Design and Control of Thermally-Coupled Reactive Distillation System

Article pubs.acs.org/IECR

Design and Control of Thermally-Coupled Reactive Distillation System for Esterification of an Alcohol Mixture Containing n‑Amyl Alcohol and n‑Hexanol Yi-Chang Wu,† Hao-Yeh Lee,‡ Chung-Han Lee,† Hsiao-Ping Huang,† and I-Lung Chien*,† †

Department of Chemical Engineering, National Taiwan University, Taipei 10617, Taiwan Department of Chemical Engineering, National Taiwan University of Science and Technology, Taipei 10607, Taiwan

‡

ABSTRACT: Semiconductor and pharmaceutical industries often produce waste alcohol mixtures. This study discusses a reactive distillation process of a mixed alcohol feed containing n-amyl alcohol and n-hexanol to react with acetic acid to produce useful esters. In an earlier paper (Lee et al. Ind. Eng. Chem. Res. 2009 48, 7186−7204), optimal design of the same Type-III mixed-alcohol reactive distillation process containing n-butanol and n-amyl alcohol was developed. In that paper, the indirect sequence containing a reactive distillation column with heavier amyl acetate bottom product and a second column to produce butyl acetate gives the lowest total annual cost. In this study, with different mixed alcohol feed (although of the same reaction type), it is found that an alternative direct sequence design is much more economically favorable with 34% reduction in total annual cost as compared to the indirect sequence. The reason for the seemingly contradictory results can be explained by differences in relative volatilities of binary pairs in the stripping sections of the reactive distillation towers. Process intensification technology is also attempted to devise a thermally coupled reactive distillation configuration to further save 13% of the operating energy. The overall control strategy using tray temperatures is also proposed to maintain high-purity products despite various disturbances. work of Tang et al.9 In the work of Lee et al.,1 a mixed n-butanol and n-amyl alcohol RD process has been developed. In that paper, an indirect sequence design containing a RD column with heavier amyl acetate bottom product and a second column to produce butyl acetate gives the lowest total annual cost. In this paper (although classified as the same Type-III system), we will demonstrate that the more economical design of this mixed alcohol system does not use the same design configuration. Investigating of heat-integrated reactive distillation design for a mixed alcohol feed in the open literature is even lesser. Mueller et al.10 was the only paper we can find to combine the RD with dividing wall configuration as reactive dividing wall column (RDWC) for simultaneous esterification of methanol and butanol with acetic acid to yield methyl acetate, butyl acetate, and byproduct water. Since no paper can be found to study the heat integration of mixed alcohol system classified as Type-III system, the investigation in this paper should be valuable to other researchers. After the optimal design of this system is accomplished, the overall control strategy of the thermally coupled RD system will also be investigated. The goal is to use tray-temperature control strategy to hold the purities of the two ester products despite various disturbances.

1. INTRODUCTION Semiconductor and pharmaceutical industries often produce waste alcohol mixtures. Therefore, esterification of the alcohol mixture to coproduce two valuable esters is a more beneficial way in reusing waste from these industries. The purpose of this paper is to use an alcohol mixture of n-amyl alcohol (AmOH) and n-hexanol (HexOH) as a feed stream to react with acetic acid (HAc) to coproduce n-amyl acetate (AmAc) and n-hexyl acetate (HexAc) using reactive distillation (RD) technology. There are numerous papers and applications of RD in the open literature. A good review paper of this technology was written by Malone and Doherty.2 From the book of Luyben and Yu,3 the above survey in Malone and Doherty2 was updated to include 1105 related publications and 814 US patents between 1971 and 2007. Luyben and Yu3 also highlighted 236 reaction systems which can be designed with RD configuration. Another book of Sundmacher and Kienle4 also surveyed over one hundred industrially or potentially important reactions for RD applications. The above references show the importance of the RD technology in industrial applications. However, existing literature on investigating the design or control of RD process for simultaneous esterification of an alcohol mixture with acetic acid is relatively scarce. There are a few patents5−7 discussed processes to coproduce ethyl acetate and n-butyl acetate. The resulting design flowsheets all have problems of having lower purity in one of the main products or the water byproduct. Wu et al.8 discussed the design and control of an improved RD process to coproduce higher purity ethyl and n-butyl acetate products. This above simultaneous esterification RD process can be classified as a mixed Type-II/Type-III system because ethyl acetate RD process was classified as Type-II system and that of n-butyl acetate was classified as Type-III system in the © 2013 American Chemical Society

2. REACTION KINETICS AND PHASE EQUILIBRIUM The parallel and simultaneous esterification reactions in this system are described as the following two reversible equations. Received: Revised: Accepted: Published: 17184

February 24, 2013 October 9, 2013 November 7, 2013 November 7, 2013 dx.doi.org/10.1021/ie4006035 | Ind. Eng. Chem. Res. 2013, 52, 17184−17197

Industrial & Engineering Chemistry Research

Article

Table 1. Kinetic Equations of the Mixed Alcohol Processa Kforward T = 402 K

system

Keq T = 402 K

5.9 × 10−6 [m6 cat/(kmol·kg cat·s)]

(1) AmAc

2

r = mcat (k1C HAcCAmOH − k −1CAmAcC H2O) k1 = 31.1667 exp(−51740/RT) k−1 = 2.2533 exp(−45280/RT) (2) HexAc

2.09 × 10−3∼4.90 × 10−5 [kmol/(kg cat·s)] (for XHexOH = 0−1)

29.5

r = mcat(k 2aHAcaHexOH − k −2aHexAca H2O) k2 = [4883.18 − 4769.06XHexOH]exp(−49000/RT) k−2 = [4883.18 − 4769.06XHexOH]exp(−46215/RT) a

R = 8.314 [kJ/kmol·K], T [K], r [kmol/s], mcat [kg cat], Ci [kmol/m3 cat], XHexOH = mole fraction of n-hexanol.

Table 2. NRTL Model Parameters for the AmAc and HexAc Process comp i

HAc

HAc

HAc

HAc

HAc

comp j

AmOH

HexOH

AmAc

HexAc

H2O

source

Lee et al.1

Schmitt and Hasse13

Lee et al.1

Schmitt and Hasse13

Lee et al.1

−2.9897 5.8822 801.34 −1264.46 0.2 AmOH

0 0 −110.6 424.02 0.2987 HexOH

aij aji bij (K) bji (K) cij comp i

0 0 41.866 150.87 0.3 AmOH

−2.8409 0.8255 1091.78 −351.26 0.2 AmOH

comp j

HexOH

AmAc

0 0 −37.94 214.55 0.2 AmOH HexAc 1

source

Aspen VLE-HOC

Lee et al.

aij aji bij (K) bji (K) cij comp i

1.0548 1.005 −536.7770 −211.1423 0.3 HexOH

0 0 −144.8 320.65 0.3009 HexOH

H2O

Aspen UNIFAC

Lee et al.

0 0 127.2314 150.269 0.3 AmAc

AmAc 1

0 0 57.328 1424.8 0.2869 AmAc

Aspen UNIFAC 0 0 67.7188 191.4062 0.3 HexAc

comp j

HexAc

H2O

HexAc

H2O

H2O

source

Schmitt and Hasse13

Schmitt and Hasse13

Aspen UNIFAC

Lee et al.1

Schmitt and Hasse13

aij aji bij (K) bji (K) cij

2.6355 −3.2599 −839.81 1225.85 0.3

−3.1777 −0.1522 1381.31 1945.07 0.3323

0 0 126.173 −111.433 0.3

0 0 254.47 2221.5 0.2

−1.3148 −1.7481 998.70 3545.58 0.2

acetic acid in vapor phase. The NTRL parameters associated with the AmOH esterification reaction are taken from the work of Lee et al.,1 and that of the HexOH esterification reaction are taken from the work of Schmitt and Hasse.13 The binary NRTL parameters associated with both esterification reaction systems were regressed in the previous papers to fit the HAc−water, HAc−alcohol, HAc−ester, and alcohol−ester pairs by VLE data, and to fit ester−water and alcohol−water by LLE data. However, the above binary NRTL parameters may have to be readjusted in order to also describe reasonably well the ternary LLE boundaries and also the azeotropic information. For the remaining mixed-pairs, the binary pair of AmOH−HexOH is taken from Aspen built-in NTRL parameters (only VLE data are available in literature to obtain the model parameters). For the other three mixed-pairs (AmOH−HexAc, HexOH−AmAc, and AmAc−HexAc) with no Aspen built-in parameters or experimental data, NRTL parameters are estimated by UNIFAC group contribution method. Table 2 summarized the NTRL parameters used in the simulation. The boiling points rankings for pure components and the azeotropic information of all the azeotropes for the studied

k1

HAc + AmOH XooY AmAc + H 2O k −1 k2

HAc + HexOH XooY HexAc + H 2O k −2

(1)

The rate equation for the AmOH esterification reaction is taken from the work of Lee et al.,1 and that of the HexOH esterification reaction is taken from Schmitt and Hasse.11 Both are heterogeneous reactions using acid ion exchanged resin as catalyst. The AmOH esterification reaction is concentrationbased, and the HexOH esterification reaction is activity-based. Table 1 summarizes the rate equations with the kinetic parameters. Both the forward reaction rate and the equilibrium constant of the HexOH esterification reaction are larger indicating that this reaction is faster and it is easier to reach equilibrium. As for predicting the phase equilibriums of this system, NRTLHOC property method is used in the Aspen Plus simulation. The NRTL model is selected to describe the nonideal VLE and possibly VLLE behavior, and the HOC (Hayden−O’Connell12) model is used to account for nonideality such as dimerization of 17185

dx.doi.org/10.1021/ie4006035 | Ind. Eng. Chem. Res. 2013, 52, 17184−17197


Article

Table 3. Compositions and Temperatures for the AmOH/ HexOH System experimental data component

mole fraction temp °C

AmOH/AmAc/H2Oa (0.046, 0.107, 0.847) AmAc/H2Oa (0.166, 0.834) AmOH/HexAc/H2Oa AmOH/H2Oa (0.146, 0.854) HexOH/HexAc/H2Oa (0.235, 0.060, 0.705) H2O/HexAca (0.926, 0.074) H2O/HexOHa (0.926, 0.074) H2O 1 HAc 1 AmOH 1 HexOH/AmAc AmAc 1 HexOH 1 HexAc 1 a

94.9 95.2

95.8 96.00 97.40 97.80 100.02 118.01 137.68 147.71 157.63 171.21

computed data mole fraction

temp °C

(0.0418, 0.1352, 0.8230) (0.1696, 0.8304) (0.1203, 0.0307, 0.849) (0.1471, 0.8529) (0.0379, 0.0534, 0.9087) (0.9168, 0.0823) (0.9268, 0.0732) 1 1 1 (0.7707, 0.2293) 1 1 1

94.76 94.90 95.79 96.00 97.37 97.58 97.96 100.02 118.01 137.68 146.41 147.71 157.63 171.20

Heterogeneous azeotrope.

Figure 2. (a) RCM and LLE of AmOH−AmAc−H2O ternary system at 1 atm. (b) RCM and LLE of HexOH−HexAc−H2O ternary system at 1 atm.

experimental data of Gmehling et al.,14 especially on the order of temperature ranking.

3. DESIGN OF RD PROCESS FOR SIMULTANEOUS ESTERIFICATION 3.1. Alternative Conceptual Design Flowsheets. For simultaneously reacting the two alcohols with acetic acid to produce two heavy esters in a single RD column, Figure 1 displays two possible conceptual design flow sheets. Figure 1a shows an indirect-sequence design to draw out the heaviest component (HexAc) at bottoms of the RD column first. The top vapor of the RD column should approach the lowest temperature (94.76 °C at 1 atm) of the system which will be the AmOH/ AmAc/H2O ternary azeotrope (as seen in Table 3 of the boilingpoint rankings). From RCM and LLE plots of Figure 2, this ternary azeotrope after condensation, can naturally be separated into two liquid phases. The aqueous phase containing mostly water can be drawn-out of the system. The organic phase should go to another purification column to obtain pure AmAc at bottoms of this second column. The distillate product of this column still contains fair amount of the unreacted AmOH should be recycled back to the RD column. Note from the experience for the study of the same RD type (Lee et al.1 for the BuOH/AmOH system), portion of the aqueous stream should be refluxed back to the RD column to carry the product (second ester) out from the top of this RD column via ternary azeotrope.

Figure 1. Alternative conceptual design flowsheet: (a) indirectsequence design; (b) direct-sequence design.

system are shown in Table 3. The predictions of the azeotropic temperatures and compositions are in good agreement with the 17186



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Figure 3. Optimal direct-sequence design flowsheet for the mixed BuOH and AmOH system in the work of Lee et al.1

simpler which contains no recycle stream from the second column to the RD column. Previous study for the BuOH/AmOH system showed that the indirect-sequence design (in Figure 1a) is much more economical than the direct-sequence design. The main reason for the drawback of the direct-sequence design is because of excess large number of stages in the stripping section of the RD column. Figure 3 shows the previous design flowsheet of direct-sequence design for the BuOH/AmOH system in the work of Lee et al.1 Note that the stripping section is from 29th stage to 116th stage. The reason for excess large number of stages in the stripping section can be explained by the yx plot of HAc/BuAc in Figure 4a. Note that there is tangent pinch near pure BuAc end. Thus in order to prevent HAc from going out of the RD column bottoms, excess large numbers of stages in the stripping section is needed. The situation is quite different for the AmOH/HexOH system studied in this paper. From Figure 4b, it is noted that the separation of HAc/AmAc is much easier. Thus, it is conjecture that the direct-sequence design for the AmOH/HexOH system should be more competitive. 3.2. Direct-Sequence Design. The optimal design flowsheet of direct-sequence design will be established in this section. Several assumptions are made in the following simulation studies. The mixed alcohols feed stream (AmOH/HexOH) is set to be equal molar, and the total flow rate is 100 kmol/h. The feed flow rate of pure reactant HAc is also set at 100 kmol/h. The decanter temperature is set at 40 °C. Product specifications are set as follows: xAmAc = 99 mol %; xHexAc = 99 mol %. Total annual cost (TAC) calculations for establish the optimal design flowsheet are based on information provided by Douglas15 with TAC includes annual total operating cost plus total installed capital cost divided by a payback period. Operating cost includes the costs of steam, cooling water, and the catalyst, and capital cost includes the costs of columns, trays, reboilers, and condensers. The payback period is assumed to be three years, and the wastewater treatment cost is assumed to be negligible in the overall TAC.

Figure 4. yx plots at 1 atm: (a) HAc/BuAc system; (b) HAc/AmAc system.

The second alternative design is the one in Figure 1b, denoted as direct-sequence design. The design thinking is to first draw out the water byproduct from the RD column. The bottom stream of the RD column is designed to contain two heavy ester products (AmAc and HexAc in this case). The purpose of second column is to separate two esters. Note that this design flowsheet is 17187



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Figure 5. Optimal design flowsheet of the direct-sequence for the mixed AmOH and HexOH system.

Figure 6. Optimal design flowsheet of the indirect-sequence for the mixed AmOH and HexOH system.

section (Nrxn), the rectifying stage numbers of RD (NR), the stripping stage numbers of RD (NS), the HAc feed location (NFHAc), the mixed alcohols feed location (NFOH), the total stage numbers of the second column (NT), and the feed location (NF) in the second column. Two degree-of-freedoms in the distillation column are used to meet two purity specifications. The reboiler duty of the RD column is another free variable which also used to minimize TAC. Exhaustively iterative optimization procedure was used to find the optimal design flowsheet. Although time-consuming, all

In the simulations, the tray weir heights of RD and distillation columns are set to be 0.1016 and 0.0508 m, respectively. For each simulation run, the reactive tray holdup has to be iteratively obtained to agree with the column sizing calculation. Possible liquid phase splitting will be checked by RADFRAC simulator on each tray. If there is a second liquid phase on a particular tray, a vapor phase composition and two liquid phase compositions will be exhibited in the simulation result. There are eight design variables needed to be determined in this design flowsheet including: the stage numbers of the reactive 17188



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Table 4. Optimal TAC and Individual Cost of Each Case direct-sequence configurations annualized capital cost for column shell ($1000/y) annualized capital cost for column tray ($1000/y) annualized capital cost for reboiler ($1000/y) annualized capital cost for condenser ($1000/y) steam cost ($1000/y) cooling water cost ($1000/y) catalyst cost ($1000/y) total reboiler duty (KW) (% difference) total steam cost ($1000/y) (% difference) TAC ($1000/y) (% difference)

C1 269.83 56.31 87.29 123.59 289.36 8.50 104.46 4725.55 (0%) 536.97 (0%) 1656.48 (0%)

indirect-sequence

C2 194.02 36.99 193.65 37.37 247.61 7.50

C1

C2

437.71 219.34 108.50 45.81 123.87 115.09 181.01 57.88 552.84 432.59 15.14 10.47 206.55 8044.58 (+70.25%) 985.43 (+70.25%) 2506.80 (+51.33%)

thermally coupled C1

C2

273.01 222.60 57.28 45.17 143.03 86.05 31.01 463.86 16.84 5.63 103.77 4103.82 (−13.16%) 463.86 (−13.16%) 1448.25 (−12.56%)

Figure 8. Conceptual design of the thermally coupled RD system.

(6) Change the reflux ratio and reboiler duty of the second column until the two product specifications (xAmAc = 99 mol %; xHexAc = 99 mol %) are met. (7) Go back to step 5 and change the RD reboiler duty until the TAC is minimized. (8) Go back to step 4 and change NT and NF until the TAC is minimized. (9) Go back to step 3 and change NFHAc and NFOH until the TAC is minimized. (10) Go back to step 2 and change NR and NS until the TAC is minimized. (11) Go back to step 1 and vary Nrxn until the TAC is minimized. Figure 5 shows the optimal flowsheet of the direct-sequence design after completing the steps above. The optimal Nrxn is 40 stages, NR is 2 stages, and NS is 5 stages in the RD column, and the feed tray locations NFOH and NFHAc are on the first and fifth stages in the upper section of the RD column. The total number of stages and the feed tray location for the second column are 37 stages and at the 20th stage, respectively. Note particularly the stripping section (NS) is only 5 stages as compared to 88 stages for the BuOH/AmOH system. 3.3. Indirect-Sequence Design. The optimal flowsheet of indirect-sequence design is also established for comparison purpose. Same feed conditions and product specifications are maintained in the simulations. There are extra design variables in this flowsheet including the feed location of recycle stream from the overhead of second column to the RD column and the aqueous reflux split ratio. The HexAc purity requirement is set the same at 99 mol % by varying the RD reboiler duty. The AmAc purity requirement is also set the same at 99 mol % by varying the

Figure 7. Liquid composition profile of the RD column: (a) directsequence design; (b) thermally coupled direct-sequence design.

combinations of the design variables were investigated to obtain the minimization of the TAC. The optimization procedure is described in the following steps: (1) Guess Nrxn. (2) Guess NR and NS. (3) Guess NFHAc and NFOH. (4) Guess NT and NF. (5) Guess the RD reboiler duty 17189



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Figure 9. Optimal flowsheet of the thermally coupled direct-sequence RD system.

Figure 10. Inventory and basic control loops of the thermally coupled RD system.

to be taller with more stages in the reactive section and also in the stripping section. The overall reboiler duty (4126.42 + 3918.16 = 8044.58 KW) is also 70.2% more as compared to that of the direct-sequence design (2560.00 + 2165.55 = 4725.55 KW). Table 4 displays the individual items in the TAC calculations for

reboiler duty of the second column. An extra operating variable (reflux ratio of the second column) is set to minimize the TAC. Similar exhaustively iterative sequential optimization procedure was used to find the optimal flowsheet for this indirectsequence design as in Figure 6. Note that the RD column needs 17190



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Figure 11. Closed-loop sensitivity analysis plots. Figure 12. Open-loop sensitivity analysis plots.

both designs. TAC of indirect-sequence design is 51.33% higher than the proposed direct-sequence design.

the total stage numbers of the second column (NT) in Figure 5 are used here. The TAC of thermally coupled design is minimized by varying the sidedraw location and the vapor draw-off rate. The two remaining variables (reboiler duty and reflux ratio) are used to maintain two product specifications at 99 mol %. The resulting thermally coupled design flowsheet is shown in Figure 9. The elimination of the remixing effect can be seen in Figure 7b. Table 4 also displays individual items in the TAC calculations for the thermally coupled design. It is shown that 12.54% in TAC, and 13.16% in steam cost can further be saved by using this thermally coupled design.

4. THERMALLY-COUPLED DIRECT-SEQUENCE DESIGN FOR FURTHER ENERGY SAVINGS By checking the liquid composition profile of RD column for the direct-sequence design in Figure 7a, the remixing effect at the bottom section is observed. AmAc composition is at its highest at 44th stage of the stripping section and then dropped to a lower purity at the column bottoms. This remixing effect can be eliminated by thermally coupled design to further save energy. The concept of thermally coupled direct-sequence design is shown in Figure 8. The top parts of two columns are remained the same with two reboilers combined into one. In order to provide vapor traffic inside of the RD column, a vapor sidedraw from second column should be designed. Bottom liquid draw-off from RD column is designed to be fed into the second column at the same location as the vapor sidedraw. In order to simplify the optimization procedure, column configuration of the direct-sequence design was kept the same. The optimal design variables of Nrxn, NR, NS, NFHAc, NFOH, and

5. OVERALL CONTROL STRATEGY Because the thermally coupled design in Figure 9 results in lower TAC and steam cost than the other two designs in Figures 5 and 6, this process is subject to further control study in this section. The proper conventional control strategy of this process will be investigated. Tray temperature control strategy will be used to indirectly hold two product purities in the face of various 17191



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Figure 13. Closed-loop responses for mixed alcohol composition changes when selecting control point at 42nd stage for the RD column.

needed. A 10 min holdup time with 50% liquid level is used to calculate the volume of each column’s sump and reflux drum. As for the decanter, 20 min holdup time is considered. Top pressures of the both columns are set at atmospheric pressure. Tray rating tool in Aspen Plus is used to calculate the pressure drop inside the column. In the dynamic simulation, three REDFRAC units in Aspen Plus Dynamics are used to represent the thermally coupled design flowsheet in Figure 9. 5.1. Inventory Control. The main control objective is to maintain AmAc and HexAc purities at 99 mol %. There are nine

disturbances. One measured disturbance and two unmeasured disturbances will be tested in this section. The measured disturbance will be the throughput change via one of the fresh feed flow rates. The two unmeasured disturbances will include the feed composition change in mix alcohol or HAc feed, and the other will be the inevitable vapor split variations to the RD column and to the second column. Pressure-driven simulation in Aspen Dynamics is used in the control strategy development. Before converting the Aspen Plus result to Aspen Plus Dynamics, sizing of all equipments are 17192



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Figure 14. Overall control strategy of the thermally coupled RD system.

disturbance tests by varying this ratio value will be introduced in the dynamic simulation. 5.2. Selection of Temperature Control Points. The remaining manipulated variables are the reflux flow rate and the reboiler duty of second column and the feed ratio of the RD column. The tough disturbance of unmeasured fresh feed composition change and vapor ratio change will be considered first. Combinations of closed-loop and open-loop sensitivity analyses are used to select tray temperature control points. The idea is to perform closed-loop simulations to perfectly control the process subjecting to unmeasured disturbances and then looking for the tray(s) that have least temperature deviations comparing to base case. By holding the temperature at these control points, it is expected that the product purities can indirectly be maintained. In order to effectively control those tray temperatures under disturbances, large enough open-loop sensitivity between the manipulated variable paired to the control point must be available. Thus, both closed-loop sensitivity plots in Figure 11 and also open-loop sensitivity plots in Figure 12 must be examined together for the determination of the control points. The way to perform closed-loop sensitivity analyses is to do simulation with the presence of one of the unmeasured disturbances. In the closed-loop simulation, reflux flow rate in DC1 is varied to hold AmAc purity at 99 mol % and reboiler duty in DC2 is used to hold HexAc purity at 99 mol %. As for the feed

inventory control loops included six level controls and three pressure loops which are shown in Figure 10. Note that the second column is broken down into two parts (DC1 and DC2) for easier simulation. Conventional inventory control strategy by picking the manipulated variable having the most influence on the controlled variable (level or pressure) is adopted. The nine inventory loops are the following: top pressure of RD column is controlled by manipulating top vapor flow; bottom level of RD column is controlled by manipulating bottom flow; decanter aqueous phase level is controlled by manipulating aqueous outlet flow; decanter organic phase level is controlled by manipulating organic reflux flow; reflux drum level of DC1 column is controlled by manipulating distillate flow rate; top pressure of DC1 column is controlled by manipulating the condenser duty; top pressure of DC2 column is controlled by manipulating brake power of a fictitious compressor; the levels of the base for both DC1 and DC2 columns are controlled by manipulating their bottom flows. Note that the fictitious compressor between DC1 and DC2 is installed in the dynamic simulation to ensure vapor traffic from DC2 to RD or DC1. This fictitious compressor was also used in the work of Ling and Luyben16 in their divided-wall column configuration. There is another decanter temperature loop setting the decanter temperature to be 40 °C by manipulating cooling duty. Note that in Figure 10, vapor split ratio to RD1 and DC1 is fixed in the simulation. Since in real situation holding this ratio at a constant value may not be feasible, 17193



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Figure 15. Closed-loop responses of the proposed control strategy for mixed alcohol and HAc feed composition changes.

ratio, it is adjusted to be at stiochiometric ratio assuming the unmeasured disturbance is known. Figure 11 shows the results of perfect product composition controls under various unmeasured disturbances. It is observed that the least temperature differences occurred at 42nd stage of RD, 16th stage of DC1, and 9th stage of DC2 for alcohol feed composition changes. However, by confirming the open-loop sensitivity of the manipulated variables with respect to the three control points in

Figure 12, there is potential problem for controlling 42nd stage temperature of RD by manipulating feed ratio. It is observed that there is almost no change in this temperature for −1% change in the feed ratio. This control problem will be confirmed by a closed-loop simulation later. For avoiding this problem, the control point must be higher in the RD so that enough open-loop sensitivity from feed ratio exists. Figure 12 also marked an alternatively control point at 35th stage with large enough open-loop sensitivity. 17194



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Figure 16. Closed-loop responses of the proposed control strategy for ±5% vapor ratio changes.

loop pairings are assumed in the above analyses with a stage temperature in RD, in DC1, and DC2 controlled by manipulating feed ratio, reflux flow, and reboiler duty, respectively. 5.3. Closed-Loop Performance. A confirmation run via closed-loop simulation is done first to see if controlling at 42nd stage of RD is feasible or not. The alcohol feed is assumed to have unmeasured composition disturbance at a time of 5 h with feed composition changed from equal molar to 55 mol % AmOH and 45 mol % HexOH or 45 mol % AmOH and 55 mol % HexOH.

There is no foreseeable problem for the other two control points. For 16th stage of DC1, the open-loop sensitivity from the assigned manipulated variable (reflux flow) to this temperature is large enough and fairly linear. For ninth stage of DC2, the openloop sensitivity from the assigned manipulated variable (reboiler duty) to this temperature is also large enough but more nonlinear than the DC1 case. The potential control problem due to nonlinearity is the reason why ninth stage of DC2 is selected but not the other stages lower than ninth stage. Note that trivial control 17195



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Figure 17. Closed-loop responses of the proposed control strategy for ±10% throughput changes.

The three control points are 42nd stage of RD, 16th stage of DC1 and 9th stage of DC2. Figure 13 shows the simulation results with these two feed composition disturbances. Note that temperature at 42nd stage of RD was not able to increase to the set point value by decreasing of the feed ratio for the 55 mol % AmOH and 45 mol % HexOH case. After around 23 h, the feed ratio dropped to zero and the control strategy fell apart. Several alternative control points were tried at RD to see which stage temperature was able to be controlled by the feed ratio with

enough open-loop sensitivity. The other goal is to have the least deviations of product compositions during the disturbance changes. It was found that 35th stage gave the best closed-loop control performance. The overall control strategy of the thermally coupled reactive column system is summarized in Figure 14. In all the closed-loop runs, the relay feedback test17 and the Tyreus−Luyben18 PI tuning rule are selected to determine the controller tuning parameters of the three crucial temperature control loops. The sequential iterative tuning procedure19 is used 17196



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to find the final controller settings. The final tuning parameters are the following: for RD, Kc1 = 4.29, τI1 = 160 (min); for DC1, Kc2 = 86.6, τI2 = 26.4 (min); for DC2, Kc3 = 57.3, τI3 = 39.6 (min). It can be seen that large reset time is associated with the feed ratio loop and the reset time for the other two loops are relatively small. This implies that the three loops are dynamically decoupled with the feed stoichiometric balance maintaining in a slow and smooth manner while the product quality is tightly controlled. Figure 15 displays closed-loop performance under three different feed composition disturbances. The solid and dashed lines are with alcohol feed composition change from equal molar to 55 mol % AmOH and 45 mol % HexOH or to 45 mol % AmOH and 55 mol % HexOH at a time of 5 h. The third feed composition change is the one with HAc feed from pure to contain 5 mol % water at a time of 5 h. The first observation from Figure 15 is that all three controlled temperatures are back to their set point values within 10 h. Compositions of three outlet streams (HexAc, AmAc, and water) are all maintained at highpurity. For alcohol feed composition changes, the ester flow rates are changed according to the distribution of the two alcohols in feed stream. For HAc feed composition disturbance, the aqueous outlet flow rate is properly increased to draw-off more water in the system. Figure 16 displays closed-loop performance for ±5% vapor fraction disturbances going up to either RD or DC1. This disturbance is inevitable because there will not be any regulation of the vapor split. Again, the closed-loop performance is satisfactory in maintaining all product compositions at high-purity despite these disturbance changes. The final closed-loop test is to change the throughput of this process by increasing or decreasing the flow rate of the alcohol feed. Figure 17 displays closed-loop results with ±10% changes of this flow rate. Note that flow rates of the two main products and water are all increased or decreased according to the demand of throughput changes. Since the throughput change is assumed to be a known disturbance, small deviations of the product compositions can be eliminated by minor adjustments of the set point values of three tray temperature loops.

compositions are still maintained at high-purity despite reasonable variations of feed and vapor split disturbances.

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*Tel.: +886-2-3366-3063. Fax: +886-2-2362-3040. E-mail: [email protected]. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS Financial supports from the National Science Council of the R.O.C. under grant no. NSC 100-2221-E-002-115-MY3 is greatly appreciated.

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REFERENCES

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6. CONCLUSIONS This study investigated the feasible design for esterification of AmOH and HexOH mixtures with acetic acid. Two design configurations are explored in this work, including “directsequence” and “indirect-sequence” designs. The results show that the direct-sequence design with the first RD column to drawoff the water byproduct from the top decanter and another column to further separate the two esters from the RD bottoms gives significant energy and TAC savings as compared with an indirect-sequence design. By further combining two reboilers into one in a thermally coupled design, further 13.16% savings in steam cost and 12.56% savings in TAC can be realized. The proposed direct-sequence thermally coupled design flowsheet is new for this Type-III mixed-alcohol RD system. As for overall control strategy of the proposed thermally coupled RD process, all three remaining manipulated variables besides the inventory control loops have to be used in tray temperature loops. The more difficult choice is to select a control point in RD column to have enough open-loop sensitivity to feed ratio but also give the least temperature deviation in closed-loop sensitivity analyses. Compromise has to be made which results in some sacrifice of closed-loop performance. The product 17197


Design and Control of Thermally-Coupled Reactive Distillation System

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