Mass Integration as a Design Heuristic ... - American Chemical Society

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Mass Integration as a Design Heuristic: Improvements in the HDA Process Carlos D. Fischer*,†,‡ and Oscar A. Iribarren*,†,‡ † ‡

Institute for Process Development and Design, INGAR UTNCONICET, Avellaneda 3657 (3000) Santa Fe, Argentina Universidad Tecnologica NacionalFacultad Regional Reconquista Acceso Parque, Industrial Calle 44 No 1000, (3560) Reconquista, Argentina ABSTRACT: This paper explores mass exchanging the outlet and inlet streams of a reactor, as a design heuristic within the hierarchical process design procedure by Douglas [AIChE J. 1985, 31 (3), 353361 and Conceptual Design of Chemical Processes; McGrawHill, 1988], who worked on the HDA process to test the proposal. The heuristic is used at an early stage of the hierarchy, when deciding the recycle and separation system structure of the process. If the reaction requires operating conditions with reactants in excess or that catalyze the reaction, which must be removed after the reaction, there is a concentration gradient between the inlet and outlet streams that may be used as the driving force in a mass exchanger (if such a device is available for the particular case). When applied to the HDA process, this methodology generated alternatives different from the previously proposed by other authors by resorting to a ceramic membrane gas permeation unit to perform the mass exchange of hydrogen. The performance of applying the heuristic was tested comparing the flow sheets proposed by several authors with and without this mass exchanger. The success of implementing this mass exchange networks synthesis concept was dependent on the concentration of the component to be transferred in the rich stream (i.e., it works if there is an appropriate driving force).

1. INTRODUCTION This paper explores the use of mass integration as a design heuristics within the hierarchical process design procedure by Douglas,1,2 who worked on the toluene hydrodealkylation (HDA) process to test the idea. When designing a new process following the hierarchical methodology of Douglas,1,2 one moves toward designing more-detailed versions of the process with an increasing number of process blocks interconnected by streams (generating new streams). The first level of decisions is the inputoutput (“i/o”) structure of the flow sheet. Raw materials, end products, and processing routes are defined at this level, yielding the overall structure of the components that enter and exit the process. The second level of decisions adds detail to the selected process alternative, deciding about the recycle structure of the flow sheet and selecting the unit operations to perform the required separations. This is done following heuristics that recommend alternatives for recycling components and criteria for selecting the unit operations. Finally, the procedure performs heat integration of the already-defined process streams. In an apparently independent approach (which is actually also a flow sheeting procedure), the technique for the synthesis of mass exchange networks (MENs)35 extrapolates pinch analysis of the heat exchange networks (HENs) to mass exchange, designing mass exchangers in a counter-current arrangement (as much as possible). As in the case of heat integration, the mass integration technique requires, as input information, the list of streams to be integrated, as well as their flow rates and concentrations (instead of temperatures, as in heat integration). Thus, both methodologies (for the integration of heat and mass) are usually applied, either to existing processes or in the last stage of the process design (once all the process streams have been r 2011 American Chemical Society

generated). Also, both approaches have an implicit assumption: the integration will not affect the operation conditions of the process. Although this is completely true in the case of heat integration (so one assumes that heating and cooling is eventually done with external sources, leaving heat integration of the process streams as the last step in the process design procedure), it is also a good approximation for most of the reported case studies of mass integration, where transfers are minimal (usually a contaminant in small concentrations in, for example, water, but not the principal components involved in the process). In this work, we explore using the technique of synthesizing MENs at an earlier stage of the hierarchy of Douglas,1,2 when deciding the recycle structure of the process. This can be done as soon as after the reaction is defined, which usually requires operating conditions with components in excess or that catalyze the reaction, which must be removed after the reactor. The concept of synthesizing MENs can be used as an additional heuristic rule (in competence with Douglas heuristics): “Explore the implementation of a Mass Exchanger between the streams exiting and entering the reactor”. This mass integration could eventually be performed with absorbers, strippers, or membrane systems arranged to operate in counter-current mode. If this material integration were possible, the subsequent separations and recycles should be minimized (or even not exist if a total integration between output and input could be achieved). Figure 1 illustrates the structures of the recycle and separation system that are determined by following the heuristics process design procedure by Douglas standing alone (Figure 1a), and Received: June 24, 2011 Accepted: October 11, 2011 Published: October 11, 2011 12664

dx.doi.org/10.1021/ie2013554 | Ind. Eng. Chem. Res. 2011, 50, 12664–12677

Industrial & Engineering Chemistry Research

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Figure 1. Douglas heuristic process design procedures: (a) standing alone and (b) including the MEN heuristics.

incorporating the MEN heuristics in this design procedure (Figure 1b). In a previous work,6 this concept was applied successfully to a biodiesel process. In this paper, we apply it to the HDA process originally used by Douglas1,2 to present his methodology, and we identify the process characteristics that enable the use of countercurrent mass exchangers in gases. Section 2 formally presents the problem to which we are going to apply the MEN heuristics, and the problem is compared with the original process synthesis problem. The methodology proposed for addressing this problem is presented and explained. Then, in section 3, we describe the traditional and custom models used for the membrane modules (the models used for other unit operations are standard). In the next section (section 4), example alternative processes are generated by applying the MEN concept to different HDA process configurations that have been proposed previously by other authors, and the variations produced in operating conditions are analyzed. Finally, section 5 compares the performance of the alternatives and section 6 presents the conclusions of this paper.

2. DESCRIPTION OF THE PROBLEM Let us consider a process where the product-forming reaction occurs between a component A and a component B (both gaseous, at an intermediate or high pressure), which react to form a product C (and eventually other byproducts). Usually, one of the reactants (let us assume component B) is present in excess, to accelerate the reaction. Therefore, this component B is present in abundance in the stream that is exiting the reactor, along with product C and small amounts of the limiting reactant if the progress of the reaction was not complete (usually,2 the reaction seldom progresses beyond 98% completion). Traditional gas recovery systems that use membranes to perform the separation work with a considerable pressure difference (30 bar or larger) through the membrane.7,8 In the case described, we take advantage of the intermediate or high pressure available at the reactor exit stream to drive the desired component through the membrane; a low-pressure permeate stream is obtained,

Figure 2. Depictions of (a) a traditional membrane separation, (b) the MEN concept used to generate a recycle and separation structure, and (c) the MEN concept plus the traditional separation system.

which must be recompressed to recycle it into the reactor. In other systems where the reaction occurs at a low pressure, the entire reactor exiting stream should be compressed (including components A and C), giving much higher compression costs and making membrane separation less attractive. For a process such as the one described, the structure of a traditional separation and recycle system using membranes is shown in Figure 2a. Note that the membrane permeates component B, which is then recompressed to be fed into the reactor, along with a fresh stream of component B. In this system, the main cost of implementing the recycle is given by the compressor and, to a lesser extent, the required area of membrane. Using the MEN concept as a heuristic at this level of the design procedure of Douglas1,2 (when deciding the structure of the recycle and separation system), we can exchange mass between the B-rich stream exiting the reactor and a B-poor reactor inlet stream. The process alternative generated is shown in Figure 2b. Component B in the stream that is leaving the reactor is at a high partial pressure, while the stream that feeds A into the reactor has 12665

dx.doi.org/10.1021/ie2013554 |Ind. Eng. Chem. Res. 2011, 50, 12664–12677

Industrial & Engineering Chemistry Research no component B at all and, therefore, its partial pressure is null. Thus, we have a considerable gradient of partial pressure to transfer component B, even if the stream entering the reactor is already pressurized. We use this partial pressure gradient to drive a mass exchange that is equivalent to recycling B from the output to the input of the reactor, without the need to separate component B from the other components present in the streams leaving the reactor nor recompressing a stream of recycled component B. Actually, we obviously could not transfer the entire amount of component B. In a limiting case, we could exchange up to the point where the partial pressure of component B is the same on both sides of the membrane. In a real case, we must maintain a sufficiently large gradient across the membrane for the mass exchange unit to have acceptable dimensions. To recover the remaining component B, we may resort to the traditional separation process with membranes. This alternative is shown in Figure 2c. In this flow sheet, the areas of mass exchange and the size of the recirculation compressor should be optimized to give a minimum total annualized cost. Depending on the particular process, the mass exchanger will be able to recover and recycle different percentages of the component that is in excess. However, in all cases, the remaining separation task performed with a traditional separation system will be smaller and have a lower cost.

3. MEMBRANE MODELS In this paper, we use Aspen Plus V7.2 to simulate the steady state and assess the performance of different process alternatives; but Aspen Plus V7.2 lacks, among its build-in operation blocks, one for the permeation of gases. However, Aspen Custom Modeler V7.2 provides, among the examples of personalized modules, one to perform gas permeation. This example module must be exported to be used as a personalized block in Aspen Plus V7.2. Exporting this module requires a C++ or FORTRAN compiler that must be installed, tested, and configured to be executed from the command line. Luyben et al.9 provided in their paper a personalized permeation module to analyze the dynamic effects on a process with cross-flow and counter-current membrane modules. They do not export the modules to Aspen Plus but, instead, simply use them in Aspen Dynamics. The models used for these gas permeation modules cannot be applied to the operation proposed in the present paper. The two principal limitations for using these existing modules are the inlet streams and the direction of permeate fluxes: (a) The counter-current arrangement used by these authors is different from the traditional one: an additional inlet stream must be modeled. (b) They model the permeate flux as a total flux that consists of different components, all permeating from the retentate side to the permeate side of the membrane. These components permeate according to their respective permeability and partial pressure gradient, with all of the gradients (and, therefore, all fluxes) being positive from the retentate side to the permeate side of the membrane. In our case, when exchanging between the streams entering and exiting the reactor, the partial pressure gradients of some components are negative. Therefore, the fluxes of each component must be modeled independently and not as a total flux with molar fractions of different components. Figure 3 schematizes

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the models used for the cross-flow membrane model (Figure 3a) and the necessary counter-current membrane module (Figure 3b). The equations that model mass transfer in each cell n are given as follows: Balances for each component in the rich stream: FRn1 yc, n1 ¼ FRn yc, n þ Nc, n Balances for each component in the lean stream: FLn1 xc, n1 ¼ FLn xc, n þ Nc, n Permeate fluxes for each component: Nc, n ¼ An Lc ðPRyc, n PLxx, n Þ FRn1 and FRn are the flow rates of the rich stream entering and exiting each cell n (expressed in units of kmol/h), while yc,n1 and yc,n are the molar fractions of each component entering the reactor. Nc,n represent the permeate flux of each component in cell n (expressed in units of kmol/h). For the lean stream, FLn1 and FLn are the flow rates entering and exiting each cell n, with molar fractions xc,n1 and xc,n, respectively. Lc are the permeability of each component. An is the membrane area (given in units of m2) of each cell, while PR and PL are the total pressures of the rich and lean streams, respectively. Using these equations, we implemented a module in Aspen Custom Modeler 7.2 (ACM), taking, as a basis, the example module provided in ACM and modifying it to represent the mass transfer described above. We compiled and exported this module to be able to use it as a block in the “ACMModels” library of Aspen Plus 7.2. The script used for this module is presented in the Appendix.

4. EXAMPLE: HDA PROCESS To illustrate the new process alternatives generated when adding the MEN concept as a heuristic when deciding the recycle and separation structure of the flow sheet, we use the well-known process HDA (hydrodealkylation of toluene to benzene) originally used by Douglas1,2 to present his methodology. To implement and assess the steady-state process alternatives generated, we use Aspen Plus V7.2 with the PengRobinson equation of state as a physical property model. The HDA process is well-known and has been the subject of many studies, with regard to the synthesis and optimization of processes,912 thus producing different and varied optimal process flow sheets, depending on the parameters used by different authors. The traditional process designed by Douglas1,2 has an extent of reaction of 70%75%. Figure 4 shows a flow sheet of this process before heat integration (heating and cooling are performed with external utilities, without exchanging heat among process streams), which is the last level of decisions in the hierarchy of Douglas.1,2 The most relevant streams for a production of 125 kmol/h of benzene are presented in Table 1. The reactions of interest are toluene þ H2 f benzene þ CH4 2ðbenzeneÞ T diphenyl þ H2 The reaction is homogeneous and takes place in the range of 621667 °C (below this temperature range, the rate of reaction 12666



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Figure 3. (a) Cross-flow gas permeation module; (b) counter-current gas permeation module.

Figure 4. HDA flow sheet reported by Douglas.1,2

is too low and above this temperature range, hydrocracking reactions are dominant) and a pressure ca. 35 atm. An excess of

hydrogen (5:1 minimum) is necessary to prevent coking at the reactor. The gas leaving the reactor must be quickly quenched to 12667



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Table 1. Stream Data in the Douglas HDA Flow Sheet Flow Rates (kmol/h) temperature (°C) pressure (bar)

H2

CH4

Composition (mol fraction)

benzene toluene diphenyl

total

H2

CH4


FFH2

30.00

43.06

225.27

6.97

0.00

0.00

0.00

232.24 0.9700 0.0300

0.0000

0.0000

0.0000

FFTOL

30.00

43.06

0.00

0.00

0.00

134.98

0.00

134.98 0.0000 0.0000

0.0000

1.0000

0.0000

RECYCTOL

146.70

41.80

0.00

0.00

0.42

31.02

0.06

31.50 0.0000 0.0000

0.0130

0.9850

0.0020

RIN

621.10

36.68

906.20

986.78

20.39

167.78

0.06

2081.21 0.4350 0.4740

0.0100

0.0810

0.0000

ROUT

667.80

36.68

774.90 1121.42

148.36

33.14

3.40

2081.21 0.3720 0.5390

0.0710

0.0160

0.0020

OUT-Q LIQUID L-SEP

621.50

35.57

775.14 1123.77

187.99

42.96

4.47

2134.33 0.3630 0.5270

0.0880

0.0200

0.0020

48.90

33.44

9.83

165.28

40.94

4.47

221.52 0.0050 0.0440

0.7460

0.1850

0.0200

49.40 48.90

38.91 33.44

0.76 7.48 774.14 1113.93

125.65 22.71

31.12 2.03

3.40 0.00

168.40 0.0050 0.0440 1912.81 0.4050 0.5820

0.7460 0.0120

0.1850 0.0010

0.0200 0.0000

GAS PURGE

1.00

73.80

41.85

93.21

134.13

2.73

0.24

0.00

230.31 0.4050 0.5820

0.0120

0.0010

0.0000

GASRECYC

73.80

41.85

680.93

979.81

19.98

1.78

0.00

1682.49 0.4050 0.5820

0.0120

0.0010

0.0000

BENZENE

105.60

2.07

0.00

0.00

124.97

0.03

0.00

125.00 0.0000 0.0000

1.0000

0.0000

0.0000

Figure 5. HDA flow sheet studied by Bouton and Luyben.9

621 °C, to prevent coking in the next heat exchanger. The fresh feed streams of toluene and hydrogen are heated and mixed with the recycle streams of toluene and hydrogen before being fed into the reactor. The stream exiting the reactor contains hydrogen, methane, benzene, toluene, and diphenyl. Most of the hydrogen and methane are separated from the aromatics using a partial condenser, followed by a flash that separates the light gases, which are recycled after purging some of this stream to prevent the accumulation of methane (which is an impurity of hydrogen feed and also produced by the reaction) in the process. Some of the liquid stream leaving the flash is used to quench the hot gases that are leaving the reactor. The remaining liquid stream goes through a distillation train. Because all of the hydrogen and methane could not be separated by the flash, they are removed in a stabilizer distillation column. Benzene is separated in the second distillation column, and the third column separates toluene from diphenyl.

Konda et al.11 proposed alternatives for this process using separation membranes and explored the optimal configuration. The alternative proposals are as follows: (a) Placement of a membrane system in the purge gas. This alternative does not change the operation conditions of the process. The benefit is a reduction of the fresh hydrogen requirement, because of the hydrogen recovered. This alternative is studied in detail by Bouton and Luyben9 and is presented in Figure 5. Steady-state simulation data are presented in Table 2. (b) Placement of a membrane system in the main gas stream (leaving from the flash gasliquid separator). This option is recommended by Konda et al.,11 even though the design of the reactor must be modified. Because the gas recycle is virtually free of methane (which acts as a heat carrier), the adiabatic reactor exceeds temperature, burning 12668



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Table 2. Stream Data for the HDA Flow Sheet Studied by Bouton and Luybena Flow Rates (kmol/h) temperature (°C) pressure (bar)

CH4


total

H2

CH4


30.00

43.06

0.00

0.00

0.00

134.83

0.00

134.83 0.9700 0.0300

0.0000

0.0000

0.0000

RECYCTOL

146.90

41.80

0.00

0.00

0.46

30.82

0.22

31.50 0.0000 0.0000

0.0000

1.0000

0.0000

RIN

621.10

36.68

906.20

984.33

20.21

167.40

0.22

2078.36 0.0000 0.0000

0.0150

0.9790

0.0070

ROUT

667.80

36.68

774.85 1118.87

148.38

32.86

3.41

2078.36 0.4360 0.4740

0.0100

0.0810

0.0000

LIQUID

48.90

33.44

165.30

40.59

4.48

221.18 0.3730 0.5380

0.0710

0.0160

0.0020

OUT-Q L-SEP

621.50

35.57

49.40

38.91

GASRECYC

48.90 73.80

PURGE

FFTOL

1.00

9.81

188.01

42.60

4.48

2131.42 0.3640 0.5260

0.0880

0.0200

0.0020

7.46

125.66

30.86

3.40

168.14 0.0050 0.0440

0.7470

0.1840

0.0200

33.44 41.85

774.10 1111.43 673.26 966.64

22.71 19.75

2.01 1.75

0.00 0.00

1910.25 0.0050 0.0440 1661.40 0.4050 0.5820

0.7470 0.0120

0.1840 0.0010

0.0200 0.0000

73.80

41.85

100.84

144.79

2.96

0.26

0.00

248.85 0.4050 0.5820

0.0120

0.0010

0.0000

PERMEATE

73.80

8.84

81.71

13.01

0.00

0.00

0.00

94.72 0.4050 0.5820

0.0120

0.0010

0.0000

RETENTAT

73.80

41.85

19.13

131.77

2.96

0.26

0.00

154.13 0.8630 0.1370

0.0000

0.0000

0.0000

105.60

2.07

0.00

0.00

124.97

0.03

0.00

125.00 0.1240 0.8550

0.0190

0.0020

0.0000

30.00

43.06

151.23

4.68

0.00

0.00

0.00

155.91 0.0000 0.0000

1.0000

0.0000

0.0000

50.00

19.59

81.72

13.01

0.00

0.00

0.00

94.73 0.8626 0.1374

0.0000

0.0000

0.0000

GAS

BENZENE FFH2 S13 a

H2

Composition (mole fraction)

775.10 1121.23 0.76

Data taken from ref 9.

Figure 6. HDA flow sheet proposed by Konda et al.11

the aromatic compounds. Therefore, they use two adiabatic reactors in series, with intermediate cooling. They do not provide a detailed flow sheet of this alternative; they describe only the more-significant features, focusing on the operation of the membrane systems. Using the information available, we have simulated the corresponding flow sheet to obtain the data that have not been supplied. However, to simplify further comparisons among alternatives, we considered a single nonadiabatic reactor, respecting the inlet and outlet temperatures of the reaction. This alternative is presented in Figure 6 and Table 3. The flow sheet has only two distillation columns, because these authors propose recycling the diphenyl, to allow it to build up to its equilibrium level.

These alternative configurations are determined by thinking of membranes as a separation unit (one inlet stream that is split into two outlet streams of different compositions). We wish to explore the effect of placing a direct counter-current mass exchange between the input and output streams of the reactor. This process has two reactions, one of which is an equilibrium one that can greatly affect the selectivity. If we separate the hydrogen at high temperature just after the reactor, we favor the conversion of benzene to diphenyl and hydrogen. Therefore, before reducing the hydrogen content of this stream, quenching and cooling is required. We consider that (i) 300 °C is cool enough to avoid benzene decomposition, yet still hot enough for all components to be gaseous, and (ii) that gas-permeation ceramic-type membranes are able to operate at this temperature and generate the following process alternatives: 12669



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Table 3. Streams Data for the HDA Flow Sheet Proposed by Konda et al.a Flow Rates (kmol/h) benzene

toluene

diphenyl

total

H2

CH4

benzene

toluene

diphenyl

0.00

0.00

143.13

0.9700

0.0300

0.0000

0.0000

0.0000

134.89

0.00

134.89

0.0000

0.0000

0.0000

1.0000

0.0000

3.61

6.22

0.0000

0.0000

0.0000

0.4190

0.5810

temperature (°C)

pressure (bar)

H2

FFH2

30.00

43.06

138.84

4.29

0.00

FFTOL

30.00

43.06

0.00

0.00

0.00

TOL-DIPH

186.00

2.35

0.00

0.00

0.00

2.61

RIN

621.10

36.68

737.98

39.04

0.00

137.50

3.61

918.13

0.8040

0.0430

0.0000

0.1500

0.0040

ROUT

667.40

33.44

603.23

173.79

134.75

2.75

3.61

918.13

0.6570

0.1890

0.1470

0.0030

0.0040

OUT-Q LIQUID

620.20

33.44

603.40

174.19

156.30

3.21

4.23

941.33

0.6410

0.1850

0.1660

0.0030

0.0040

48.90

33.44

1.18

2.70

146.57

3.14

4.23

157.82

0.0070

0.0170

0.9290

0.0200

0.0270

GAS

49.50 48.90

38.91 33.44

1.01 602.22

2.30 171.49

125.02 9.73

2.68 0.08

3.61 0.00

134.62 783.52

0.0070 0.7690

0.0170 0.2190

0.9290 0.0120

0.0200 0.0000

0.0270 0.0000

PURGE

69.40

33.44

3.08

136.74

9.73

0.08

0.00

149.62

0.0210

0.9140

0.0650

0.0010

0.0000

PERMEATE

48.90

3.00

599.14

34.75

0.00

0.00

0.00

633.89

0.9450

0.0550

0.0000

0.0000

0.0000

GASRECYC

50.00

41.85

599.14

34.75

0.00

0.00

0.00

633.89

0.9452

0.0548

0.0000

0.0000

0.0000

BENZENE S13

105.60

2.07

0.00

0.00

124.93

0.07

0.00

125.00

0.0000

0.0000

0.9990

0.0010

0.0000

50.00

19.59

599.14

34.75

0.00

0.00

0.00

633.89

0.9452

0.0548

0.0000

0.0000

0.0000

L-SEP

a

CH4



(a) The HDA flow sheet of Douglas,1,2 taking all the gas flow coming from the reactor after quenching and cooling to 300 °C as the rich stream and the feed of toluene to the reactor heated to 300 °C as the lean stream. (b) The HDA flow sheet studied by Bouton and Luyben,9 taking the same rich and lean streams as in the previous case. (c) The HDA flow sheet proposed by Konda et al.,11 taking the gas stream leaving the flash heated to 300 °C as rich stream (here, we explored using a stream with a larger H2 content to perform the mass exchange) and again the toluene feed to the reactor heated to 300 °C as the lean stream. Placing a membrane module between the described streams to perform mass exchange (mainly hydrogen), and knowing the concentrations and partial pressure of hydrogen in these flows, we can easily compute the amount of hydrogen that is transferable. This recovery of hydrogen produces a variation in the concentrations of streams, such that all of the subsequent separation system and recycle must be recalculated. To compute the properties of all streams with Aspen Plus V7.2, we use the custom module developed in ACM and add it to the flow sheet with the parameters of a membrane capable of performing the separation in the described conditions. There are different types of ceramic membranes that can be used to carry out the separation at this temperature. Among them, the literature reports ceramic membranes that are composed of zeolite, coal, silica, and metallic alloys. Membranes are customized for each particular application, although none have been developed for the specific application that we are proposing. Nevertheless, a ceramic membrane of zeolite ZSM-5 is likely to operate under these conditions. The literature13 has reported a ceramic membrane of zeolite ZSM-5 with a total permeability of 0.244 kmol/(m2 bar h), with a H2-to-CH4 selectivity of 16, that operates in the process for steam reforming of methane with a feed composition of 76% H2 and 24% CH4, producing a permeate stream enriched to 98% in H2. Considering the pressure gradients and flow rates of each component, we can compute the component permeabilities to be 0.31463 for hydrogen and 0.02033 for methane. Although these figures may vary with pressure, temperature, and composition, using these permeabilities

gives us an approximation to the performance that can be expected in the HDA process. 4(a). The HDA Flow Sheet of Douglas.1,2 The traditional process cools the reactor outlet stream up to a temperature of 48.9 °C. We modify the process by performing an initial cooling up to 300 °C, which is appropriate to perform the mass integration between the gaseous streams leaving and entering the reactor. This alternative is shown in Figure 7 and Table 4. The stream values correspond to an exchange area of 120 m2, with a hydrogen partial pressure gradient of >0.5 bar all along the unit. We see that the mass exchange can recover only a small percentage of the hydrogen present in the rich stream. Because the amount of hydrogen exchanged is only 79.59 kmol/h (10.61% of the 750.05 kmol/h of hydrogen in the rich stream), the main impact on the process is in the gas recovery system, with a slight reduction in the compression system. In addition to the hydrogen, 36.50 kmol/h of methane also permeates. The flow of gas driven by the compressor is reduced from 1912.81 kmol/h to 1794.95 kmol/h. This results in a decrease of required power from 451.01 kW to 421.94 kW (taking an isentropic efficiency equal to 0.72), which amounts to a power reduction of 6.45%. Direct recovery and recycle of some of the hydrogen reduced the hydrogen feed to a flow rate of 13.58 kmol/h (5.85% of the original flow rate of 232.24 kmol/h). The rest of the separation system (the separation of liquids) remains without significant changes. 4(b). The HDA Flow Sheet Studied by Bouton and Luyben9. In this alternative, we perform mass exchange between the same streams as in the previous case. The other membrane unit in the flow sheet (to recovery hydrogen from the gas purge) is a polymer membrane with a permeability of 0.201 and a H2-toCH4 selectivity of 110 (as used by Konda et al.11) operating at 73.60 °C with the retentate stream at 41.85 bar and the permeate stream at 8.84 bar (according to data presented in the flow sheet). This alternative is shown in Figure 8 and Table 5. The new mass exchange unit has an area of 120 m2. Here again, the mass exchange unit recovers a low percentage of the hydrogen that is present in the rich stream. The amount of hydrogen exchanged is 82.22 kmol/h (11.82% of the 695.33 kmol/h of hydrogen in the reactor outlet stream); in addition to 12670



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Figure 7. HDA flow sheet of Douglas,1,2 incorporating the MEN concept.

Table 4. Streams Data in Douglas HDA Flow Sheet Incorporating the MEN Concept Flow Rates (kmol/h) temperature (°C) pressure (bar)

H2

CH4



total

H2

CH4


30.00

43.06

212.10

6.56

0.00

0.00

0.00

218.66 0.9700 0.0300

0.0000

0.0000

0.0000

30.00

43.06

0.00

0.00

0.00

134.53

0.00

134.53 0.0000 0.0000

0.0000

1.0000

0.0000

RIN

147.30 621.10

41.80 36.68

0.00 0.00 880.59 1012.45

0.02 18.78

32.62 168.93

0.06 0.06

32.70 0.0000 0.0000 2080.81 0.4230 0.4870

0.0010 0.0090

0.9980 0.0810

0.0020 0.0000

ROUT

667.20

36.68

749.82 1146.53

146.24

34.85

3.37

2080.81 0.3600 0.5510

0.0700

0.0170

0.0020

INRICH

300.00

33.44

750.05 1149.03

185.65

45.21

4.43

2134.36 0.3510 0.5380

0.0870

0.0210

0.0020

OUTRICH

300.00

33.44

670.46 1112.53

185.65

45.21

4.43

2018.27 0.3320 0.5510

0.0920

0.0220

0.0020

INLEAN

300.00

40.02

0.00

0.00

0.02

167.15

0.06

167.23 0.0000 0.0000

0.0000

1.0000

0.0000

OUTLEAN LIQUID

300.00

40.02

79.59

36.50

0.02

167.15

0.06

283.32 0.2810 0.1290

0.0000

0.5900

0.0000

48.90

33.44

0.94

10.44

164.32

43.19

4.43

223.32 0.0040 0.0470

0.7360

0.1930

0.0200

GAS

49.40 48.90

38.91 33.44

0.72 7.94 669.52 1102.08

124.91 21.33

32.83 2.02

3.37 0.00

169.76 0.0040 0.0470 1794.95 0.3730 0.6140

0.7360 0.0120

0.1930 0.0010

0.0200 0.0000

PURGE

73.60

41.85

80.61

132.70

2.57

0.24

0.00

216.12 0.3730 0.6140

0.0120

0.0010

0.0000

GASRECYC BENZENE

73.60

41.85

588.90

969.39

18.76

1.78

0.00

1578.83 0.3730 0.6140

0.0120

0.0010

0.0000

100.60

2.07

0.00

0.07

124.89

0.04

0.00

125.00 0.0000 0.0010

0.9990

0.0000

0.0000

FFH2 FFTOL RECYCTOL

L-SEP

the hydrogen, 36.13 kmol/h of methane also permeates. Similar to the previous case, there is a slight reduction in the compression system. The gas flow driven by the main compressor is reduced from 1912.25 kmol/h to 1822.94 kmol/h, with a reduction from 450.42 kW to 428.88 kW in the required compressor power (with an isentropic efficiency of 0.72 in all cases). In turn, the permeate gas stream driven by the side compressors reduces from 94.72 kmol/h to 85.04 kmol/h with a reduction in the powers of the compressors from 96.32 kW and 83.89 kW to 86.38 kW and 75.26 kW. These reductions in the permeate stream produce a reduction in the total power of compression from 630.63 kW to 590.91 kW, which amounts to 6.36% of the power. Although the mass exchange unit recovers and recycles some of the hydrogen directly, the hydrogen feed is not reduced significantly; however, the route of recycle changed, reducing the compression system. The hydrogen feed flow rate is reduced by 1.12 kmol/h (0.71% of 155.91 kmol/h). The rest of the separation system (separation of liquids) remains unchanged.

4(c). The HDA Flow Sheet Proposed by Konda et al.11 In this alternative, proposed by Konda et al.,11 the membrane separation is on the main gas stream leaving the gasliquid separation at 48.9 °C and 33.44 bar. The retentate stream does not experience appreciable pressure losses (0.34 bar), whereas, on the permeate side, the pressure is only 3 bar. They used a polymeric membrane with a permeability of 4.5 (m3(STP))/(m2 bar h) (0.201 kmol/(m2 bar h)) and a H2-to-CH4 selectivity of 110. A membrane area of 766 m2 is needed to obtain the required recycle stream. In this flow sheet, all the gaseous streams are richer in hydrogen and, therefore, have a higher partial pressure. We explored choosing the stream leaving the gasliquid separator for mass exchanging with the toluene feed to the reactor, because it is richer in hydrogen than the stream leaving the reactor. Therefore, we place the mass exchange unit as shown in Figure 9. With this configuration, a higher percentage of the hydrogen available in the rich stream can be transferred to the lean stream. The rest of the hydrogen is recovered with a 12671



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Figure 8. HDA flow sheet studied by Bouton and Luyben,9 incorporating the MEN concept.

Table 5. Streams Data in the Flow Sheet Studied by Bouton and Luyben, Incorporating the MEN Concept Flow Rates (kmol/h) temperature (°C) pressure (bar) FFH2 FFTOL RECYCTOL

H2

CH4



total

H2

CH4


30.00

43.06

150.14

4.64

0.00

0.00

0.00

154.79

0.9700 0.0300

0.0000

0.0000

0.0000

30.00 147.00

43.06 41.80

0.00 0.00

0.00 0.00

0.00 0.25

135.71 32.40

0.00 0.06

135.71 32.70

0.0000 0.0000 0.0000 0.0000

0.0000 0.0080

1.0000 0.9910

0.0000 0.0020

RIN

621.10

36.68

908.20 1014.93

18.95

169.92

0.06

2112.05 0.4300 0.4810

0.0090

0.0800

0.0000

ROUT INRICH

666.90

36.68

777.31 1149.02

146.64

35.83

3.26

2112.05 0.3680 0.5440

0.0690

0.0170

0.0020

300.00

33.44

777.55 1151.49

186.12

46.48

4.28

2165.92 0.3590 0.5320

0.0860

0.0210

0.0020

OUTRICH

300.00

33.44

695.33 1115.36

186.12

46.48

4.28

2047.56 0.3400 0.5450

0.0910

0.0230

0.0020 0.0000

INLEAN

300.00

40.02

0.00

0.00

0.25

168.11

0.06

168.41

0.0000 0.0000

0.0010

0.9980

OUTLEAN

300.00

40.02

82.22

36.13

0.25

168.11

0.06

286.76

0.2870 0.1260

0.0010

0.5860

0.0000

LIQUID

48.90 49.40

33.44 38.91

0.97 0.74

10.36 7.88

164.62 125.15

44.39 33.75

4.28 3.26

224.62 170.76

0.0040 0.0460 0.0040 0.0460

0.7330 0.7330

0.1980 0.1980

0.0190 0.0190 0.0000

L-SEP GAS

48.90

33.44

694.36 1104.99

21.50

2.09

0.00

1822.94 0.3810 0.6060

0.0120

0.0010

GASRECYC

73.60

41.85

603.91

961.05

18.70

1.82

0.00

1585.46 0.3810 0.6060

0.0120

0.0010

0.0000

PURGE

73.60

41.85

90.46

143.95

2.80

0.27

0.00

237.48

0.0120

0.0010

0.0000

PERMEATE

73.60

8.84

71.93

13.10

0.00

0.00

0.00

85.04

0.8460 0.1540

0.0000

0.0000

0.0000

RETENTAT

73.60

41.85

18.52

130.84

2.80

0.27

0.00

152.44

0.1220 0.8580

0.0180

0.0020

0.0000

BENZENE S13

100.20

2.07

0.00

0.07

124.90

0.03

0.00

125.00

0.0000 0.0010

0.9990

0.0000

0.0000

50.00

19.59

71.93

13.10

0.00

0.00

0.00

85.04

0.8459 0.1541

0.0000

0.0000

0.0000

conventional membrane system, as proposed by Konda et al.11 Table 6 shows the new streams data for this alternative. Analyzing these results, we see that 272.28 kmol/h of hydrogen and 38.06 kmol/h of methane are exchanged. These values correspond to a 45.24% of the hydrogen and 19.37% of the methane present in the rich stream. In this way, the gas stream compressed by the compressor is reduced from 633.89 kmol/h to 347.01 kmol/h. This resulted in a reduction in compressor power from 1638.62 kW and 563.10 kW to 896.34 kW and 308.18 kW. These reductions in the flow rates of the gaseous streams produce a decrease in the total power of compression from 2201.72 kW to 1204.52 kW, which amounts to a decrease of 45.29%. Because the gas recycle handles a much smaller gas volume, the original 766 m2

0.3810 0.6060

of polymeric membrane is no longer needed and only 460 m2 suffices, operating at the same pressure levels. Ceramic zeolite membranes have less selectivity than the polymeric membranes; therefore, 1.48 kmol/h of extra feed hydrogen are needed to provide the same reaction conditions. The rest of the separation system (separation of liquids) remains without significant changes.

5. ECONOMIC ANALYSIS To analyze the success (or failure) of implementing the MEN concept as a heuristic when deciding the structure of the recycle and separation system, we compute the costs of the different alternatives. Analyzing the streams data and equipment involved 12672



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Figure 9. HDA flow sheet proposed by Konda et al.,11 incorporating the MEN concept.

Table 6. Streams Data in the Flow Sheet Proposed by Konda et al.,11 Incorporating the MEN Concept Flow Rates (kmol/h) CH4

benzene

diphenyl

total

H2

CH4

benzene

toluene

diphenyl

0.00

0.00

144.61

0.9700

0.0300

0.0000

0.0000

0.0000

135.19

0.00

135.19

0.0000

0.0000

0.0000

1.0000

0.0000

2.67 137.85

3.63 3.63

6.30 943.44

0.0000 0.7820

0.0000 0.0680

0.0000 0.0000

0.4230 0.1460

0.5770 0.0040

135.09

2.76

3.63

943.44

0.6390

0.2110

0.1430

0.0030

0.0040

10.14

0.08

0.00

808.65

0.7440

0.2430

0.0130

0.0000

0.0000

158.46

10.14

0.08

0.00

498.31

0.6610

0.3180

0.0200

0.0000

0.0000

0.00

0.00

0.00

137.85

3.63

141.48

0.0000

0.0000

0.0000

0.9740

0.0260

40.02

272.28

38.06

0.00

137.85

3.63

451.82

0.6030

0.0840

0.0000

0.3050

0.0080

33.44

1.15

3.00

146.49

3.14

4.26

158.03

0.0070

0.0190

0.9270

0.0200

0.0270

38.91 33.44

0.98 4.20

2.56 136.88

124.95 10.14

2.68 0.08

3.63 0.00

134.80 151.30

0.0070 0.0280

0.0190 0.9050

0.9270 0.0670

0.0200 0.0010

0.0270 0.0000

48.90

3.00

325.43

21.59

0.00

0.00

0.00

347.01

0.9380

0.0620

0.0000

0.0000

0.0000

GASRECYC

50.00

41.85

325.43

21.58

0.00

0.00

0.00

347.01

0.9378

0.0622

0.0000

0.0000

0.0000

BENZENE S13

103.10

2.07

0.00

0.04

124.95

0.01

0.00

125.00

0.0000

0.0000

1.0000

0.0000

0.0000

50.00

19.59

325.43

21.58

0.00

0.00

0.00

347.01

0.9378

0.0622

0.0000

0.0000

0.0000

temperature (°C)

pressure (bar)

H2

30.00

43.06

140.27

4.34

0.00

30.00

43.06

0.00

0.00

0.00

RIN

185.50 621.10

2.35 36.68

0.00 737.98

0.00 63.98

0.00 0.00

ROUT

667.40

33.44

602.89

199.07

INRICH

300.00

33.44

601.91

196.52

OUTRICH

300.00

33.44

329.63

INLEAN

300.00

40.02

OUTLEAN LIQUID

300.00 48.90

PURGE

49.50 68.50

PERMEATE

FFH2 FFTOL TOL-DIPH

L-SEP

in each alternative, we notice that the main changes occurred in the gas recycle flows, which impacts the size of the compressors and the number and size of the membrane units. We also note a variation in the amount of fresh hydrogen necessary to provide the same reaction conditions. We did not consider the changes in the heat exchange network (HEN), because it is a well-known methodology, applied in a lower level of the process design procedure. Nevertheless, we can observe that neither the total energy exchanged nor the initialfinal temperature targets change significantly, so one can expect that the total area of heat exchange will also remain approximately the same (even if the number of heat exchangers does not). In this way, we perceive that the cost of heat exchange will not change much. Therefore, to assess the performance of the proposed approach in each of the studied examples, we compare the annualized costs of

toluene


compressors and membrane units, plus the cost savings of hydrogen feed and electric energy consumption for compression, before and after implementation of the methodology. For the compressor installed cost, we use the correlation by Douglas:2 MS compressor installed cost ð$Þ ¼ 517:5 280 ðbhpÞ0:82 ð2:11 þ FdÞ where $ denotes U.S. dollars, M&S is the Marshal and Swift inflation index (M&S = 1477.7 for the year 2009), bhp is the brake horsepower (calculated with Aspen Plus V7.2 with an isentropic efficiency of 0.72), and Fd is the design factor (Fd = 1 for centrifugal compressors). 12673



ARTICLE

Table 7. Comparison for the HDA Flow Sheet of Douglasa

Table 9. Comparison for the HDA Flow Sheet of Konda et al.a

Case

parameter

HDA by

HDA by Douglas +

Douglas

MEN

Case HDA by parameter

hydrogen recovery (kmol/h)

13.58

ceramic membrane area (m2)

120.00


421.94

ceramic membrane area (m2) polymeric membrane area (m2)

brake power (kW) compressor

451.01

compressor 1

1638.62

896.34

compressor 2

563.10

308.18

1 002 299.51

1 082 197.35

ceramic membrane

7.97

polymeric membrane

187 528.07

reduction (%) compressor 1

200 451.07

installed cost (dollars, $)

6.45

hydrogen savings ($ per year) 552 258.20

net annual savings ($ per year) a

121 676.80

reduction (%)

567 379.34

compressor 2 reduction (%)

Table 8. Comparison for the HDA Flow Sheet of Lyuben et al.a Case HDA by Luyben +

Luyben

MEN


1.12

ceramic membrane area (m2) brake power (kW)

120

compressor principal

450.42

428.88

compressor 1

96.32

86.38

compressor 2

83.89

75.26

installed cost (dollars, $) ceramic membrane compressor principal reduction (%) compressor 1

133 200.00 1 001 209.56

961 776.41

282 631.08

3.94 258 472.94

reduction (%) compressor 2

8.55 252 338.35

reduction (%) total capital investment (dollars, $)

1 536 179.00 280,277.88

reduction (%) hydrogen savings ($ per year) total annual cost ($ per year) net annual savings ($ per year) a

1 584 294.30 3.13

reduction (%) annual energy cost ($ per year)

230 844.95 8.52

262 449.45 6.36 10 035.20

819 476.70

818 536.75 10 975.15


For the polymeric membrane unit installed cost, we used the data from Bouton and Luyben,9 which is $550 per m2. Konda et al.11 used an average cost of $55 per m2 for the membrane module alone, which corresponds to an installation factor of 10 to

a

253 000.00

2 886 908.85

39.95 1 760 313.76 39.02 733 460.01 39.00

total capital investment (dollars, $) reduction (%)

4 510 583.48

3 079 773.77 31.72

annual energy cost ($ per year) reduction (%) hydrogen savings ($ per year) total annual cost ($ per year)

978 541.14

535 342.20 45.29 13 260.80 1 616 342.79

net annual savings ($ per year)

HDA by

333 000.00 421 300.00

1 202 374.64

106 555.66

Data taken from refs 1 and 2.

parameter

300.00 460.00

948 997.35 5.32

reduction (%) total annual cost ($ per year)

766.00

1 002 299.51

reduction (%) annual energy cost ($ per year)

1.48

133 200.00

ceramic membrane

total capital investment (dollars, $)

Konda + MEN

brake power (kW)

installed cost (dollars, $) compressor reduction (%)

HDA by Konda

2 561 755.94

932 152.35


get to the US $550 per m2 figure used by Bouton and Luyben.9 For the cost of the ceramic zeolite membrane modules, we resorted to the research of Caro et al.,14 who reported a value of approximately $3000 per m2; however, it is common practice to use a value that is 10 times lower15 to account for the fact that the cost of this type of membrane decreases rapidly with intensified use. Therefore, to predict the cost of installed units of zeolite membranes, we use a cost for the module of $300 per m2 and an installation factor of 3.7, obtaining a cost of approximately $1110 per m2 of installed area. Investment costs are annualized using a capital charge factor of 0.351. To compute the energy consumed by the compressors, we consider a driver efficiency of 0.9 and $0.05/kWh. To compute the savings in hydrogen consumption, we take a cost of $1.12/kmol, as reported by Bouton and Luyben.9 Taking the difference (between the alternatives with and without incorporating MEN) in annualized equipment cost, and adding the annual savings in hydrogen and electricity, yields a net annual savings, which we use as a performance index to assess the success of incorporating the MEN concept in the process design procedure. Tables 7, 8, and 9 present the comparisons for the flow sheets of Douglas,1,2 Bouton and Luyben,9 and Konda et al.,11 respectively. Analyzing the tables, we see that the alternatives that include the MEN concept have a positive net annual savings in all cases: • A moderate savings of $106 555.66/yr for the flow sheet of Douglas1,2 is observed, where there is a net increase of 7.97% in capital costs (the cost of the new membrane unit is not compensated by the reduction in the compressor size) but a significant benefit from the savings in hydrogen. • An almost-negligible savings of $10 975.15/yr is observed for the flow sheet of Bouton and Luyben,9 where the 12674


Industrial & Engineering Chemistry Research increase in capital costs is barely compensated by the savings in electric energy and hydrogen consumption. • And an interesting savings of $932 152.35/yr is observed for the flow sheet of Konda et al.,11 where there is a significant 31.72% reduction in capital costs; the reduction in energy consumption is also important, and we have a small increase in hydrogen consumption. We see that the HDA flow sheet that benefited the most from the incorporation of a gas-permeation counter-current mass exchange unit was the one for Konda et al.,11 where the gas recycle has larger concentrations of hydrogen and therefore provided a larger driving force for this exchange: a significant portion of the hydrogen can be recycled through this unit, reducing recompression capital and operative costs.

6. CONCLUSIONS This paper explores incorporating the concept of mass exchange networks (MENs) as a heuristic in the early second level of the hierarchical design procedure of Douglas1,2 to synthesize a new process, when deciding the recycle and separation structure of the flow sheet. If the reaction requires operating conditions with reactants in excess or components that catalyze the reaction, which must be removed after the reactor, the MEN heuristic rule would read: “Explore the implementation of a mass exchanger between the streams exiting and entering the reactor”. And the heuristic was explored on the HDA process case study, for which several authors proposed different designs. The methodology for the synthesis of MEN is a useful conceptual tool that finds a new application if used in an initial stage of the hierarchical process design procedure, rather than at the end of a completely designed process. The first level of decisions defines the inputoutput structure of the flow sheet: raw materials, products, and processing routes. After this first level, the only defined streams are the input and output streams of the reactor: this is the point where the counter-current mass exchange should be explored, because it is a decision with a large impact on the structure of the process (this is the concept of hierarchy). If it were explored at a later level of the design procedure, it would require revising most of the decisions taken previously (about recycles, separation units, heat integration, etc.). When applied to the HDA process, this methodology generated alternatives different from the previously proposed by other authors. Implementing the mass exchange with gas permeation through a counter-current ceramic membrane unit proved to allow the recovery of some of the hydrogen present in the stream leaving the reactor without requiring recompression, but this recovery was partial, thus not eliminating the need of a further separation system. At this point, we adopted the alternative separation systems proposed by Douglas,1,2 Bouton and Luyben,9 and Konda et al.11 redesigned to perform the remaining separation task, and compared the designs proposed by these authors with those adding the mass exchanger. Analyzing the results on the examples presented qualitatively and quantitatively, we note a reduction in the size of the separation and recycle system needed after the mass exchanger, reducing its cost. The best result was obtained with the alternative proposed by Konda et al.,11 where the gas recycle has the larger mole fraction of hydrogen, with a significant reduction in 45.29% of the compression system and with net annual savings of $ 932 152.35. When adopting the separation system proposed

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originally by Douglas,1,2 and afterward by Bouton and Luyben,9 where the mole fraction of hydrogen in the rich stream is lower, the reduction of the compression system was, in the best case, 6.45% with a net annual savings of $ 106 555.66. Therefore, in the HDA case and giving the membrane technology available to perform the mass exchange, the success of implementing the MEN concept depends on how high is the concentration of the component to be transferred in the rich stream, i.e., it will work if you have a proper driving force for justifying its implementation. We compared three alternative HDA process flow sheets proposed by other authors with and without incorporating the MEN concept. We did not compare the different alternatives among themselves, because they implement different criteria for designing the separation system that affect operating conditions of the reaction, the recycle (or not) of diphenyl, etc., which was beyond the scope of this paper to analyze. What we do want to point about these alternative designs, within the context of this paper, is that the alternatives that use membrane units employ these as separators (one inlet stream, two outlet streams), and not as mass exchangers (two inlet streams contacted countercurrent). The alternative proposed by Konda et al.11 with a membrane separation in the main gas recycle could be considered (within Douglas1,2 design procedure) as selecting a membrane unit for gas separation (a technology not available at the time of Douglas’ original publications1,2) in the second level of the hierarchical design procedure (when deciding the recycle and separation structure of the flow sheet). However, the alternative proposed by Bouton and Luyben,9 with a membrane separation in the gas purge, could be considered (within the Douglas1,2 design procedure, as modified by Fischer et al.6) as selecting a membrane for process mass integration in the last level of the hierarchical design procedure (when refining the design by mass integration). In this last one, the membrane unit would be an “operator” within a sourcesink mass integration procedure (a separation unit that changes the composition of a stream, permitting that the source of hydrogen “gas purge” be acceptable for the sink of hydrogen “reactor”) and without changing the process conditions (a characteristic of low hierarchical level decisions). As further work, we think that there is a need for better conceptualization of the terms “separator” and “mass exchanger”, because they are going to be applied within different mass integration techniques (i.e., sourcesink and MEN synthesis). For example, a gas absorber column is certainly a counter-current gasliquid mass exchanger; however, if the liquid stream is an external mass separating agent, which is regenerated and recycled, then the “absorber plus mass separating agent recovery system” is a separator. Moreover, with less-theoretical commitment, membrane technology for gas permeation and, in particular, hydrogen recovery is a very active area of research and development. Probably other real-world (than our academic HDA) processes could benefit from using the MEN heuristics at an early stage of process design.

’ APPENDIX Chart A1 shows the script used in Aspen Custom Modeler V7.2 to model/describe the exchange operation, for N number of cells. It is important to notice that, for N > 20 cells, this discrete model approaches the distributed parameter model. Therefore, this is the quantity of cells that we use in our simulations in the steady state. 12675



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Chart A1

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’ AUTHOR INFORMATION Corresponding Author

*E-mails: cfi[email protected] (C.D.F.), iribarr@ santafe-conicet.gov.ar (O.A.I.).

’ ACKNOWLEDGMENT The authors greatly acknowledge financial support from Consejo Nacional de Investigaciones Científicas y Tecnicas of Argentina (through Grant No. PIP 1817). ’ REFERENCES (1) Douglas, J. M. A Hierarchical Decision Procedure for Process Synthesis. AIChE J. 1985, 31 (3), 353–361. (2) Douglas, J. M. Conceptual Design of Chemical Processes; McGrawHill: New York, 1988. (3) El-Halwagi, M. M.; Manousiouthakis, V. Synthesis of Mass Exchange Networks. AIChE J. 1989, 35 (8), 1233–1244. (4) El-Halwagi, M. M. Pollution Prevention through Process Integration; Elsevier Science: San Diego, CA, 1997. (5) El-Halwagi, M. M. Process Integration; Elsevier Science: New York, 2006. (6) Fischer, C. D.; Iribarren, O. A. Synthesis of a Mass Integrated Biodiesel Process. Ind. Eng. Chem. Res. 2011, 50 (11), 6849–6859. (7) Yampolskii, Y.; Freeman, B. Membrane Gas Separation; John Wiley & Sons, Ltd.: Chichester, U.K., 2010. (8) Pabby, A. K.; Rizvi, S. S. H.; Sastre, A. M., Handbook of Membrane Separations; CRC Press: Boca Raton, FL, 2009. (9) Bouton, G. R.; Luyben, W. L. Optimum Economic Design and Control of a Gas Permeation Membrane Coupled with the Hydrodealkylation (HDA) Process. Ind. Eng. Chem. Res. 2008, 47 (4), 1221–1237. (10) Murthy Konda, N. V. S. N.; Rangaiah, G. P.; Krishnaswamy, P. R. Plantwide Control of Industrial Processes: An Integrated Framework of Simulation and Heuristics. Ind. Eng. Chem. Res. 2005, 44 (22), 8300–8313. (11) Murthy Konda, N. V. S. N.; Rangaiah, G. P.; Lim, D. K. H. Optimal Process Design and Effective Plantwide Control of Industrial Processes by a Simulation-Based Heuristic Approach. Ind. Eng. Chem. Res. 2006, 45 (17), 5955–5970. (12) Phimister, J. R.; Fraga, E. S.; Ponton, J. W. The synthesis of multistep process plant configurations. Comput. Chem. Eng. 1999, 23 (3), 315–326. (13) Welk, M. E.; Nenoff, T. M.; Bonhomme, F. Defect-free zeolite thin film membranes for H2 purification and CO2 separation. In Recent Advances in the Science and Technology of Zeolites and Related Materials, Proceedings of the 14th International Zeolite Conference, Cape Town, South Africa, April 2530, 2004; van Steen, E., Claeys, M., Callanan, L. H., Eds.; Studies in Surface Science and Catalysis, Vol. 154; Elsevier: Amsterdam, 2004; Part 1, pp 690694. (14) Caro, J.; Noack, M.; K€olsch, P.; Sch€afer, R. Zeolite membranes— State of their development and perspective. Microporous Mesoporous Mater. 2000, 38 (1), 3–24. (15) Meindersma, G. W.; de Haan, A. B. Economical feasibility of zeolite membranes for industrial scale separations of aromatic hydrocarbons. Desalination 2002, 149 (13), 29–34.

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