Investigation of the Energy-Saving Design of an Industrial 1,4

Article pubs.acs.org/IECR

Investigation of the Energy-Saving Design of an Industrial 1,4-Dioxane Dehydration Process with Light Feed Impurity Yi Chang Wu, Hsiao-Ping Huang, and I-Lung Chien* Department of Chemical Engineering, National Taiwan University, Taipei 10617, Taiwan ABSTRACT: In this work, several energy-saving designs of an industrial 1,4-dioxane dehydration process with small amounts of triethylamine (TEA) impurity were compared. The original process contains a preconcentrator column to remove TEA and two other columns operated at different pressures by pressure-swing distillation to separate the remaining azeotrope-containing mixture. To save on the operating costs of this process, two alternative designs were investigated. The first combines the condenser of a high-pressure column with the reboiler of a low-pressure column. The second design uses a completely different method for this separation by heterogeneous azeotropic distillation. This second design can further be thermally coupled into a dividing-wall column with only one column shell. It was found that significant savings in the operating costs (49.09%) and in the total annual cost (43.31%) can be realized by using this second design as compared to the original process. The proposed method uses the light impurity as an entrainer, so no foreign component is added into the system. The operation and control of the proposed design were also studied. An overall conventional tray-temperature control strategy was proposed to maintain highpurity product streams despite disturbances in feed composition and throughput.

1. INTRODUCTION Distillation is the most widely used separation process in the chemical industry. In one work,1 it was estimated that separation processes account for 40−70% of both capital and operating costs in petrochemical processing and that distillation is used to perform 90−95% of all separations in the chemical process industry. In another article,2 it was stated that distillation columns and their support facilities can account for about one-third of the total capital costs and more than half of the total energy consumption. Consequently, the design and control of the distillation train has a critical impact on the economics of the entire process. Among all distillation processes, mixtures containing azeotropes are most difficult to separate. Simple distillation cannot be used to achieve complete separation. A recent book by Luyben and Chien3 summarizes the feasible techniques used in industry to achieve such separations. For any particular azeotropic mixture, there are multiple ways to achieve the separation task. However, significant savings on the steam cost and total annual cost can be achieved if the most suitable separation method can be selected. Quite a few illustrative examples have been published in the open literature. For the separation of C5s and methanol in a tert-amyl methyl ether (TAME) reactive-distillation process, Luyben4 compared the separation methods of pressure-swing distillation and extractive distillation. In a follow-up article by Wu et al.5 on the same separation, it was demonstrated that, by using the separation method for binary heterogeneous azeotropes of a twostripper/decanter flowsheet, the total annual cost can be decreased by a factor of 4 compared to that of the corresponding pressure-swing system. As another example, Arifin and Chien6 studied a heterogeneous azeotropic system for separating isopropanol and water using cyclohexane as the light entrainer. However, this separation method resulted in a significantly higher energy requirement than the competing separation method of extractive distillation.7 Kiss and Suszwalak8 compared © 2014 American Chemical Society

the dehydration separation of bioethanol by heterogeneous azeotropic distillation and extractive distillation. Similar results showing extractive distillation to be more economical were obtained. Other comparative examples of pressure-swing distillation or extractive distillation include separations of acetone/ methanol,9 acetone/chloroform,10 and methylal/methanol11−14 and recovery of organic solvents from an aqueous waste mixture.15 In this work, an industrial solvent recovery system for the separation of 1,4-dioxane and water with small amounts of impurity (triethylamine, TEA) was studied to determine the most efficient separation method for this mixture. Only a few works in the open literature have studied possible separation methods for this mixture system,16−18 specifically by using pervaporation membranes or adding another component to form a buffer solution. However, those works did not provide a complete design flowsheet, let alone identify the optimal design and further study the operation and control of this separation system. Another article19 in the open literature used triethylamine as an intermediate entrainer to aid the separation; however, the mixture to be separated was methanol/toluene, which is not the same as the system studied here. In this article, the optimal design using pressure-swing distillation (PSD) for our separation system is compared with a heterogeneous azeotropic distillation (HAD) system using the feed impurity component as the entrainer. To further save energy in the azeotropic separation process, various heat-integration schemes were also explored for this separation system. One approach is to operate two columns in the process at different pressures so that the condenser of the Received: Revised: Accepted: Published: 15667

May 4, 2014 September 11, 2014 September 16, 2014 September 16, 2014 dx.doi.org/10.1021/ie501831f | Ind. Eng. Chem. Res. 2014, 53, 15667−15685

Industrial & Engineering Chemistry Research

Article

Figure 1. T−x−y plots and conceptual design flowsheet of the pressure-swing distillation process.

high-pressure column can be combined with the reboiler of the low-pressure column to save energy. This energy-saving design can naturally be applied to pressure-swing distillation.12,20,21 Another feasible approach is to devise a thermally coupled (dividing-wall) design to save energy. A good complete reference for the subject of dividing-wall columns was published by Yildirim et al.22 They provide a comprehensive review of current industrial applications of dividing-wall columns and related research activities including column configurations, design, modeling, and control issues. Wu et al.23 proposed a heterogeneous azeotropic dividing-wall column for the separation of pyridine and water using toluene as the entrainer. We investigated this process intensification method for the 1,4-dioxane/water separation system using heterogeneous azeotropic distillation.

Table 1. NRTL Model Parameters for This Separation System component i/component j source aij aji bij (K) bji (K) cij

1,4-dioxane/water

1,4-dioxane/TEA

TEA/water

ASPEN VLE-IG −3.3099 6.5419 1348.1772 −1699.4196 0.3

ASPEN VLE-IG 0 0 522.6913 −127.3588 0.3

ASPEN LLE-IG 7.5958 14.7698 −1966.3192 −3310.1699 0.3

two-column sequence. The Aspen Plus built-in NRTL thermodynamic model parameters for predicting the T−x−y plots are listed in Table 1. However, because of the small amounts of light component (TEA) in the feed stream, the design with only a two-column sequence will not work. The reason for this failure is the accumulation of TEA in the top recycle loop of the two-column sequence. A preconcentrator column is necessary to prevent the impurity from going into the two-column sequence. From the residue curve map (RCM) and liquid−liquid envelope (LLE) for this ternary system at 1 atm (see Figure 2), a small loss of 1,4-dioxane is inevitable through the top stream of the preconcentrator column. The feed flow rate was 50 kmol/h (1096.04 kg/h) with a feed composition of 30 wt % 1,4-dioxane, 69 wt % water, and 1 wt % TEA. The preconcentrator column was a stripper with

2. BASE-CASE DESIGN FLOWSHEET USING PRESSURE-SWING DISTILLATION The conceptual design flowsheet of this industrial separation system is illustrated in Figure 1. The feed stream contains mostly aqueous 1,4-dioxane with small amounts of the light impurity triethylamine (TEA). As shown in the top part of Figure 1, 1,4-dioxane and water form a minimum-boiling azetrope. As can be seen in Figure 1, the azetropic composition is 19.54 wt % water at 1 atm. This azetropic composition shifts to 31.31 wt % water at 8 atm. Thus, with a low-pressure column operated at 1 atm and a high-pressure column operated at 8 atm, these two components can easily be separated using a 15668

dx.doi.org/10.1021/ie501831f | Ind. Eng. Chem. Res. 2014, 53, 15667−15685


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is a sample plot of several results from various simulations obtained by fixing NT3 = 11 and NF3 = 4 and finding the best values of NT2 and NF2 to minimize the TAC of the low-pressure column. Figure 4d is a sample plot of several results from various simulations obtained by fixing NT2 = 14 and NF2 = 7 and finding the best values of NT3 and NF3 to minimize the TAC of the lowpressure column. By collecting many results from such plots, the best NT2, NF2, NT3, and NF3 values that minimize the overall TAC of this process can be obtained. For the pressure-swing distillation system, two operating pressures can also be design variables. In this study, the operating pressure for the low-pressure column was set to atmospheric pressure because cooling water can still be used in the condenser. For the high-pressure column, the operating pressure was set to 8 atm because the difference in the two azeotropic compositions becomes flatter at higher pressures. The TAC calculations include annual total operating cost and total installed capital cost divided by the payback period. Operating costs include the costs of steam and cooling water and also a stream cost accounting for the dioxane lost from the top of the preconcentrator column. The capital costs include the costs of columns, reboilers, and condensers. The payback period was assumed to be 3 years, and the wastewater treatment cost was assumed to be negligible in the overall TAC. The formulas for the installed capital costs for all of the process equipment can be found in Chapter 5, Economic Basis, of Luyben’s book Principles and Case Studies of Simultaneous Design.24 The unit prices of $7.78/GJ for low-pressure steam (used at the preconcentrator column and at the low-pressure column) and $9.88/GJ for high-pressure steam (used at the high-pressure column) were also from the same book. The unit price for calculating the 1,4-dioxane loss was $1.95/kg. Note from the flowsheet in Figure 3 that the ratio of the recirculation rate (D3) to the feed of the two-column sequence (B1) was only 0.584. It was not obvious that this was an economically unfavored design flowsheet before this investigation. We used this flowsheet as a base case to compare with other alternative designs proposed in this article.

Figure 2. RCM and LLE of the studied system (weight-based).

only one control degree of freedom (DOF). This DOF was used to set the bottom composition at 1 × 10−7 wt % TEA to avoid the accumulation problem in the two-column sequence. Both the low-pressure column and the high-pressure column had two degrees of freedom. One of the degrees of freedom for each column was used to set the product purities of water and 1,4-dioxane all at 99.9 wt %. The other degrees of freedom (two reflux ratios, RR2 and RR3) were varied to minimize the total reboiler duty of the two-column sequence. The design flowsheet of the industrial separation process is summarized in Figure 3. The flowsheet was established by minimizing the total annual cost (TAC) of this process. The design variable of the preconcentrator column is its total number of stages (NT1). The design variables of the twocolumn sequence are the total numbers of stages for these two columns (NT2 and NT3) and also the two feed locations (NF2 and NF3). The summarized plots for determining the flowsheet are displayed in Figure 4. Figure 4a shows that NT1 should be 7 to minimize the TAC of this preconcentrator column. Figure 4b is a sample plot demonstrating how to find RR2 and RR3 with the set of values of NT2, NF2, NT3, and NF3. Figure 4c

3. MULTIEFFECT PRESSURE-SWING DISTILLATION Because the previous design requires two columns operated at different pressures, it is quite natural to consider multieffect

Figure 3. Design flowsheet of the pressure-swing distillation process. 15669



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Figure 4. Plots for determining the design and operating variables of the pressure-swing distillation process.

with Qr2. Figure 5 summarizes the simulation results, showing that the optimum RR2 and RR3 values are 0.40 and 0.776, respectively. At these values, the corresponding net reboiler duty (Qr3) is the lowest value of 304.0 kW (net required reboiler duty for the two-column sequence). Figure 6 shows the multieffect design flowsheet with complete heat integration. Note that, in this figure, the reboiler of the low-pressure column and the condenser of the high-pressure column seem to be two separate process units. However, in actual implementation, these two process units are instead one process-to-process heat exchanger. A comparison of the total operating costs and TACs of the designs in Figures 3 and 6 is provided in Table 2. The more detailed itemized TAC terms can be found in Table 3. In developing the multieffect design, the total numbers of stages for the high- and low-pressure columns were all fixed as in the original design flowsheet in Figure 3 to clearly show the benefit of the savings in reboiler duty using this multieffect design. According to Table 2, significant savings in the total reboiler duty (−35.97%) can be obtained by using the multieffect design flowsheet with complete heat integration. The TAC can also be reduced by 16.22% with the multieffect design in Figure 6.

designs to save reboiler duty by combining the condenser of the high-pressure column with the reboiler of the low-pressure column. According to Figure 3, the temperature at the top of the high-pressure column is 433.2 K, and that at the bottom of the low-pressure column is 375.5 K. The temperature difference (57.7 K) is more than enough for the heat-transfer requirement. Therefore, a partial heat-integration design can be used to save portions of the reboiler duty at the low-pressure column (savings of 173.4 kW). An auxiliary reboiler is needed at the low-pressure column to supply the discrepancy (duty of 275.4 − 173.4 = 102.0 kW). Another way to save reboiler duty of the high-pressure column is to add a feed−effluent heat exchanger to transfer some heat from the bottom of this column (T = 461.5 K in Figure 3) to the feed stream of this column (T = 361.0 K). In the following simulation studies, a countercurrent heat exchanger with a minimum approach temperature of 10 K was assumed. To further save total reboiler duty, a complete heatintegration design can be used to alter the two reflux ratios in the design flowsheet so that the condenser duty of the highpressure column (Qc3) can be exactly matched (with opposite sign) by the reboiler duty of the low-pressure column (Qr2). For each RR2 value, there is an RR3 value to exactly match Qc3 15670



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concepts (feed−effluent heat exchanger and complete multieffect heat intergration) are implemented in this design flowsheet. An intermediate design obtained from the base case by adding a feed−effluent heat exchanger and then employing partial heat integration by combining the condenser of the high-pressure column with the reboiler of the low-pressure column to save portions of the reboiler duty at the low-pressure column can be made. This alternative design, although it cannot save the most reboiler duty, should be preferable in terms of preserving the important control degree of freedom (auxiliary reboiler) at the low-pressure column. In this study, our purpose was to develop a design flowsheet that can save the most reboiler duty using pressure-swing distillation and then compare it with the use of other separation methods.

4. HETEROGENEOUS AZEOTROPIC DISTILLATION SYSTEM According to the ternary diagram in Figure 2, the feed impurity forms a binary azeotrope with water having the lowest azeotropic temperature (75.25 °C) of this ternary system. Also, this binary azeotrope is within the liquid−liquid boundary, which means that this azeotrope can naturally be separated in a decanter into an aqueous phase and another organic phase rich in TEA. However, another important observation is that it is very difficult to reach this binary heterogeneous azeotrope in the top vapor stream of a heterogeneous azeotropic column by going through a very narrow funnel. 4.1. Conventional Two-Column Design. Wu and Chien25 studied a pyridine/water separation system using heterogeneous azeotropic distillation with toluene as the entrainer.

Figure 5. Plots for determining RR2 and RR3 of the multieffect PSD with complete heat integration.

The flowsheet in Figure 6 is the design that can save the most reboiler duty from the base case because two energy-saving

Figure 6. Design flowsheet of the multieffect PSD with complete heat integration.

Table 2. Comparison of Various Designs for This Separation System dioxane loss (kg/h)

total operating cost (1000 $/year)

Q (kW)

base case multieffect PSD with complete heat integration

5.04 5.04

Pressure-Swing Distillation 580.7 (0%) 229.213 (0%) 371.8 (−35.97%) 184.291 (−19.60%)

conventional two-column stripping-column design thermally coupled design dividing-wall column design

2.94 3.40 2.49 2.11

Heterogeneous Azeotropic Distillation 514.8 (−11.35%) 166.129 (−27.52%) 470.8 (−18.93%) 163.578 (−28.63%) 375.4 (−35.36%) 126.497 (−44.81%) 359.7 (−38.06%) 116.682 (−49.09%) 15671

total installed capital cost (1000 $)

TAC ($1000/year)

167.440 (0%) 163.542 (−2.33%)

285.026 (0%) 238.805 (−16.22%)

160.142 (−4.36%) 156.321 (−6.64%) 138.104 (−17.52%) 134.663 (−19.57%)

219.510 (−22.99%) 215.685 (−24.33%) 172.532 (−39.47%) 161.570 (−43.31%)



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Table 3. Itemized TAC Terms for the Base Case and Multieffect PSD Design with Complete Heat Integration Design multieffect PSD design with complete heat integration

pressure-swing distillation C1 installed capital cost for column (1000 $) installed capital cost for reboiler (1000 $) installed capital cost for condenser (1000 $) installed capital cost for heat exchangers (1000 $) steam cost (1000 $/year) cooling-water cost (1000 $/year) dioxane loss (1000 $/year) total reboiler duty (kW) (difference) total steam cost (1000 $/year) (difference) TAC (1000 $/year) (difference)

9.037 10.603 − − 15.473 − 78.624 580.7 (0%) 147.078 (0%) 285.026 (0%)

CL

CH

C1

CL

CH

35.454 29.173 26.225 − 62.815 2.240

22.461 24.348 10.139 − 68.790 1.271

9.037 10.603 − − 15.473 −

34.370 − 25.363 6.006 − 2.128

25.626 28.589 − 23.948 88.066 −

371.8 (−35.97%) 103.539 (−29.60%) 238.805 (−16.22%)

Figure 7. Conceptual design flowsheet and material balance lines of the heterogeneous azeotropic distillation system. 15672



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Figure 8. Plots for determining the design and operating variables of the heterogeneous azeotropic distillation system.

Figure 9. Design flowsheet of the heterogeneous azeotropic distillation system.

The RCM type was similar to that of this system, with three distillation regions and three binary azeotropes and with the azeotropic temperature of the heterogeneous azeotrope (TEA/ water) having the minimum temperature of the system. We used the idea from Wu and Chien25 to develop a feasible design flowsheet using heterogeneous azeotropic distillation. There is one important difference between the flowsheet in Wu and Chien25 and that for our system. In the work of Wu and Chien,25 the entrainer (toluene) was a new component added to the system. However, in our system, the candidate entrainer (TEA) is not a foreign component. Not adding a foreign component to the system is preferable in industry. For this

proposed design, a purge stream needs to be designed to prevent TEA accumulation inside our system. Figure 7 shows the material balance lines and the conceptual design flowsheet for this dioxane dehydration system. In this two-column design, one column serves as a preconcentrator column for the feed and as a recovery column for the aqueous outlet stream from a decanter. The top vapor of a heterogeneous azeotropic column does not strictly need to go through a very narrow funnel to the minimum binary azeotrope of TEA/water, but can be at location V1 in the top plot in Figure 7. After this top vapor V1 has been condensed to liquid, aqueous and organic phases can be formed, with the organic 15673



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Table 4. Itemized TAC Terms for the Four Heterogeneous Azeotropic Designs conventional twocolumn system C-1 installed capital cost for column (1000 $) installed capital cost for reboiler (1000 $) installed capital cost for condenser (1000 $) installed capital cost for compressor (1000 $) steam cost (1000 $/year) cooling-water cost (1000 $/year) electricity cost (1000 $/year) dioxane loss (1000 $/year) total reboiler duty (kW) (difference) total steam cost (1000 $/year) (difference) TAC (1000 $/year) (difference)

C-2

33.355 28.506 29.929 26.656 25.201 16.495 − − 63.039 54.385 1.730 1.111 − − 45.864 514.8 (0%) 117.424 (0%) 219.510 (0%)

stripping-column design C-1

C-2

35.080 29.509 25.832 27.428 33.515 − − 4.957 50.575 56.828 2.906 − − 0.224 53.045 470.87 (−8.53%) 107.403 (−8.53%) 215.685 (−1.74%)

thermally coupled design

dividing-wall column design

C-1 + C-2

C-1

C-3

32.895 28.245 18.931 27.023 26.102 − − 4.908 31.081 54.549 1.809 − − 0.214 38.844 375.4 (−27.08%) 85.63 (−27.08%) 172.532 (−21.40%)

C-2 + C-3

23.504 41.267 17.554 27.002 25.336 − − − 27.573 54.481 1.712 − − − 32.916 359.7 (−30.12%) 82.054 (−30.12%) 161.570 (−26.39%)

Figure 10. Design flowsheet of the heterogeneous azeotropic distillation system with stripping-column design.

the number of stages in the other column (high-pressure column) in the pressure-swing design flowsheet. The reason for this assumption was to eliminate the trade-off from savings of reboiler duty when increasing the total number of stages. In other words, if the total reboiler duty requirement of this HAD system can be reduced even when arbitrarily setting the total numbers of stages to be the same as in the PSD system, both the capital and operating costs of the HAD system should be more favorable. Four design and operating variables need to be determined for this design flowsheet. They are the feed location of the HAD column (NF1), the feed location of the combined column (NF2), the organic reflux fraction (ORF), and the reflux ratio of the combined colum (RR). The reboiler duty of the HAD column was used to fix the dioxane purity at 99.9 wt %, and the

phase mostly refluxing back to the heterogeneous azeotropic column and a small amount purging out of the system to balance TEA from the feed stream. The aqueous phase is designed to be combined with fresh feed stream entering the combined column for further purification into water product. The bottom product of the heterogeneous azeotropic column can be pure dioxane with specified purity because dioxane has the maximum temperature in the same distillation region as V1. For simplicity in investigating of the energy-saving design for this separation system, the total number of stages for the combined column was set to be the same as the number of stages in the low-pressure column in the optimized PSD design flowsheet because the purposes of these two columns in obtaining purifying water product are similar. The total number of stages for the heterogeneous column was set to be the same as 15674



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Figure 11. Design flowsheet of the heterogeneous azeotropic distillation system with thermally coupled column design.

reboiler duty of the combined column was used to fix the water purity also at 99.9 wt %. Many simulation runs were evaluated with varying ORF and RR to determine the optimal NF1 and NF2 values to allow the total reboiler duty to reach its minimum value or with varying NF1 and NF2 to find the optimal ORF and RR values to achieve the same goal. Figure 8 shows samples of many simulation runs with either varying ORF and RR or varying NF1 and NF2. In Figure 8b for NF1 = 6 and NF2 = 6, ORF cannot be set below 0.99 because of an inability to maintain product purities. Figure 9 displays the design flowsheet for this conventional two-column design. Note that, in this figure, NT1 = 10 actually has the same number of stages as for the highpressure column in Figure 3. Because the condenser and decanter are simulated outside the heterogeneous azeotropic column, stage 1 is defined here as the top tray, not as the condenser in Figure 3. Table 2 also summarizes the main economic data for this conventional two-column design using HAD, with the more detailed itemized TAC terms found in Table 4. Note that the total operating cost and total installed capital cost are even less than for the multieffect design with complete heat integration.

The reasons for the lower total operating cost are that the dioxane losses are lower and also that both reboilers require only low-pressure steam. According to Table 2, compared with the base case, the total operating cost can be further reduced by 27.52%, with the TAC also further reduced by 22.99%. 4.2. Stripping-Column Design. An alternative design to save additional energy using a heterogeneous azeotropic distillation system was proposed by Chang et al.,26 who incorporated a stripping column in their three-column system for isopropyl alcohol dehydration. Their idea from the perspective of energy conservation was that a stripping column is a tower without a condenser, in which energy carried by the overhead vapor is conserved instead of being removed in a condenser, so that it is more energy-efficient than the conventional recovery column. This design concept was adapted here for the dioxane dehydration system. The resulting flowsheet with the same feed conditions and product purity specifications is shown in Figure 10. Note that the feed location for this stripping column has to be the first stage to provide liquid traffic down the column. A small compressor is needed to transport overhead vapor from the stripping column to the feed location of the HAD column. 15675



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Figure 12. Plots for determining the design and operating variables of the heterogeneous azeotropic distillation system with thermally coupled column design.

column will also be higher. The flowsheet in Figure 10 shows exactly the same observations as predicted. For the reasons mentioned above, the reboiler duty of the heterogeneous column could only be reduced from 276.4 to 221.7 kW, even though the energy carried by the overhead vapor of the stripping column was conserved instead of being removed in a condenser for a regular column. This result is qualitatively different from the result found by Wu et al.23 for the pyridine/water system. In that work,23 a significant reduction of the reboiler duty for the heterogeneous column was achieved using the stripping-column design. The result can be predicted because, in our case (Figure 7), the distillation boundary is farther from the FF + AQ point. A regular column will be preferred to purify the top product so that it is richer in dioxane instead of a stripping-column design with only one equilibrium stage for the top product. Table 2 also summarizes

In the conventional two-column system, the distillate composition of the combined column was optimized so that the dioxane content could be as high at 76.73 wt %. However, for the stripping-column configuration, the design degree of freedom at the top of the conventional recovery column is lost. Thus, the top vapor composition for this stripping column was only 71.00 wt % dioxane. By observing the material balance lines in the ternary diagram of Figure 7, the location of D2 for the strippingcolumn design will be at a lower point with less dioxane. From the intersection of the two material balance lines of D2−OR and dioxane−V1 in the top plot in Figure 7, it can be expected from the lever rule that the top vapor flow rate of this strippingcolumn design will be larger than that of the original design in Figure 9. With a higher top vapor flow rate, the organic reflux flow rate and the aqueous outlet flow rate to the stripping 15676



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Figure 13. Design flowsheet of the heterogeneous azeotropic distillation system with dividing-wall column design.

respectively. Note that, in Figure 11, there are three column sections. The upper part of the heterogeneous column is denoted as C-1, and the lower part is denoted as C-2. The second column is denoted as C-3. The reason for this naming convention is because the diameters of upper and lower parts of the heterogeneous column are different. Because of the higher vapor traffic for the upper part, its diameter is larger. From the design flowsheet in Figure 11, it can be observed that the top vapor flow rate of the heterogeneous column is significantly reduced in comparison with that for the strippingcolumn design in Figure 10. Because the energy carried by the overhead vapor of column C-3 is still conserved in this thermally coupled design, the reboiler duty of the heterogeneous column (136.3 kW) is significantly reduced in comparison with those in the original heterogeneous design in Figure 9 (276.4 kW) and in the stripping-column design in Figure 10 (221.7 kW). Table 2 also summarizes the main economic data for this thermally coupled design, with the more detailed itemized TAC terms found in Table 4. 4.4. Dividing-Wall Column Design. To have the additional benefit of combining two columns into a single column shell to save on space requirements, the number of stages for the heterogeneous azeotropic column below the liquid sidedraw stage should be the same as the total number of stages in column C-3. In this way, a dividing-wall column as in the flowsheet in Figure 13 with a dividing wall at the lower part of a single column can be designed. Because there are only four stages below the liquid side-draw stage for the thermally coupled design (Figure 11), an additional nine stages on the left side of the dividing-wall column are needed to complete the design. An equivalent diameter for the lower part of the dividing-wall column can be calculated as in Wu et al.23 The idea is to have the overall cross-sectional area be equal to

the main economic data for this stripping-column design, with the more detailed TAC terms itemized in Table 4. 4.3. Thermally Coupled Design. An alternative way to have a top product richer in dioxane is to still have the freedom of lowering the feed location of the combined feed so that more stages can be allowed above the feed stream. However, with this concept, a liquid side draw from the heterogeneous column should be designed to provide liquid traffic in the upper part of the second column. As in the stripping-column design, the overhead vapor of the second column can be fed into the heterogeneous column at the same side-draw location. To transport this overhead vapor, a small compressor is needed in the design flowsheet. The flowsheet of this thermally coupled design can be seen in Figure 11. Again, for simplicity in investigating the energysaving design for this separation system, the total numbers of stages of the two columns were fixed at the same value as in Figure 9. Three design and operating variables need to be determined for this design flowsheet: the liquid side-draw flow rate (L2), the side-draw and feed location of the HAD column (NF1), and the feed location of the combined column (NF3). The reboiler duty of the HAD column was used to fix the dioxane purity at 99.9 wt %, and the reboiler duty of the combined column was used to fix the water purity also at 99.9 wt %. The ORF was also fixed as in Figure 9. An iterative sequential optimization procedure was used to determine the optimal values of these three variables to minimize the total reboiler duty. The liquid side-draw flow rate (L2) was designed as the variable for the outermost iterative loop. For a given L2 value, optimal values of NF1 and NF3 were obtained. Figure 12 summarizes the results of many simulation runs. The optimal L2 value was found to be 3.0 kmol/h, and the optimal NF1 and NF3 values were the sixth and fifth stages, 15677



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Again, the ORF was also fixed as in Figure 9. Figure 14 summarizes simulation runs used to determine these two variables to minimize the total reboiler duty. The results show that L2 should be 3.5 kmol/h, and NF3 should be the fifth stage. Table 2 also summarizes the main economic data for this dividing-wall design, with the more detailed itemized TAC terms found in Table 4.

5. COMPARISON OF ALTERNATIVE DESIGNS In this section, the reboiler duties, total operating costs, total installed capital costs, and TACs of all of the cases investigated in the preceding sections are compared with the base-case design in Figure 3. Total operating costs include steam costs, cooling-water costs, and stream costs representing the dioxane loss. The total installed capital cost indicates the overall capital investment for this separation system. TAC is defined as the total operating costs plus the total installed capital cost divided by the payback period. The payback period was assumed to be 3 years in this study. The results are summarized in Table 2. More detailed itemized TAC terms can be found in Table 3 for the base case and multieffect PSD with complete heat integration and in Table 4 for the HAD system, stripping-column design, thermally coupled design, and dividing-wall column design. By combining the condenser of the high-pressure column and the reboiler of the low-pressure column and adding another feed−effluent heat exchanger to reduce the reboiler duty at the high-pressure column (Figure 6), a significant 35.97% reduction

Figure 14. Plots for determining L2 and NF3 of the heterogeneous azeotropic distillation system with dividing-wall column design.

the combination of two cross-sectional areas on the two sides of the dividing wall. Again, for simplicity in investigating the energy-saving design for this separation system, the total numbers of stages in columns C-1 and C-3 were fixed at the same values as in Figure 11. Only two remaining design and operating variables need to be determined for this design flowsheet: the liquid side-draw flow rate (L2) and the feed location of the combined column (NF3).

Figure 15. Overall control strategy of the conventional heterogeneous azeotropic distillation system. 15678



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Figure 16. Overall control strategy of the dividing-wall column system.

expected. The original idea by Chang et al.26 was that a stripping column is a tower without a condenser, in which energy carried by the overhead vapor is conserved instead of being removed in a condenser, so that it can be considered more energy-efficient than a conventional recovery column. However, because the stripping column has only one equilibrium stage for the feed stream to purify into top product, the residue curve maps and liquid−liquid envelope of the studied system can be used to predict whether significant savings of the reboiler duty can be achieved. The full explanation can be found in section 4.2. For this studied system, not much of a savings can be realized by using the design in Figure 10. Significant savings in both operating and capital costs can be obtained by using either the thermally coupled design in Figure 11 or the dividing-wall column design in Figure 13. The idea of energy conservation for the stripping-column design can be preserved by using a thermally coupled design without the deficiency of not being able to purify the top product of the second column. From the results, significant savings in both the total operating costs (44.81%) and the capital cost (17.52%) in comparison with the base case can be realized by using a thermally coupled design. By further combining two thermally

of the total reboiler duty can be realized. However, because of savings on only the lower-grade steam and also of larger dioxane stream cost, the savings of the total operating cost are reduced to 19.60%. As for the capital cost, because of savings of a condenser and another reboiler but addition of two heat exchangers, there are only negligible savings (2.33%) on the total capital cost. The overall savings on the TAC are still considerable at 16.22%. By using the different separation method of heterogeneous azeotropic distillation as in Figure 9, the total number of columns can be reduced from three to two. Although the total reboiler duty is higher than the case using PSD with complete heat integration, the total operating cost is even lower than for the design in Figure 6. This is because of lower dioxane losss and also the need for just lower-grade steam for the twocolumn system. The overall TAC of the case in Figure 9 is also lower than that of the case in Figure 6. These results demonstrate that, by selecting a more appropriate separation method, a simpler conventional design can even be better than a more complicated heat-integration design of another separation method. This work also demonstrates that the stripping-column design might not save as much of the total reboiler duty as first 15679



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Figure 17. Closed-loop responses of the conventional heterogeneous azeotropic distillation system with ±20% changes in feed composition.

6. CONTROL OF A DIVIDING-WALL COLUMN SYSTEM In this section, the dynamics and control of the proposed design flowsheet in Figure 13 are examined to determine whether any sacrifice needs to be made in comparison with the conventional two-column system in Figure 9. We did not investigate the control performance of the multieffect PSD with complete heat integration in Figure 6. This type of complete heat-integration design is known to have worse control behavior than a partial heat-integration scheme.20 Moreover, this design gives a much inferior economic performance in comparison with the proposed design in Figure 13. Only the conventional control structure using multiloop tray-temperature control was implemented in the following control study. The overall control strategy for the conventional two-column system was adapted from the study of Wu et al.23 The only difference is that there was no entrainer makeup stream in our study. Instead, a purge stream was needed to avoid accumulation

coupled columns into one column shell in a dividing-wall column, additional savings in the total operating costs (49.09%) and also in the capital cost (19.57%) in comparison with the base case can be achieved. The savings in overall TAC for the dividing-wall column design are also very significant at 43.31%. In the calculation of the column cost for the dividing-wall column with only one column shell, the additional costs for more complex tray internals including the partition wall inside the column were neglected. Dejanović et al.27 provided more detail calculations to account for this more complex column design. Even with a design complexity factor of 1.2 as suggested in Dejanović et al.27 for the lower part (C-2 + C-3) of the dividing-wall column, the dividing-wall column design is still most economical, not to mention its additional benefits of lower space requirements at the plant site and less instrumentation and control equipment. 15680



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Figure 18. Closed-loop responses of the conventional heterogeneous azeotropic distillation system with ±20% changes in feed flow rate.

dynamic study of dividing-wall column systems. In reality, the vapor flow will naturally move from the right side of the dividing wall (section C-3) to the upper section (section C-1) as a result of the pressure drop. Note also that, in this dividing-wall column configuration, the original feed to column C-1 is now an internal vapor flow from section C-3 (right side of dividing wall) to section C-1. This internal flow is realistically considered as unmeasurable, so that the ratio of the organic flux to the column C-1 feed cannot be fixed. Therefore, the level of the organic phase in the dividingwall column system is set to be controlled by the organic reflux flow so that this flow can be varied under disturbances. The purge flow is then set to maintain it at a constant ratio with the fresh feed. Similar open-loop sensitivity analyses were performed to determine the two tray-temperature control points. The closed-loop dynamic responses under various disturbances for the conventional two-column system and for the dividingwall column system are discussed next.

of TEA within the process system. The overall control strategy for the conventional two-column system is summarized in Figure 15. The purge flow (an external flow) was used to maintain the organic level, whereas the organic reflux flow was fixed to a constant ratio with respect to the column C-1 feed. For column C-2, the reflux ratio was fixed at the base-case value. The reboiler duties of columns C-1 and C-2 were used to hold the temperature control point inside columns C-1 and C-2, respectively. The control points were determined by open-loop sensitivity analyses with one-at-a-time perturbation of the reboiler duties. The most sensitive and nearly linear stages in columns C-1 and C-2 were selected as the control points. The overall control strategy for the dividing-wall system is summarized in Figure 16. Note that there is a fictitious compressor at the top of section C-3 to ensure that vapor traffic can correctly be flowed from section C-3 (right side of dividing wall) to section C-1. This type of arrangement is commonly used in the open literature (e.g., Ling and Luyben28) for the 15681



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Figure 19. Closed-loop responses of the dividing-wall column system with ±20% changes in feed composition.

observed for the throughput changes. These control performances were compared with the closed-loop results for the dividing-wall column system under same disturbances. The closed-loop dynamic simulations with ±20% changes in the feed composition for the dividing-wall column system are shown in Figure 19. Both temperature control points were able to quickly return to their set points. The two product flow rates (B2 and B3) increased/decreased according to the unmeasured feed composition changes. Both products were maintained at high purity despite these unmeasured disturbances. The only

The control performances for disturbances in the feed composition and feed flow rate for the conventional two-column system are shown in Figures 17 and 18, respectively. These closed-loop results were obtained using Aspen Plus Dynamics by pressure-driven simulation. For the unmeasured disturbances in feed composition in Figure 17, it was found that the two temperature control points quickly returned to their set-point values. More importantly, both products could be maintained at high purities despite these large ±20% unmeasured disturbance changes. Similarly good closed-loop dynamic results were 15682



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Figure 20. Closed-loop responses of the dividing-wall column system with ±20% changes in feed flow rate.

in the feed flow rate. Both temperature control points were able to quickly return to their set points. Note that both products were maintained at high purity with corresponding changes in the production rates. The third disturbance considered was the liquid split ratio. We assumed that the actual ratio differed from the designed value by either +10% or −10%. The closed-loop dynamic simulations with these two unmeasured disturbances are shown in Figure 21. Again, the control performances were highly satisfactory, with both products still maintained at high purity.

minor problem was in the dynamic transient response for the dioxane product purity. The dioxane purity dipped to 0.9935 (still at high purity) during the −20% change in feed dioxane composition but quickly returned to a higher purity of over 0.999. The final purity deviations from the corresponding basecase values were even smaller than for the conventional twocolumn system. The second disturbance considered was throughput changes, which were achieved by changing the feed flow rate. Figure 20 displays the closed-loop dynamic simulations with ±20% changes 15683



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Figure 21. Closed-loop responses of the dividing-wall column system with ±10% changes in liquid split ratio.

7. CONCLUSIONS In this article, an industrial 1,4-dioxane/water separation system was used as an example to demonstrate two commonly used methods for separating azeotropes. Separation by pressureswing distillation requires the use of three columns to achieve the separation. Specifically, a preconcentrator column is necessary to avoid the accumulation of the light feed impurity within the pressure-swing distillation system. For the completely different separation method of heterogeneous azeotropic distillation, only a two-column system is needed because this light feed impurity can be used as the entrainer for the separation.

Further heat integration of both the pressure-swing distillation system and the heterogeneous azeotropic distillation system was thoroughly investigated. Although a 19.60% reduction in total operating costs and a 16.22% reduction in total annual cost can be achieved by using a design flowsheet of multieffect PSD with complete heat integration, this type of complete heatintegration design is known to have worse control behavior than the original system because of the loss of an important control degree of freedom. Using a more appropriate separation method for this example, namely, heterogeneous azeotropic distillation, can already cut the total operating costs by 27.52% 15684



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(8) Kiss, A. A.; Suszwalak, D. J.-P. C. Enhanced Bioethanol Dehydration by Extractive and Azeotropic Distillation in DividingWall Columns. Sep. Purif. Technol. 2012, 86, 70. (9) Luyben, W. L. Comparison of Extractive Distillation and Pressure-Swing Distillation for Acetone−Methanol Separation. Ind. Eng. Chem. Res. 2008, 47, 2696. (10) Luyben, W. L. Comparison of Extractive Distillation and Pressure-Swing Distillation for Acetone/Chloroform Separation. Comput. Chem. Eng. 2013, 50, 1. (11) Wang, Q. Y.; Yu, B. R.; Xu, C. J. Design and Control of Distillation System for Methylal/Methanol Separation. Part 1: Extractive Distillation Using DMF as an Entrainer. Ind. Eng. Chem. Res. 2012, 51, 1281. (12) Yu, B. R.; Wang, Q. Y.; Xu, C. J. Design and Control of Distillation System for Methylal/Methanol Separation. Part 2: Pressure-Swing Distillation with Full Heat Integration. Ind. Eng. Chem. Res. 2012, 51, 1293. (13) Xia, M.; Yu, B. R.; Wang, Q. Y.; Jiao, H. P.; Xu, C. J. Design and Control of Extractive Dividing-Wall Column for Separating Methylal− Methanol Mixture. Ind. Eng. Chem. Res. 2012, 51, 16016. (14) Xia, M.; Xin, Y.; Luo, J.; Li, W.; Shi, L.; Min, Y.; Xu, C. Temperature Control for Extractive Dividing-Wall Column with an Adjustable Vapor Split: Methylal/Methanol Azeotropic Separation. Ind. Eng. Chem. Res. 2013, 52, 17996. (15) Modla, G.; Lang, P. Removal and Recovery of Organic Solvents from Aqueous Waste Mixtures by Extractive and Pressure Swing Distillation. Ind. Eng. Chem. Res. 2012, 51, 11473. (16) Veerapur, R. S.; Patil, M. B.; Gudasi, K. B.; Aminabhavi, T. M. Poly(vinyl alcohol)−Zeolite T Mixed Matrix Composite Membranes for Pervaporation Separation of Water + 1,4-Dioxane Mixtures. Sep. Purif. Technol. 2008, 58, 377. (17) Taha, M.; Lee, M. J. Solubility and Phase Separation of 2-(NMorpholino)ethanesulfonic Acid (MES) and 4-(N-Morpholino)butanesulfonic Acid (MOBS) in Aqueous 1,4-Dioxane and Ethanol Solutions. J. Chem. Eng. Data 2011, 56, 4436. (18) Taha, M.; Teng, H. L.; Lee, M. J. Buffering-out: Separation of Tetrahydrofuran, 1,3-Dioxolen, or 1,4-Dioxane from their Aqueous Solutions Using EPPS Buffer at 298.15 K. Sep. Purif. Technol. 2013, 105, 33. (19) Modla, G. Energy Saving Methods for the Separation of a Minimum Boiling Point Azeotrope Using an Intermediate Entrainer. Energy 2013, 50, 103. (20) Luyben, W. L. Design and Control of a Fully Heat-Integrated Pressure-Swing Azeotropic Distillation System. Ind. Eng. Chem. Res. 2008, 47, 2681. (21) Li, W.; Shi, L.; Yu, B.; Xia, M.; Luo, J.; Shi, H.; Xu, C. New Pressure-Swing Distillation for Separating Pressure-Insensitive Maximum-Boiling Azeotrope via Introducing a Heavy Entrainer: Design and Control. Ind. Eng. Chem. Res. 2013, 52, 7836. (22) Yildirim, Ö .; Kiss, A. A.; Kenig, E. Y. Dividing Wall Columns in Chemical Process Industry: A Review on Current Activities. Sep. Purif. Technol. 2011, 80, 403. (23) Wu, Y. C.; Lee, H. Y.; Huang, H. P.; Chien, I. L. Energy-Saving Dividing-Wall Column Design and Control for Heterogeneous Azeotropic Distillation Systems. Ind. Eng. Chem. Res. 2014, 53, 1537. (24) Luyben, W. L. Principles and Case Studies of Simultaneous Design; Wiley: New York, 2011. (25) Wu, Y. C.; Chien, I. L. Design and Control of Heterogeneous Azeotropic Column System for the Separation of Pyridine and Water. Ind. Eng. Chem. Res. 2009, 48, 10564. (26) Chang, W. T.; Huang, C. T.; Cheng, S. H. Design and Control of a Complete Azeotropic Distillation System Incorporating Stripping Columns for Isopropyl Alcohol Dehydration. Ind. Eng. Chem. Res. 2012, 51, 2997. (27) Dejanović, C.; Matijašević, Lj.; Halvorsen, I. J.; Skogestad, S.; Jansen, H.; Kaibel, B.; Olujić, Ž . Designing Four-Product Dividing Wall Columns for Separation of a Multicomponent Aromatic Mixture. Chem. Eng. Res. Des. 2011, 89, 1155. (28) Ling, H.; Luyben, W. L. New Control Structure for DividedWall Columns. Ind. Eng. Chem. Res. 2009, 48, 6034.

and the total annual cost by 22.99% as compared to the base case. Furthermore, applying process intensification by thermally coupling the two column into a dividing-wall column allows much better savings in the total operating cost and total annual cost to be realized. It was found that significant 49.09% savings in the operating costs and 43.31% savings in the total annual cost could be realized by this proposed design. The proposed method uses the light impurity as the entrainer, so no foreign component is added into the system. An additional benefit is that the original three-column system can be reduced to a dividing-wall column with only one column shell. This article also demonstrates that the stripping-column design proposed by Chang et al.26 might not save much of the total reboiler duty for all heterogeneous azeotropic column systems. With the stripping-column design, energy carried by the overhead vapor can be conserved into the heterogeneous azeotropic column instead of being removed in a condenser. However, because the stripping column has only one equilibrium stage for the feed stream to purify into the top product, the residue curve maps and liquid−liquid envelope of the studied system can be used to predict whether significant savings of reboiler duty can be achieved. Regarding the control performance of the proposed heterogeneous dividing-wall column design, the two original important control degree of freedoms (two reboiler duties) are still preserved as compared to the conventional two-column system by heterogeneous azeotropic distillation. It was found that, by using just the conventional tray-temperature control strategy, the two products can be maintained at high purity despite disturbances in feed composition, feed flow rate, and liquid split ratio.

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AUTHOR INFORMATION

Corresponding Author

*Tel.: +886-3-3366-3063. Fax: +886-2-2362-3040. E-mail: [email protected]. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS Research funding from the National Science Council of R.O.C. under Grant NSC 103-2218-E-002-010 is greatly appreciated.

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REFERENCES

(1) Tyreus, B. D. DistillationEnergy Conservation and Process Control, A 35 Year Perspective. Presented at the AIChE Annual Meeting, Minneapolis, MN, Oct 16−21, 2011. (2) Julka, V.; Chiplunkar, M.; O’Young, L. Selecting Entrainers for Azeotropic Distillation. Chem. Eng. Prog. 2009, 105 (3), 47. (3) Luyben, W. L.; Chien, I. L. Design and Control of Distillation Systems for Separating Azeotropes; Wiley: New York, 2010. (4) Luyben, W. L. Comparison of Pressure-Swing and ExtractiveDistillation Methods for Methanol-Recovery System in the TAME Reactive-Distillation Process. Ind. Eng. Chem. Res. 2005, 44, 5715. (5) Wu, Y. C.; Chien, I. L.; Luyben, W. L. Two-Stripper/Decanter Flowsheet for Methanol Recovery in the TAME Reactive-Distillation Process. Ind. Eng. Chem. Res. 2009, 48, 10532. (6) Arifin, S.; Chien, I. L. Combined Preconcentrator/Recovery Column Design for Isopropyl Alcohol Dehydration Process. Ind. Eng. Chem. Res. 2007, 46, 2535. (7) Arifin, S.; Chien, I. L. Design and Control of an Isopropy Alcohol Dehydration Process via Extractive Distillation Using Dimethyl Sulfoxide as an Entrainer. Ind. Eng. Chem. Res. 2008, 47, 790. 15685


Investigation of the Energy-Saving Design of an Industrial 1,4

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