feed enters the column at the top of the reactive section. The TAC, annualized capital cost, and operating cost are all further reduced from the result in Section 3.2.1. Also notice that the catalyst cost is reduced from the base case in Section 3.1. The reason is that although the number of stages of the reactive section is increased from 21 to 23, the column diameter is also decreased from 1.117 m to 1.048 m, thus resulting in the reduction of catalyst cost from $23,200 to $22,400. With this simple process design alternation by considering feed-splitting, the TAC can be reduced from $358,700 to $321,400 (a reduction of 10.4%) with annualized capital cost from $207,400 to $186,200 (a reduction of 10.2%) and operating cost from $128,100 to $112,800 (a reduction of 11.9%). The optimized design flowsheet is summarized in Fig. 12. The comparison of the specific reaction rate for the original base case with no feed-splitting and the optimized results from Sections 3.2.1 and 3.2.2 can be seen in Fig. 13. Notice that without feedsplitting, the maximum value occurs at 5th stage (between the two feed tray locations) and decreases on both directions upward and downward of the reactive distillation column. However, in the two feed-splitting cases, the profile of specific reaction rate has two peaks when a portion of HOPr feed is moved downward. And according to the composition profiles of optimized design flowsheet with feed-splitting shown in Fig. 14, the HOPr composition is still higher below the BuOH feed stage in reactive zone than the case without feed-splitting shown in Fig. 9. It means that high concentration of both reactants would help the specific
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Table 3 Summarized optimization results at various main HOPr feed tray locations.
Total stages NR NRXN NS HOPr feed flow (kmol/h) Split HOPr feed flow (kmol/h) HOPr feed tray Split HOPr feed tray BuOH feed tray Column diameter (m) Reboiler duty (kW) TAC ($1000/yr) Capital cost ($1000/yr) Operating cost ($1000/yr) Catalyst cost ($1000/yr)
Base case
HOPr at 1st stage
HOPr at 2nd stage
HOPr at 3rd stage
HOPr at 4th stage
28 1 21 6 50 – 4 – 6 1.117 1131 358.7 207.4 128.1 23.2
27 1 25 1 15 35 1 6 6 1.052 1002 327.2 189.2 113.5 24.5
26 1 23 2 20 30 2 6 6 1.048 995.9 321.4 186.2 112.8 22.4
27 1 24 2 25 25 3 6 6 1.046 990.0 324.3 188.9 112.2 23.3
26 1 24 1 40 10 4 6 6 1.067 1030 331.4 190.5 116.7 24.2
Fig. 11. Summarized optimization plot at various main HOPr feed tray locations.
Fig. 12. Optimized design flowsheet with feed-splitting.
Fig. 13. Comparison of the specific reaction rate for three design cases.
Fig. 14. Liquid composition profile and reaction rate of the optimized design with feed-splitting.
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Table 4 Summary of kinetic models, k1, and Keq used in the study. System
Kinetic model (catalyst)
k1 (T = 363 K)
BuOPr
Quasi-homogeneous model (Amberlyst 35) r ¼ mcat ðk1 aHOPr aBuOH k1 aBuOPr aH2 O Þ 1 1 7954 1:6786 1010 exp k1 ¼ 1000 60 T 1 1 7135 3:1085 108 exp k1 ¼ 1000 60 T
8.52 105 [kmol/(kgcat s)]
Keq (T = 363 K)
5.66
Pseudo-homogeneous model (Amberlyst 15) r ¼ mcat ðk1 aHAc aBuOH k1 aBuOAc aH2 O Þ BuOAc
70; 660 RT 74241:7 ¼ 1:0135 106 exp p RT
k1 ¼ 3:3856 106 exp k1
2.32 104 [kmol/(kgcat s)]
10.9
Quasi-homogeneous model (Amberlyst 15) r ¼ mcat ðk1 C HAc C AmOH k1 C AmOAc C H2 O Þ AmOAc
51; 740 RT 45; 280 ¼ 2:2533 exp RT
k1 ¼ 31:1667 exp k1
reaction rate increasing for this 2nd order reaction system. By calculating the summation of the specific reaction rate at each stage, it is found that the design flowsheet in Fig. 12 gave the largest value representing most efficient usage of the reactive volume. 3.3. Extension of this feed-splitting strategy to other heavy ester processes The design flowsheet for the production of butyl proprionate belongs to a general class of the reactive distillation flowsheet called Type III as in the paper by Tang et al. (2005). In that paper, the design flowsheets of butyl acetate (BuOAc) and amyl acetate (AmOAc) have been studied. We have also preliminary investigated the alternative feed-splitting strategy of these two heavy ester flowsheets. Same kinetic and thermodynamic models as in Tang et al. (2005) are used in this study and the design flowsheet of the two processes with no feed-splitting are established. These two base cases are compared with the results from feed-splitting operating strategy. For the feed-splitting design, the main portion of the acid feed is designed to enter the column at the same location as the alcohol feed does. The remaining acid feed is designed to enter the column at the top of the reactive section. This is the design experience from the result of Section 3.2. With no detailed optimization search, the above feed-splitting design has already seen the saving of TAC in comparison with the no feed-splitting design. For the case of butyl acetate, significant saving of TAC over 30% can be realized with this simple design alternation. For the case of amyl acetate, however, the TAC saving is much less (only 2.4%). The reason for the discrepancy in TAC savings on butyl proprionate, butyl acetate, and amyl acetate can be explained by their equilibrium rate constants. Table 4 shows the kinetic models and their model parameters used in the simulation. The table also listed their forward reaction rate constants and the equilibrium rate constants at 363 K. Notice that the equilibrium rate constant of the BuOAc system are the greatest and that of the AmOAc system are the smallest. Thus, the BuOAc system has the most potential for improving the reaction performance while the AmOAc system has little room for any further improvement.
1.13 106 [m6/(kmol kgcat s)]
1.6
4. Overall control strategy The inventory control loops of the design flowsheet in Fig. 12 will be determined first. The three level control loops are: bottom level controlled by the bottom product flow; organic level controlled by the organic reflux flow; and the aqueous level controlled by the aqueous product flow. Other control loops include the top column pressure controlled at 1.1 atm by manipulating the top vapor flow and the decanter temperature controlled at 40 8C by manipulating the condenser duty. There are also two ratio-control schemes implemented in the overall control strategy. One of it is to maintain the ratio of flow rates between the splitting HOPr feed (at 6th stage) and the HOPr feed (at 2nd stage) at 1.5. The other ratio scheme is to keep the ratio of flow rates between the total HOPr feed and the BuOH feed at 1.0. Notice that the above ratio scheme of keeping the ratio of flow rates between the total HOPr feed and the BuOH feed at 1.0 is not suitable when either of the feed compositions change. For example, if BuOH feed composition changes from pure to include some water in this feed stream, this ratio should be decreased to allow less HOPr feed from entering the column. This set value of the ratio should be adjusted by a product quality control loop. In this paper, we assumed that the on-line composition measurement is not readily available, thus, the product quality control loop will be replaced by some tray temperature control loop. Beside the ratio mentioned above can be adjusted by a tray temperature control loop, another temperature control loop is needed to hold the temperature profile inside the column by manipulating the reboiler duty. Therefore, a dual-temperature control strategy commonly used in reactive distillation column will be investigated in this system. The two temperature control points can be determined by open-loop sensitivity analysis. Figs. 15 and 16 show the open-loop sensitivity test by changing either of the ratio of the two feeds or the reboiler duty for 0.1% and plotting the tray temperature differences versus the base case condition. From these two figures, it is determined that the 10th stage temperature should be controlled by manipulating the reboiler duty and the 22nd stage temperature should be controlled by manipulating the ratio of the two feeds.
H.-Y. Lee et al. / Journal of the Taiwan Institute of Chemical Engineers 41 (2010) 403–413
Fig. 15. Open-loop sensitivity test with 0.1% changes in the feed ratio.
411
Fig. 16. Open-loop sensitivity test with 0.1% changes in the reboiler duty.
Fig. 17. Closed-loop simulation with 10% changes in the BuOH or HOPr feed composition.
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Fig. 18. Closed-loop simulation results with and without feedforward compensation for +10% throughput changes.
Fig. 17 shows 10% changes in the BuOH and HOPr feed compositions. In this two simulation runs, the feed compositions were changed to include 10% water in the original pure feed stream. It is noticed that the dual-temperature control strategy works well despite these feed composition changes. The purity of the BuOPr product stream was closely maintained (99.5 wt% BuOPr and 100 ppm HOPr impurity). For the throughput changes, the above dual-temperature control strategy was not able to hold the high-purity product specifications. However, throughput change is considered as a measurable disturbance, thus some feedforward compensation as in Huang et al. (2004) can be used to circumvent the above problem. The set-point of the 22nd stage temperature was determined as a linear function of the throughput (nominal case at 50 kmol/h BuOH feed flow rate). For a 10% BuOH feed flow rate change, the set-point of this temperature loop was adjusted
from 121.96 8C to 125.72 8C in order to maintain the high-purity product specifications. The set-point of the other temperature control loop (10th stage temperature) does not need to be changed. Fig. 18 shows such simulation runs with and without this feedforward compensation for 10% throughput changes. It is clearly shown that the HOPr impurity is reduced from 155 ppm to 93 ppm with the implementation of this feedforward compensation. The final proposed overall control strategy can be seen in Fig. 19. With HOPr feed being splitting into two sub-streams, an additional degree of freedom can be utilized by the overall control strategy. However, simply holding the two sub-stream flow rates at a constant ratio, the abovementioned overall control strategy can already properly handle feed composition and throughput disturbance. Thus, there is no need to develop other more complicated control strategy to use this extra degree of freedom.
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Fig. 19. Proposed overall control strategy.
5. Conclusions In this study, the design and control of a reactive distillation column system for the production of butyl propionate has been investigated. Significant TAC and operating costs can be reduced by properly locate the two feed streams inside the reactive section. By simply splitting the HOPr feed into two sub-streams and entering into the reactive distillation column at different appropriate locations can further save TAC by 10.4% and cut the required operating cost by 11.9%. This study has been extended to other heavy ester systems and saving of TAC has also been shown. The percentage of the saving can be qualitatively approximated by the value of the equilibrium rate constant. Bigger saving can be realized with this simple feed-splitting design if the value of the equilibrium rate constant is large. A simple overall control strategy for this process with dual-temperature control has also been proposed. The high-purity product specifications can be maintained in the face of feed compositions and throughput changes. References Cheng, Y. C. and C. C. Yu, ‘‘Effects of Feed Tray Locations to the Design of Reactive Distillation and its Implication to Control,’’ Chem. Eng. Sci., 60, 4661 (2005). Douglas, J. M., Conceptual Design of Chemical Processes, McGraw-Hill Inc., New York, USA (1988).
Gmehling, J., J. Menke, J. Krafczyk, and K. Fischer, Azeotropic Data,’’ 2nd ed., Wiley–VCH, Weinheim, Germany (2004). Gmehling, J., U. Onken, and W. Arlt, Vapor–Liquid Equilibrium Data Collelection, Chemistry Data Series,’’ vol. 1. DECHEMA, Frankfurt/Main, Germany (1977). Gonzales, E. and J. Ortega, ‘‘Densities and Isobaric Vapor–Liquid Equilibria for the Mixtures Formed by Four Butyl Esters and 1-Butanol,’’ J. Chem. Eng. Data, 41, 53 (1996). Hayden, J. G. and J. P. O’Connell, ‘‘A Generalized Method for Predicting Second Virial Coefficients,’’ Ind. Eng. Chem. Proc. Des. Dev., 14, 209 (1975). Huang, S. G., C. L. Kuo, S. B. Hung, Y. W. Chen, and C. C. Yu, ‘‘Temperature Control of Heterogeneous Reactive Distillation,’’ AIChE J., 50, 2203 (2004). Lee, M. J., J. Y. Chiu, and H. M. Lin, ‘‘Kinetics of Catalytic Esterification of Propionic Acid and n-Butanol over Amberlyst 35,’’ Ind. Eng. Chem. Res., 41, 2882 (2002). Lee, M. J., L. H. Tsai, G. B. Hong, and H. M. Lin, ‘‘Multiphase Equilibria for Binary and Ternary Mixtures Containing Propionic Acid, n-Butanol, Butyl Propionate, and Water,’’ Fluid Phase Equilibria, 216, 219 (2004). Liu, W. T. and C. S. Tan, ‘‘Liquid-Phase Esterification of Propionic Acid with n-Butanol,’’ Ind. Eng. Chem. Res., 40, 3281 (2001). Luyben, W. L., Plantwide Dynamic Simulators in Chemical Processing and Control, Marcel Dekker, New York, USA (2002). Malone, M. F. and M. F. Doherty, ‘‘Reactive Distillation,’’ Ind. Eng. Chem. Res., 39, 3953 (2000). Sharma, O. N., G. D. Nageshwar, and P. S. Mene, ‘‘Esterification of Butanol with Propionic Acid Catalysed by a Cation Exchange Resin,’’ Indian J. Technol., 11, 360 (1973). Sørensen, J. M. and W. Arlt, Liquid–Liquid Equilibrium Data Collelection, Chemistry Data Series,’’ vol. 5. DECHEMA, Frankfurt/Main, Germany (1979). Sundmacher, K. and A. Kienle, Reactive Distillation: Status and Future Directions, Wiley– VCH, Weinheim, Germany (2003). Tang, Y. T., Y. W. Chen, H. P. Huang, C. C. Yu, S. B. Hung, and M. J. Lee, ‘‘Design of Reactive Distillations for Acetic Acid Esterification,’’ AIChE J., 51, 1683 (2005).