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Economic estimation of various membranes and distillation for propylene and propane separation Hamid Reza Amedi, and Masoud Aghajani Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.7b04169 • Publication Date (Web): 07 Mar 2018 Downloaded from http://pubs.acs.org on March 8, 2018

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Industrial & Engineering Chemistry Research

Economic estimation of various membranes and distillation for propylene and propane separation Hamid Reza Amedi*, Masoud Aghajani Department of Gas Engineering, Petroleum University of Technology (PUT), Ahvaz, IRAN

Abstract The separation of propylene and propane is very costly due to similar physical-chemical properties. In this study, the effect of various membranes on the economic performance of the olefin unit was evaluated. The Robeson’s line cannot judge the economic performance of membrane separation systems. A precise model of co-current and counter-current hollow fiber membrane was developed and implemented in a commercial chemical engineering software to estimate the economics of the best inorganic (ZIF-8) and polymeric (6FDA-TrMPD) membranes compared with the existing distillation unit. At the same area, the co-current flows have higher recovery and the counter-current flows have higher purity. The use of the alternative polymer membrane results in a drastic increase in costs. The ZIF-8 membrane performance leads to a 30% increase in the total capital costs while the annual cost considerably decreased by 46%. When the membrane performance is appropriate in terms of selectivity and permeability, the use of membrane processes alone is more economical than the use of membrane-distillation systems. But when this is mediocre, cost savings in hybrid systems are higher than the membrane system. Also, to achieve high purity, the use of high selectivity materials is more efficient than high permeability materials.

Key Words Membrane separation, Distillation, Propylene and propane separation, Economic evaluation, ZIF8

*

Corresponding author. Tel. & fax: +98 61 35550868 E-mail address: [email protected] (H. Amedi)

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1. Introduction Lightweight olefins such as propylene and ethylene are among the most important intermediate products of the petrochemical industry. In the production of olefins, there always exists an amount of paraffin as an unreacted feed. The global need for the lightweight olefins has grown rapidly with the growth of industries and technologies. Ethylene is commercially produced in petrochemicals with natural gas feeds, and propylene is the most important byproduct. The production of ethylene from hydrocarbon sources, such as methane, ethane or naphtha, consists of four basic steps: 1) thermal cracking, 2) cooling and quenching, 3) gas compression, and 4) separation and purification. Two grades of olefins are used in industries: chemical grade with 94% purity, and polymer grade requiring a purity of 99.8%. Therefore, separation is required to achieve high purity. Low-temperature distillation is commonly used in the separation of olefins from paraffin. Figure 1 shows a simplified schematic of the separation process of light olefins and paraffin1, 2.

T = -30 0C

Light gas

T = -13 0C

Ethylene

C2 Splitter Feed

120 Trays P= 20 bar

Deethanizer 56 Trays P= 25 bar

Ethane

T = 48 0C T = 48 0C

Propylene C3 Splitter 180 Trays P= 20 bar

Depropanizer 55 Trays P= 20 bar

Propane C4+ Product

Figure 1. Cryogenic distillation process for olefin paraffin separation

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The separation of propylene and propane is very costly due to similar physical-chemical properties. For example, a distillation column with over 180 trays and a reflux flow of about 15 has been used in C3 stripping tower. In C2 stripping tower, the operating temperature is around 25 °C and the tower has more than 100 trays. In general, energy consumption is about 40% of total energy used in petrochemicals, estimated at about 1014 Btu per year. It also has an investment cost of over $50 million. In addition to operating cost and high energy consumption, creating an impact and changing product properties are considered undesirable features of this approach3, 4. The study of methods that improves the process efficiency is also of interest to many researchers. In situations where the system performance is very low, a new process should be created. For example, the bottom flow of the distillation tower is 75% propylene and 25% propane. In modern processes such as absorption and membrane, it is possible to recover propylene in this line or to use the membrane as an increase in the purity of the entrance flow of the tower5, 6. Membranes distinguish gas molecules based on their size, shape, or physical-chemical properties. Gas separation membranes are divided into two categories: polymer and inorganic according to their constituent material. Polymeric membranes, zeolite, carbon molecular sieves, metal organic frameworks (MOFs), facilitated transport and mixed matrix membranes (MMMs) account for the greatest share of gas separation studies5, 7, 8. Figure 2 compares the separation performance of propylene and propane using inorganic, rubbery and glassy membranes2, 4, 6, 9-13. Micro pore molecular sieve materials such as zeolites or organic-metal structures are capable of separation olefins and paraffin with high permeability. These porous structures have higher thermal and chemical resistance than polymers; however, the preparation of defect free inorganic membranes is not technologically feasible at large scales. Polymeric membranes are used in most of the membrane processes. Generally, rubbery polymers have higher permeability and glassy are more selective. Inexpensive polymers have optimal mechanical properties. The manufacturing process is also simpler. The mechanism of separation of gases in polymer membranes is based on solution and diffusion14. The most important problems are limited permeability and selectivity with a line called Robeson. Also, the separation performance of polymeric membrane will be reduced in the presence of heavy hydrocarbons under high pressure and temperature operating conditions15. hence, various methods have been used to improve the properties of polymer membranes in the gas separation industry such as composite materials. The Robeson’s line cannot properly judge the cost-effectiveness of membranes. Therefore, it is essential to simulate and estimate the membrane cost-effectiveness of propylene and propane gas separation.



100

C3H6/C3H8 Selectivity

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10 Inorganic Glassy

Rubbery 1 0.01

0.1

1

10

100

1000

10000

C3H6 Permeability (barrer)

Figure 2. C3H6/C3H8 experimental upper bound based on pure gas permeation data

According to previous studies and references, the best performance in both permeability and selectivity for the separation of propylene and propane gases is related to the ZIF-8 16, 17. The best propylene/propane separation performance was reported by Hara et al.18, 19 from Japan's AIST Center in 2016. In this study, the ZIF-8 membrane was constructed through the counter diffusion of its component parts into a porous tube of alpha-alumina. The best sample is related to 0.4 molar zinc concentration with 21 gpu and selectivity factor 42. Many studies have been conducted on the fabrication of a MMM containing the ZIF-8 particle20-23. In 2012, Koros et al.24 reviewed the separation of propylene and propane using the 6FDA-DAM/ZIF-8 MMM. In this study, it has been shown that there is a good compatibility between ZIF-8 and polymer matrix and, by adding it to the polymer, the properties of both permeability and selectivity separation are increased. At 35 ° C and the feed pressure of 2 atmospheres, with a loading of 48%, the selectivity and permeability nanoparticles were obtained at 31 and 56 barrer, respectively. Investigation of hybrid separation processes is growing to achieve more economic and sustainable chemical products. Unlike a simple connection between different units, hybrid systems can use their synergistic effects and overcome the operational constraints of a single unit25, 26. Among membranes used in the hybrid distillation-membrane system, vapor permeation (VP) and pervaporation (PV) are more common than other membrane processes. Researchers have studied the effect of various membrane arrangements and the operational parameters for the separation of lightweight27-30. Kookos et al.31 have shown that a 17% reduction in propylene and propane separation is possible with hybrid systems. Ben Ali et al.32 have shown that the initial and annual cost can be reduced by 67% and 14% respectively. They used a facilitated transport


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membrane containing silver ion with a permeance of 328 gpu and a propylene/propane selectivity of 374. But in practice, the implementation of these membranes is impossible. Especially silver ions can react with impurities in the feed and do not guarantee the long-term stability of the membranes. Facilitated transport membranes containing ionic liquids such as BMImBF4–Ag+ have shown outstanding performance, with permeability and selectivity of around 7,000 barrer and 70033. The practical mathematical model of these membranes can be found in the references34, 35. Moganti et al. 36 found that a suitable hybrid system should have a flow ratio of feed to the membrane area of about 0.1, in propylene and propane separation. It is also possible to reduce the trays by 30%, and the annual cost is reduced by 20%. Caballero et al.37 studied the effect of permeate pressure and found that with decreasing flow pressure, the surface of the membrane is reduced, but instead the compressor costs increase. Design and optimization of the distillation membrane process are very diverse. Membranes can be used sequentially before or after distillation. Pedram et al.38 have studied the effect of different configurations of hybrid distillation-membrane system on energy consumption in propylene and propane separation. In the same conditions without changing the number of trays in the tower, the series configuration has the best energy reduction. Also, Motelica et al.39 studied the effects of various membrane configuration. the series system has the highest performance under the same conditions. In series arrangement, the total system cost reduction is about 61%, as in the Cascade system, this cost reduction is about 54%. The design and optimization of the hybrid systems are more complex and flexible than a single separation unit so that there is no general approach to design and optimize them. Therefore, for each specific separation case, the probable states should be evaluated by simulation. The main objective of this study is to investigate the effect of various membranes as well as to examine the capability of membrane technology for propylene and propane separation. Also, as mentioned above, The Robeson’s line cannot judge the economic performance of membrane. In most simulations, up to now, membranes are not capable of mass production, as well as operating and feed conditions differ from petrochemical olefin units. In this study, a precise comprehensive model of the hollow fiber membrane was developed in the co-current and counter-current flow. Next, the model was implemented in a common commercial chemical engineering software product. To ensure the validity of the model and reliability of its implementation, two references were compared in terms of model simulation results and experimental data. To compare the performances of the membrane under distillation conditions, the olefin-paraffin separation unit was simulated in Hysys, and the feed data was used to simulate the alternative membrane units. The performances of the best inorganic and organic membranes in the references were compared to those of the distillation unit after simulation and economic estimation procedures.



2. Models and methods The first step in modeling and simulating of the engineering processes is gaining knowledge of the theories, models, and relations governing the process. In fact, a simulation engineer must completely analyze the existing models before taking any measure. The details of this analysis reveal the applications of the models, their weaknesses and strengths, and the information required for using those models26, 40. Hence, it was tried to simulate a precise master model of the hollow fiber membrane in a common chemical engineering software product and then estimate the economic performance of the best membranes to decide on their current status. 2.1. Base case distillation unit In this study, a real distillation system is simulated based on the petrochemical olefin unit data. This unit has two similar distillation towers, with a total of 196 trays, a reboiler, and a condenser under reflux ratio 16. The feed contains 91.2% propylene, 3.8% propane, and less than half a percent acetylene, butane and heavier compounds than hexane. The tower with a very high energy flow (about 6500kcal/s) is capable of purifying propylene up to 99.8%. This is a purity suitable for polymeric reactions. 2.2. Membrane unit model In the membrane gas separation industry, the hollow fiber modulus is commonly used and, in some cases, a spiral wound module is used while tubular or frame and plate modules are rarely used. The surface/volume ratio is high for both hollow fiber (500-9000 m2/m3) and spiral module (200-800 m2/m3); however, hollow fiber module is more cost-effective so that these modules cost 2-25$/m2. This is far lower than spiral modules, which cost 10-100$/m2. Hollow fiber membranes usually have fibers with the diameter about 50-500μm. Similar to the high volume/area ratio, the diameter has high area/length ratio, too so that one-kilometer fiber is required to create 1m2 module membrane. Hollow fiber membrane fabrication methods are divided into two general methods: 1) Solvent Tissue Method 2) Melted Polymer Tissue. These membranes can also form a dense homogeneous structure and form a microscopic structure with a very thin selective layer. The dense layer can be covered integrated or separated. Finally, the fibers are classified and squeezed to be used for membrane modulation41. Since the 1950s, many models have been developed for membrane processes for the separation of two or more gases. The basis for categorizing models is the flow pattern. Initial models are based on the complete mixing of gases; however, in practice, these models were not accurate enough. Stern et al.42, 43 proposed a model for all flow modes in separating two-component gas separation. Pan et al.44 considered the pressure drop in his model. Then, Shindo et al.45 developed


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these models for multi-component systems. an excellent review of all existing gas separation models was presented by Li et al.46. The numerical solutions were proposed by Marriott and Sorensen47 for the simplified models. In the next years, the researchers examined the numerical solution methods of these equations with acceptable accuracy and volume of computations. In 1998, Coker and Freeman48 presented a very precise numerical solution to hollow fiber modules with two co and counter flows. The numerical solutions are more accurate than others and capable of simulating membranes at high stage cut, multi-component mixtures, and various permeability rates. There are many models available to describe the gas transfer through the membrane. In gas separation membrane, mass transfer governs the process, which includes partial and total mass conservation laws. Due to the maximum driving force and the smallest area required, most simulations are based on counter-current flow. The flow of fluid in the membrane can be either co-current or counter-current. In addition, the feed is placed in one of the following forms: 1) shell-side feed; 2) bore-side feed. The former state with counter-current flow was considered in this research. Figure 3 shows the schematic of mass transfer in the counter-current hollow fiber membrane.

Permeate Permeate Feed

Feed dA=π Do Nf dz

Qi(PLxi-Pvyi) L, xi PL

V, yi PV

Support Layer Selective Layer

Retentate

Axial (z) Radial (r) Retentate

Figure 3. the schematic of mass transfer in the counter-current hollow fiber membrane

The assumptions governing the model are as follows42-44, 49: 

Any axial influence of feed flow is neglected.





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Mass transfer in the radial direction is accomplished by the mechanism of solution and diffusion.



The permeability is constant and the effects of temperature and pressure are ignored.



The pressure drop in the feed section is also neglected.



The build-up pressure of permeate flow is obtained by the Hagen–Poiseuille equation. The equation is based on the analysis of the laminar flow through the channels with porous walls.



All hollow fibers have a uniform shape and hollow fiber membrane deformation has been ignored in pressure conditions.



The mass transfer resistance is only related to the selective layer and support layer resistance has been ignored.



The membrane module is stable under steady operational conditions.

As it is shown in diagrams, feed components from the shell side with higher permeability enter the hollow fiber. According to the permeability definition, the flux passing through the membrane for the ith species is equal to Equation 1.

Ji = Pi (pL xi − pV yi )

(1)

The surface element in the modules can be defined according to equation 2.

dA = Area ∗ dz/L

(2)

z is the flow direction and L is the effective membrane length. Equation 3 is used to calculate the total effective surface of the membrane with N hollow fiber and Do external diameter.

Area = πDo Nf L

(3)

Table 1 shows the differential equations governing the counter-current flow. In the mass conservation law for the ith component (equation 4 and 5), the molar variation per unit area is equal to the amount of its flux due to lack of reaction (production/consumption) and steady conditions. Similarly, the same conditions exist in the general mass equation and the total flow differential equal to the sum of the inlet and outlet fluxes. As stated earlier, in Equation 8, feed pressure drop is ignored. Flow pressure drop is calculated based on the Hagen–Poiseuille equation, where Di is the internal diameter of each fiber. For modeling co-current flow, the


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governing equations are similar to the counter current flow. In the retentate flow, since feed current is long the Z axis, the differentials are applied with the negative signs. Table 1. The governing equation of the counter-current flow

Name

Equation

Condition

Permeate side mass balance for component i

d(V. yi ) = Ji dA

@A = 0

dyi =0 dA

(4)

Retentate side mass balance for component i

d(L. xi ) = Ji dA

@A = 0

xi = xf

(5)

Total mass balance of permeate side

d(V) = ∑ Ji dA

@A = 0

V=0

(6)

Total mass balance of retentate side

d(L) = ∑ Ji dA

@A = 0

L=F

(7)

Permeate pressure

d(pV ) 128 μRTV = − dZ πNf D4𝑖 pV

@z = 𝑙

pV = pv 𝑜

(8)

N

i=1

N

i=1

2.3. Simulation and economic calculations The numerical solution method is based on Cooker's work48 in Supporting Information file (section1). The equations were solved in MATLAB and Hysys. In addition to user-friendly environment and support for thermodynamic data, Hysys was used due to the adaptability of membrane modulus with other equipment. The differential and algebraic equations of the membrane are defined as a custom model based on the Visual Basic language, which has a good combination of speed and stability. Optimization has been done in two ways: 1) using the genetic algorithm of MATLAB software 2) using the case study of Aspen Hysys. In the first method, Hysys was integrated with MATLAB in a client– server relationship. Then, the cost objective function is minimized using the optimization tool and the genetic algorithm. The values obtained are the same in both ways. But, the second method is recommended due to its better run time. Peng-Robinson and Soave-Redlich-Kwong equations are widely used in the petrochemical industry to accurately describe the system of light and non-polar hydrocarbons. In many simulations, SRK has been used to describe light hydrocarbons, hence it is chosen as the Equation



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of State and the thermodynamic package to estimate the behavior of gases. The feed and operating conditions are similar to the base case. In order to prevent the condensation of heavy components, the process flow is heated to 60 oC by E-101 heat exchanger. At each stage, the difference in pressure between the two sides of the membrane is the maximum value, so that the pressure of the permeate flow adjust to 1 bar. Heavier compositions of Pentane are also simulated with the same hexane; this simplification results in a little error in the calculation because of its very low percentage (about a few ppm). On the other hand, one of the disadvantages of using the parallel arrangement is that the maximum flow to the membrane unit is reduced, which is due to the hydrodynamics of the tower. But in the series configuration, all feed is in contact with the surface of the membrane, hence the combination of membrane and distillation is in series. Aspen Icarus (another Aspen Tech software), which is fully integrated with Hysys, has been used for economical estimation. Other parameters, such as factory location, common currency, working hours per year, steam and cooling costs, etc., are based on the default software settings. The economical estimation of membrane modules has been done from references. Table 2 shows how to calculate the costs associated with the membrane modules. Table 2. The membrane module costs50-52 Name

Cost

Total membrane module cost [USD]

10 $/ft2

Membrane replacement cost [USD]

2.5 $/ft2

Installed membrane module Cost [USD]

1.12* membrane module Cost

Project contingency [USD]

0.2* Installed membrane module Cost

Total capital cost [USD]

Installed module Cost + Project contingency

The most important parameters affecting the performance of a membrane process are the membrane permeability and selectivity, so we need to increase permeability and membrane selectivity in order to increase the recovery and purity of the products respectively. In the following, in four different scenarios, we study the performance of the best polymer membrane and the best inorganic membrane in the simulation and economical estimation of propylene/propane gas separation unit.


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3. Results and discussion 3.1. Simulation of industrial unit Specifications of feed flow, matter, and energy are shown in Figure 4. After simulation, the composition and properties of flows completely matched with the unit's PID data. The propylene purity of overhead flow is 99.77 % with a flow rate of 462 kmol/h. In this condition, the tower has the ability to recover 86.7% of the propylene and stage cut is about 79.3%. The result details of the economic estimation are in the Tables S1-S2 (supporting information).

Figure 4. the simulation of the propylene-propane separation unit

3.2. Validation of membrane model In this section, the accuracy of the simulation is first examined in two sections and then compared with the experimental data. The effective parameters in the membrane performance are as follows: the permeance of gases, operating temperature and pressure of feed and permeate flow, geometric dimensions and flow pattern in the membrane. In membrane simulation, gas permeation is the only parameter that is related to material properties, and other parameters do not depend on the membrane material. In other words, the model is independent of the membrane material and penetrating gases and their interaction is included in the rate of permeation. Hence, the validation of the model is correct using different gases and membranes. Also, in this study, permeability has a constant value and the effect of temperature and pressure has been ignored. This assumption is reasonable because the laboratory conditions in which the permeability of the components is measured is close to the operating conditions.



To ensure the accuracy of equations and results, Cooker and Freeman48 work were simulated in Hysys. In that study, feed was on the shell side and the following specifications were taken into account: permeate pressure, 4.24 bars; the retentate pressure, 7.9 bars; feed temperature, 50°C; effective fiber length, 80 cm; external diameter, 300 microns; internal diameter, 150 microns; and the number, 500,000. Figure S2 (supporting information) shows the implementation of hollow fiber module in Aspen Hysys. Figure 5 shows the system re-implementation for comparison. As it can be seen, the numerical solution of the equations is correctly implemented. To compare the model with experimental data, the empirical study by Pan et al.44 was used at semi-industrial scale. The asymmetric cellulose acetate membrane was designed to dispose of hydrogen from ammonia at the purge gas. Conventionally, the gas mixture was at the pressure of 70 bars, the temperature of 25 ° C. It consisted of 52% hydrogen, 25% nitrogen, 20% methane, and 3% ammonia. The hydrogen/ acetone selectivity is more than 100 at the cellulose membrane. Other design parameters and feed specifications are available in the article. Figure 6 shows the comparison of lab data and the numerical model. The proper matching is found between the experimental data and the model in a wide range of stage cut, especially at the low value. There is a relative error between model and experimental data at the high hydrogen stage cut (i.e. high removal factor), especially in the co-flow over 5%, associated with the exclusion of diffusion backflow in the porous support membrane. In fact, with more separation, a greater concentration gradient occurs within the membrane. In most industrial applications, hydrogen recycling is less than 95%. At this recovery, the required area of membrane is low and the purity of hydrogen is greater than 90%. The error between the model and the experimental data in this range is less than 2%.


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100 95

H2 Purity (%)

90

Pf/Pp=5.3

85

Pf/Pp=1.8

80 75 70 65 0

20

40

60

80

100

H2 Recovery(%)

Figure 5. comparison of implemented model data (line) and Coker-Freeman data48 (dot)

100

80 90 60

80

H2 Recovery (%)

100

H2 Purity (%)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


40

0.3

0.35

0.4

0.45

0.5

0.55

0.6

Stage Cut Exp. Purity(Counter-Current) Model Purity(Counter-Current) Exp. Recovery (Counter-Current) Model Recovery (Counter-Current)

Exp. Purity (Co-Current) Model Purity (Co-Current) Exp. Recovery (Co-Current) Model Recovery (Co-Current)

Figure 6. comparison of implemented model and experimental data44



3.3. Polymeric membrane separation As mentioned in Introduction, the best propylene/propane separation performance was obtained using polymeric membranes based on the koros et al. 53 results in polyimide membranes. In that study, propylene and propane permeability was reported about 30 and 2.7 barrers using a 6FDATrMPD dense membrane with a thickness of 20 microns. In practice, to increase the permeation, the membrane thickness can be reduced to about 4 microns. it should be noted that with a further reduction in thickness, the support mass transfer resistance cannot be neglected and affects the membrane performance54. In this section, the simulated membrane has the same performance with a thickness of 4 microns. The permeability of other components is estimated based on other studies55. It should be noted that the sum of the other components except propylene and propane in the process feed is less than half a percent, so it can be possible to ignore the estimation error. One-Stage At first, their performance was checked to determine the type of co-current or counter-current flow. In the hollow fiber membrane, the outer and inner diameter is 300 and 150 microns respectively, and the useful length of the membrane is one meter. In Table 3, the purity and recovery of propylene are given in percentage based on the change of required area. When the surface of the membrane is low, the behavior of both flow is the same, and it can be said that the membrane performance is independent of the type of modulus and flow, and it is considered as dead-end flow. By increasing the membrane surface according to the conditions, these changes slightly increase, so even in a membrane with 60,000 square meters, the recovery and purity difference can hardly be expressed. In comparing co-current flow with the counter-current flow, it can be generally said that in the same area, the co-current flows have higher recovery and the counter-current flows have higher purity. On the other hand, the assumption of a plug flow in the counter-current flow is more reliable than the co-current flow. And this phenomenon reduces membrane performance in co-current flow. In summary, due to the relative advantages of the hollow fiber modulus at counter-current flows compared to the co-current flow, hollow fiber membrane with counter-current flow has been used. In Figure 7, the performance of this membrane has been shown based on surface change. As it is clear, with increasing the membrane surface, the amount of flow passed through the membrane (i.e., the stage cut) increases, while the purity of the propylene decreases. To reach 462 kmol/h, the flow passed through the top of the distillation tower, requires a surface of 36550 m2, in which case the outlet mole fraction of propylene is 97.9%. Therefore, at one stage of membrane separation for each surface, 99.8% purity cannot be reached by flow 462 kmol/h and an increase in the number of separation steps


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is required. The flow properties of the membrane input and output are shown in Figure S3 (supporting information). Table 3. The performance of hollow fiber membrane in co-current and counter-current flow Hollow Fiber Counter-Current

Hollow Fiber Co-Current

Y (%)

Recovery (%)

Y (%)

Recovery (%)

0.1

98.66

0.00

98.66

0.00

1000

98.67

2.60

98.67

2.60

20000

98.61

50.28

98.62

50.24

40000

97.60

90.36

97.61

90.28

60000

93.97

99.88

93.99

99.69

100

100

98

80

96

60

94

40

92

20

90

0

10000

20000

30000

40000

50000

0 60000

Area (m2)

Figure 7. the performance of single stage hollow fiber membrane

Two-Stage


Recovery (%)

A (m2)

Purity (%)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60



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As mentioned in the previous section, propylene purity and recovery were not achieved for each surface, so another membrane unit was added to the previous stage. To maintain the pressure and adjust the temperature, the compressor K-100 and the E-100 heat exchanger have been used. To find the area in each of the membranes, an optimization problem with the conditions of equations 9 is established.

nin At = A1 + A2 Subject to: Purity > 99.75 StageCut ≥ 79.3

(9)

The objective is to calculate the minimum required surface so that the desirable purity and flow rates of the output are maintained. To solve the optimization issue, the data can be sent simultaneously to MATLAB software and it can be solved using the optimization tools of the software, such as the genetics algorithm. But since the area of each membrane is 50 square meters, case study environment of Hysys software was used to search directly and find the optimal area. Due to the conditions and low selectivity of the membrane, in this case, similar to the one-stage state, for each area, purity and recovery conditions cannot be met at the same time. Figure S4 (supporting information) shows the flow properties in a two-stage membrane. If the first membrane surface is 36,000 square meters and the second membrane surface is 25,000 square meters, the highest concentration is achieved with 99.4%. Table 4 summarizes the performance of the two-stage membrane in the best possible state. The area for the first and second membranes are calculated in such a way that, in addition to the permeate flow of 462 km/hr., the highest value of propylene mole fraction is also obtained. In order to reach a flow rate of at least 462 kmol/h, we need to use a surface of 43,000 and 33,000 square meters. In this case, the composition rate of output is equal to 99.1%. In any case, the concentration of 99.8% is not achieved. Therefore, it is necessary to increase the separation stages. Table 4. Two-stage membrane performance in the best possible state state

A1 (m2)

A2 (m2)

Y (%)

The highest propylene composition

43000

33000

99.1

472.3

The minimum required area

38000

33000

98.9

463.3


Q (kmol.hr-1)

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Four-Stage Similar to single-stage and two-stage state, purity and recovery cannot be achieved at the same time using the three stages of membrane separation. In four steps, the optimization problem was calculated to be 48000, 37400, 32500, and 31000 square meters respectively for four membranes, to achieve a composition of 99.78 propylene and a flow rate of 462 kmol/h. Figure 8 shows the schematic of the process. After simulation and design, the economical estimation of equipment and energy has been done by Aspen Icarus software. The costs of the propylene, propane separation using a polymer membrane unit are shown in Tables S3-S4 (supporting information).

Figure 8. the input and output flow conditions of four-stage polymeric membrane

3.4. Polymeric membrane-distillation system A complete replacement of the membrane with distillation is not possible at the moment due to the poor performance of the best possible polymeric membrane. In this section, the design, simulation and economical estimation of the process of polymeric membrane combined with distillation are discussed. If the membrane is placed after distillation (whether in the direction of the condenser or in the direction of the reboiler or in the middle of the tower), the inlet flow of membrane decreases hence in this condition requires a higher area than the condition that the membrane is located before the tower. The membrane area required in this situation is more than the condition of the membrane is located before the tower. As a result, the effect of different distillation-membrane configuration is not evaluated, and in all cases, the tower is placed after the membrane. The optimization issue is described in equation 10.



min Total Cost = MemCost + ElectCost + SteamCost + CWCost Subject to: Purity > 99.75

(10)

StageCut ≥ 79.3 In this part, the purity and recovery are obtained by distillation tower in each input condition by changing the amount of heating and cooling and reflux flow to the tower. Adjustable parameters are the feed trays and the optimum membrane surface. The goal is to reduce the total operating costs resulting from heating and cooling as well as the fixed cost of the membrane module. The effect of different parameters on the current cost of the polymeric membrane- distillation system is shown in Figure 9. After the solution, the optimal surface was found 13450 m2 and permeate and retentate flows were optimally connected to the tower in trays 149 and 35 respectively. In Figure 10, the mass composition and energy conditions of the flows are indicated. Capital and

12 10 8 6 4 2 0 0

5000

10000

15000

20000

25000

30000

10 9 8 7 6 5 4 3 2 1 0 35000

Cost (105 $/year)

annual costs related to the simulated unit are specified in Tables S5-S6 (supporting information).

Cost (106 $/year)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 18 of 29

Area (m2) Total Cost

SteamCost

MemCost

ElectCost

CWCost

Figure 9. Effect of different parameters on polymeric membrane costs


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Figure 10. the input and output flow conditions of distillation-organic membrane system

3.5. Inorganic membrane separation In the third scenario, the best performance of the best inorganic membrane has been investigated. As explained in the introduction, the performance of inorganic membranes is much better than polymer membranes, and their application in the gas industry is limited because of their high price and brittleness. By assuming that the limitation and restriction will be removed, the performance of this type of membrane has been studied in the following. The ZIF-8 membrane constructed by Hara has shown a remarkable performance so that propylene permeance and its selectivity to propane were reported 21 gpu and 42 respectively at the equimolar feed condition and 25 °C. As in the previous scenario, the permeability of other species is estimated based on other references55-57. One-stage In Figure 11, the purity and recovery changes are shown by increasing the surface of the membrane at one stage. The highest purity is 99.5 obtained zero area, in which the output flow of the membrane is also zero. By increasing the area up to about 15000, stage cut and recovery have risen sharply to over 90%. From this point onwards, with increasing area for recovery of 10% propylene, purity is reduced significantly. In order to have a flow rate of 462 kmol/min, we need a minimum surface of about 12500 square meters, since, at one stage of separation, the purity equivalent to this level is 99.3%, and so we definitely need to have two stages for separation. In this case, as in 3.3, an optimization problem is raised. The goal is to optimize the minimum total



surface area of the membrane in two stages so that the purity and recovery are met. The condition of the problem is similar to the base case. After optimization, the minimum surface area is 23850 square meters, with an effective surface area of 13450 m2 and a second membrane surface of 10400 m2. In this case, the purity was 99.99% and the removal coefficient was 79.3%. The properties of input and output flows is shown in Figure 12. After the design, economic results have been shown in Tables S7-S8 (supporting information).

100

100

99

60 97 40 96 20

95 94 0

10000

20000

30000

0 40000

Area (m2)

Figure 11. the performance of single stage inorganic membrane


Recovery (%)

80

98

Purity (%)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 20 of 29

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Figure 12. inlet and outlet flow conditions of two stage inorganic membrane

3.6. Inorganic membrane-distillation system As shown in Equation 10, here we deal with an optimization issue. The objective function is to minimize the costs of membrane purchase and maintenance, steam supply, cooling water flow and electricity. Providing purity and stage cut are the conditions governing the problem. Variables include the surface of the membrane and the location of the feeds. In Figure 13, the process of changing the cost of the membrane, steam, water dispenser, and electricity is shown in terms of the change in the membrane surface. Steam costs are $ 1 million and other costs are $ 100,000. Therefore, the total cost is affected by the amount of steam used in reboiler. After solving the problem, the optimal area was found around 8350 m2 and the flow of permeate and retentate were connected to the tower in trays 165 and 33, respectively. In Figure 14, the condition of mass flow and energy is specified. The purchase costs of equipment installation, energy consumption and total current and fixed costs are specified in Tables S9-S10 (supporting information).


Page 22 of 29

12

12

10

10

8

8

6

6

4

4

2

2

0 0

5000

10000

15000

Cost (105 $/year)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Cost (106 $/year)


0 20000

Area (m2) Total Cost

SteamCost

ElectCost

MemCost

CWCost

Figure 13. Effect of different parameters on inorganic membrane costs

Figure 14. inlet and outlet flow conditions of distillation-inorganic membrane system


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3.7. Comparison A range of membranes with different selectivity and permeability was investigated in Aspen Hysys software and compared with the conventional distillation process. The summary of the comparison of the costs of different states with the distillation operation results is shown in Table 5 and Figure S5 (supporting information). In the first scenario, the use of the alternative polymer membrane results in a drastic increase in costs. The cost of purchase and installation of the organic membrane is higher than distillation, and only the cost of the steam used for the polymer membrane is lower than distillation. However, the annual costs of the unit are higher than the distillation annual costs because of the high membrane maintenance costs. Hence, complete replacement of the polymer membrane with distillation is an out-of-reach goal in this industry. In the second scenario, the performance of a hybrid system of the best polymer membrane is compared to distillation. In this state, the total capital costs are approximately doubled and a onemillion-dollar decrease in the total operating costs is observed per year. In the third scenario, the ZIF-8 membrane performance leads to a 30% increase in the total capital costs but reduces the annual costs by 5 million dollars. When the membrane performance is appropriate in terms of selectivity and permeability, i.e. the third scenario, the use of membrane processes alone is more economical than the use of membrane-distillation systems. But when this is mediocre, for example the second scenario, cost savings in hybrid systems are higher than the membrane system. In the last scenario, the combination of the ZIF-8 membrane with distillation results in a 50% increase in the capital costs and a 4-million-dollar decrease in the annual current costs. If the initial costs are not taken into account, the capital cost of the fourth scenario is lower than the third scenario, but its annual cost is still higher due to steam consumption. Table 5. A summary of capital and operational costs for different scenarios

Name (million $)

DISTILLATI ON

ORGANIC MEMBRANE

ORGANIC MEM. /DIST.

INORGANIC MEMBRANE

INORGANIC MEM. /DIST.

Equipment cost [USD]

3.0

23.9

6.2

5.7

5.4

Total installed cost [USD]

4.1

26.8

7.8

6.5

6.8

Total capital cost [USD]

9.6

43.2

16.7

12.8

14.9

Total utilities cost [USD/year]

7.8

5.3

6.4

2.6

4.1



Total operating cost [USD/year]

9.7

12.1

8.9

Page 24 of 29

5.2

6.2

4. Conclusions The separation of propylene and propane is very costly due to similar physical-chemical properties. Membrane technology can be a suitable option for replacement. The Robeson’s line cannot judge the cost-effectiveness of membranes. In this study, a precise comprehensive model of the hollow fiber membrane was developed in the co-current and counter-current flow. In comparing co-current flow with the counter-current flow, it can be generally said that in the same area, the co-current flows have higher recovery and the counter-current flows have higher purity. Next, the model was implemented in a common commercial chemical engineering software product. Using four different scenarios, the performances of the best inorganic (ZIF-8) and polymeric membranes (6FDA-TrMPD) in the references were compared following the simulation, design, and economic estimation with the distillation unit. In the first scenario, the use of the alternative polymer membrane results in a drastic increase in costs. In the polymeric membranedistillation system, the total capital costs are approximately doubled and utility costs fell by 18% due to reduced steam consumption. In the third scenario, the ZIF-8 membrane performance leads to a 30% increase in the total capital costs while the annual cost considerably decreased by 46%. In the last scenario, the combination of the ZIF-8 membrane with distillation results in a 50% increase in the capital costs and a 36% decrease in the annual costs. When the membrane performance is appropriate in terms of selectivity and permeability, i.e. the third scenario, the use of membrane processes alone is more economical than the use of membrane-distillation systems. But when this is mediocre, for example the second scenario, cost savings in hybrid systems are higher than the membrane system. When the goal is to achieve high purity, it is better to use higher selectivity materials

instead of using higher permeability materials. Therefore, it is

recommended to increase membrane selectivity instead of permeability.


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Industrial &Graphical EngineeringAbstract Chemistry Research

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EH S1

Permeate M

Feed

Ec

Compressor T

Qc (kcal/s) Qr (kcal/s)

Distillation 6426 6494

Qc

HeatExchanger

Propylene outlet

S2 Hybrid 3260 2771

SideStream

Qr Distillation Tower HF Membrane

Retentate

ACS Paragon Plus Environment

Propane outlet

Economic Estimation of Various Membranes and ... - ACS Publications

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