Separation of Carbon Dioxide and Methane Mixture by an Adsorbent

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Separation of Carbon Dioxide and Methane Mixture by an Adsorbent/Membrane Hybrid (AMH) System using Zeolite 5A Pellets and FAU-Zeolite Membrane Yun-Jin Han, Jun-Ho Kang, Hee-Eun Kim, Jong-Ho Moon, Churl Hee Cho, and Chang-Ha Lee Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.6b04608 • Publication Date (Web): 08 Feb 2017 Downloaded from http://pubs.acs.org on February 9, 2017

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Industrial & Engineering Chemistry Research

Separation of Carbon Dioxide and Methane Mixture by an Adsorbent/Membrane Hybrid (AMH) System Using Zeolite 5A Pellets and FAU-Zeolite Membrane Yun-Jin Han1, Jun-Ho Kang1, Hee-Eun Kim1, Jong-Ho Moon1,2, Churl-Hee Cho3, Chang-Ha Lee1* 1

Department of Chemical and Biomolecular Engineering, Yonsei University, Seoul, Korea 2

3

Low Carbon Process Laboratory, Korea Institute of Energy Research, Daejeon, Korea

Graduate School of Green Energy Technology, Chungnam National University, Daejeon, Korea Tel. 82-2-2123-2762, E-mail: [email protected]

Abstract The permeation and separation of a CO2/CH4 mixture in an adsorbent/membrane hybrid (AMH) system, comprising zeolite 5A pellets packed into a tubular-type FAU-zeolite membrane, were compared to those in the FAU-zeolite membrane-only system. The separation factor of the AMH system was higher than that of the FAU-zeolite membrane under the same condition because of the difference in the propagation of each gas and the radial contact efficiency. The integration of two equilibrium separating media (zeolite 5A pellets as a bulk separator and FAU-zeolite membrane as a purifier) led to improved separation efficiency and reduced performance-fluctuation by varying input conditions because of the strong interaction between zeolite 5A and CO2. However, the permeation flux in the AMH was reduced due to the gas adsorption by the packed adsorbents. The dynamic behavior in the AMH system was well-predicted by integrating the mathematical models for an adsorbent-packed bed and a tubular-type membrane.

Keywords: hybrid system, FAU zeolite membrane, zeolite 5A, carbon dioxide, methane

1

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1. Introduction Bio-gases generated from biomass, mainly methane and carbon dioxide, can be used as energy sources. Furthermore, a large amount of methane is emitted from landfills and waste deposits, accounting for approximately 11−12% of the global anthropogenic CH4 emissions.1 Therefore, the recovery of CH4 from various bio-gases, including landfill gas (LFG), is a prospectively attractive method for addressing environmental problems and simultaneously generating energy.2, 3 Various membranes based on polymers, zeolites, silica, and carbonized materials have been reported for the separation and/or filtration of gas mixtures containing CO2 and CH4,4-16 and various polymeric membranes with high permeability and selectivity for CO2 and CH4 have been reported.11, 16-18

However, enhancing the durability of polymeric membranes under adverse and rigorous

conditions is the main challenge that needs to be addressed to overcome the drawbacks of aging and plasticization of these membranes, which leads to a decline of the separation performance with time.8, 17, 19

Therefore, although polymeric membranes possess excellent performance potential, further

enhancement of their separation efficiency and thermal/chemical stability remain worthwhile challenges to augment the efficiency of separation processes. Inorganic membranes have been suggested as an alternative candidate for gas separation based on their ability to withstand high temperature and pressure conditions.19 Carbon molecular sieve membranes can yield high selectivity in mixtures with a high kinetic difference because of the molecular sieving effects.8,

15, 20

Molecular kinetic control has also been documented for a

methyltriethoxysilane templating silica (MTES) membrane based on the pore size of the thin templating silica layer.12, 21, 22 Because kinetic separation by molecular sieving effects is affected by the amount of adsorbate (adsorption amount) in porous materials, information regarding the kinetics and equilibrium in pores is important for elucidation of the separation mechanism of inorganic membranes.19, 22, 23 To achieve the desired gas separation efficiency, combined processes employing two separate units, such as membrane and adsorption units or membrane and distillation units, etc., have been evaluated in a number of studies. In combined processes, one unit works as a bulk separator and the other plays the role of a purifier. However, such combined processes generally have greater space requirements and the capital and operating costs should be carefully evaluated. On the other hand, hybrid processes, in which two separate processes are integrated into a single unit, is attractive due to the possibility of significant improvements in terms of efficiency by reducing the equipment size, energy consumption, and product losses. This principle has been successfully applied to membrane reactors and sorption-enhanced water gas shift reactors, where the reaction and separation steps occur simultaneously. As a similar concept, an MTES membrane packed with zeolite 5A has been applied to hydrogen separation,24 where it was found that the combination of a kinetic separator and an 2


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equilibrium separator was more efficient for a mixture with high kinetic difference than for a mixture with high equilibrium difference. However, due to the complexity of operation in such hybrid systems, detailed study should be carried out to understand the separation mechanism and behavior. It has been reported that zeolite membranes offer higher selectivity for carbon dioxide in comparison to other membrane materials and are also mechanically and thermally stable.25,

26

Therefore, zeolite membranes are particularly applicable in hybrid systems in which more adsorptive molecules are produced during the separation process.10 In addition, among various commercial adsorbents, zeolites have earned distinction as more efficient adsorbents for the separation of CO2/N2 and CO2/CH4 at relatively low pressure,27 and zeolite 5A was shown to be effective in the separation of a CO2/CH4 mixture.3 Therefore, an adsorbent/membrane hybrid (AMH) system employing a zeolite membrane with a zeolite pellet is prospectively useful for combining the bulk separation and purification steps in a single unit and for performing cyclic processes, such as a pressure swing adsorption (PSA) using multi-AMH units. The AMH system can be consisted of two ways: one is having adsorbents packed inside a membrane and the other is having adsorbents packed outside the membrane. Depending on the AMH configuration, the role of each substance is changed to either as a bulk separator or as a purifier. However, when a sweeping gas is demanded for a system, the AHM configuration, adsorbents packed outside the membrane, is not recommended because the sweeping gas works as a purge (desorption) gas for the adsorbents. In this study, an AMH system is developed with the objective to achieve enhanced membrane performance for the separation of a CO2/CH4 binary mixture (50/50 vol.%). A FAU-zeolite membrane packed with zeolite 5A pellets, which integrates two equilibrium separators, is selected as the model AMH module. Bulk separation by zeolite 5A and purification by the FAU-zeolite membrane are expected to occur simultaneously in this AMH system. Permeation and separation experiments are executed from unsteady- to pseudo-steady-state permeation of a CO2/CH4 mixture, and the results are compared to those obtained with the FAU-zeolite membrane only. The effects of temperature, pressure, the sweep gas, and the retentate gas flow rate on the separation performance are also investigated for the AMH system and the membrane-only system. A mathematical model of the mass and energy balances of a packed bed system that considers the axial and radial directions is integrated with a tubular membrane model using the generalized Maxwell–Stefan (GMS) and dust gas model (DGM). To illustrate the dynamic behavior of the AMH system, the model is applied to analysis of the permeation and separation behavior of a CO2/CH4 mixture under unsteady- and pseudo-steady-state conditions.

2. Mathematical model of the AMH system

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To understand the dynamic behavior of the AMH system, the following assumptions were adopted in the mathematical model of the packed adsorbents: (i) the gas phase behaves as an ideal gas mixture, (ii) thermal equilibrium between the adsorbents (or membrane) and bulk flow is assumed, (iii) the flow patterns are described by the axially and radially dispersed flow model, (iv) the pressure drop along the bed is negligible, and (v) the friction between the molecules is less important than friction 28

with the membrane wall.

Model for packed zeolite 5A pellets The component and overall mass balances for the bulk phase in the packed zeolite 5A pellets are given by the following equations:29, 30 Component mass balance of the packed bed:

∂ci  1− εt  ∂ qi +  ρp ∂t  εt  ∂ t ∂  ∂c  1 ∂  ∂ci  ∂(uzci ) ∂(urci ) = DL  i  + Dr − − r ∂z  ∂z  r ∂r  ∂r  ∂ z ∂r

(1)

Overall mass balance of the packed bed:

 1− εt  n ∂ qi ∂C + ρp  ∑ ∂t  εt  i=1 ∂ t

(2)

∂2 C 1 ∂  ∂ C  ∂ (uz C) ∂ (ur C) = DL 2 + Dr − r − ∂z r ∂r  ∂r  ∂z ∂r

where εt is the total void fraction (= ε + (1 − ε) × α), α is the particle porosity, and ρB is the bed density (= (1 − ε) × ρp). The energy balances of the gas and solid phases are as follows:

(

)

ε t ρ g Cˆ vg + ρ B C ps

)

n ∂T ∂C ∂ qi − ε t RT − ρ B ∑ −∆ H i ∂t ∂t ∂t i =1

(

)

) ∂ 2T ∂  ∂T  ∂T 2 hi = KL + KR − (T − Tw ) r  − ερ g C pg u 2 ∂z ∂r  ∂ r  ∂z RBi

(3)

where KL and KR are the effective axial and radial thermal conductivities, respectively. The adsorption equilibrium and adsorption rate are described by the Langmuir isotherm model and the linear driving force (LDF) model.29, 31

qi* =

q isat Bi Pi

(4)

n

1 + ∑ B j Pj j =1

4


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(

∂ qi = ω i q i* − q i ∂t

)

, ωi =

15 D ei r p2

(5)

Model for FAU-zeolite membrane A dynamic model considering the axial and radial concentration gradients was applied to the permeation mechanism.21, 23, 28 The mass balance of a component, i, in an infinitesimal volume is:

ε ∂Pi RT ∂t

+ (1 − ε )

∂qi 1 ∂(rNitot ) ∂( Nitot ) 1 ∂( Pi ) =− − −u (i = 1,L, n) ∂t r ∂r ∂z RT ∂z

where Pi is the partial pressure, qi is the adsorbed phase concentration,

ε

(6)

is the porosity, and

N itot is the overall molar flux. In Eq. (6), the total permeation flux is composed of pore diffusion and surface diffusion,

ε  N itot =   ⋅ N iP + (1 − ε ) ⋅ N iS τ 

(7)

P S where N i is the flux by pore diffusion, N i is the flux by surface diffusion, and

τ

is the

tortuosity factor. In Eq. (7), the multi-component pore diffusion consisting of Knudsen diffusion and Poilseulle diffusion can be expressed by the Dusty Gas Model (DGM).

ε 1  Kn Bi0   1 ∂rPi ∂Pi  N =− Pi  ⋅ +  Di + τ RT  ηi   r ∂r ∂z  P i

For a homogeneous porous structure, the two parameters,

(8)

DiKn

(Knudsen diffusivity) and

B0,i

(viscous diffusion parameter) in the pore diffusion model can be determined from the following equations: DiKn =

Bi0 =

2rP 8 RT ⋅ 3 πM i

(9)

rP2 8

(10)

where, rP is the pore radius, R is the ideal gas constant, T is the temperature, M is the molecular weight, and

η i is the viscosity of each gas.

In Eq. (7), the molar flux by surface diffusion can be calculated by applying the Generalized Maxwell–Stefan (GMS) model combined with the surface coverage. According to the assumptions of this model,21, 32 the surface flux of the binary mixture can be described by Eq. (11) because

5


ÐijS

(=∞)


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is much larger than

ÐiS

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.

N1S = − ρ ⋅ (1 − ε ) ⋅ q1sat ⋅

Ð1S ⋅ [ (1 − θ 2 )∇θ1 + θ1∇θ 2 ] 1 − θ1 − θ 2

(11)

In Eq. (11), the adsorptive surface coverage, θ i , is calculated from the Langmuir isotherm model. The boundary conditions of the model considering the axial and radial directions simultaneously are presented in Table 1.24, 30 According to the experimental method, the clean bed condition was used as the initial condition. Table 1. Boundary and initial conditions of AMH model Boundary

axial direction

conditions

at z = 0,

 ∂c  −DL  i  = u(ci  ∂z  z =0

z =0−

− ci

z = 0+

 ∂T  ) , u z =0 = u0 , −KL   =ερgCpg u(T z=0− −T z=0+ )  ∂z  z=0

at z = L,

 ∂ ci  =0  ∂z    z=L

 ∂u  =0    ∂z  z = L

 ∂T  =0    ∂z  z = L

radial direction at r = 0,

 ∂ ci   ∂u   ∂T  = 0,  = 0,  =0      ∂r  r =0  ∂r  r =0  ∂r  r = 0 at r = Rin ,

 ∂u   ∂ci  = 0 , hi (Tw − T =0,      ∂r  r =Rin  ∂r  z = Rin Initial condition

r = Rin

) = KR

∂T ∂r

r = Rin

ci ( z , r , t0 ) = 0 , qi ( z , r , t0 ) = 0 , T = Tatm , Twall = Tatm

The Maxwell-Stefan diffusivities and adsorption isotherm parameters for the AMH system are listed in Table 2. Using the gPROMS modeling tool (Process Systems Enterprise, Ltd.), the packed adsorbent bed model and inorganic membrane model were solved simultaneously to predict the transient (unsteady) and pseudo-steady-state performance of the system.

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Table 2. Diffusivities and adsorption isotherm parameters of CO2 and CH4 for AMH system FAU-zeolite

Ði [m2/sec]

‫ݍ‬௜௦௔௧ [mol/g]

‫ܤ‬௜ [1/Pa]

313 K

8.90×10-11

6.11×10-3

1.54×10-5

323 K

1.05×10-10

5.99×10-3

1.52×10-5

333 K

1.06×10-10

5.73×10-3

1.50×10-5

313 K

1.20×10-10

1.10×10-3

1.40×10-5

323 K

1.32×10-10

9.80×10-4

1.38×10-5

333 K

1.33×10-10

9.71×10-4

1.32×10-5

membrane CO2

CH4

‫ݍ‬௜௦௔௧ [mol/g]

‫ܤ‬௜ [1/Pa]

313 K

4.421×10-3

3.059

323 K

4.026×10

-3

2.995

3.840×10

-3

2.938

zeolite 5A CO2

333 K CH4 313 K

2.102×10-3

1.641×10-1

323 K

1.983×10-3

1.384×10-1

333 K

1.864×10-3

1.178×10-1

3. Experimental 3.1. FAU-zeolite membrane and zeolite 5A pellet A seed stock solution was prepared with 1 mg of a commercial FAU-zeolite powder (HiSiv 1000, UOP, USA) in a Teflon pot using a vibratory mill (5400, Red Devil Equipment Co., USA). A small amount of Al was also added to modify the membrane structure by slowing down the nucleation and growth but enhancing crystal intergrowth.25 Furthermore, it is known that the α-Al2O3 supported zeolite layer has much fewer inter-crystalline defects, which leads to a more effective separation process.33 In the secondary growth process, the hydrothermal temperature and time were 80 oC and 24 hr, respectively. The hydrothermal solution was 0.75Al2O3-7.5SiO2-14Na2O-840H2O in a molar basis. A porous α-Al2O3 tube (Ceracomb Co., Korea) with a length of 50 mm (35% porosity, and 120 nm mean pore diameter) was applied as a membrane support. The cleaned α-Al2O3 support was seeded on the outer surface using a vacuum assisted filtration process. In the seeding process, 4 ml of seed stock solution was diluted in 150 ml of water, the cleaned support was fully immersed in the diluted seed 7



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stock solution, and the interior of the support was evacuated at 0.026 bar (25 Torr) for 20 min. Some seeds in the diluted seed stock solution were filtered with the porous α-Al2O3 support and were uniformly coated onto the outer surface. The seeded α-Al2O3 support was dried at 383 K for more than 4 h in a convection oven and then subjected to the secondary growth process. Zeolite layers were grown over the seeded α-Al2O3 support from a hydrothermal solution comprising water glass (193-08185, Wako Pure Chemical Industries Ltd., Japan), sodium aluminate (1923-3260, Showa Chemicals Inc., Japan), sodium hydroxide (39155-0350, Junsei Chemical Co., Ltd., Japan), and water. A 330 ml sample of the hydrothermal solution was introduced into a mini-autoclave into which the seeded α-Al2O3 support was vertically installed. During the hydrothermal process, the seeds on the outer surface of the porous support grew into a zeolite layer. The prepared zeolite membrane, a porous α-Al2O3 tube coated with a dense zeolite layer, was washed three times with water, at which point the pH of the wash-water was less than 8. The washed zeolite membrane was dried at 323 K for 4 h in an oven. The membrane synthesis has been previously described in detail and the permeation of CO2 in the as-prepared membrane was similar to the previous results.34 In addition, considering the separation factor of CO2/CH4 system in the membrane, which was discussed later, defects in the membrane can be negligible. The amount of the formed FAU zeolite layer was simply calculated by using the theoretical density, thickness, and the outer diameter of the porous support. The thickness of the membrane including the support was 0.8 mm while the effective thickness of the formed zeolite layer was 0.15μ m from the SEM image of the fractured section membrane. The tortuosity of Eq. (7) for calculating the total permeation flux was obtained from the permeation flux experiments as a fitting parameter. However, it was fixed at 2 as listed in Table 3. The commercial zeolite 5A pellet (Grace & Davison Co., 4−8 mesh) was selected for the AMH system and the adsorbents were packed inside the FAU-zeolite membrane. Adsorption equilibria were measured for CH4 and CO2 on zeolite 5A by using a volumetric technique at 293.15K.24 The measured adsorption isotherms with Langmuir model are shown in Figure 1. The adsorption amount and affinity of CO2 on zeolite 5A are much stronger than those of CH4. The characteristics of the zeolite 5A pellet and FAU zeolite membrane are listed in Table 3. Table 3. Characteristics of zeolite 5A pellet and FAU-zeolite membrane Adsorbent (zeolite 5A) Type

Sphere

Pellet size [mesh]

4-8

Pellet density [g/cm3]

1.16

8


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Bulk density [g/cm3]

0.746

Heat capacity [cal/g K]

0.22

Average macropore diameter [Ǻ]

1972

Micropore diameter [Ǻ]

4.2~4.4 FAU-zeolite membrane

Average area of membrane [m2]

2.357×10-3

Effective thickness of the zeolite layer [µm]

0.15

Porosity of the zeolite layer [-]

0.20

Porosity of the support [-]

0.35

Tortuosity of the zeolite layer [-]

2

Density of the zeolite layer [g/m3]

1.0×106

Thickness of the membrane [mm]

0.8

Inner diameter of the membrane [mm]

9.1

Outer diameter of the membrane [mm]

10.7

Bed porosity [-]

0.357

Total void fraction [-]

0.77

5

Amount adsorbed [mol/kg]

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


4 CH4

CO2

CH4 293K CH4 313K CH4 333K

3

CO2 293K CO2 313K CO2 333K

2

1

0 0

500

1000

1500

2000

2500

3000

Pressure [kPa]

Figure 1. Adsorption isotherms for CH4 and CO2 on zeolite 5A (symbol for experimental data24, lines for Langmuir model)

3.2. Permeation and separation experiments A schematic of the AMH system and apparatus is presented in Figure 2. Before the 9



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experiments, the adsorbent (zeolite 5A) was regenerated for more than 12 h at 673 K. After packing the adsorbents inside the FAU-zeolite membrane, the assembly was placed in the stainless steel tubular module, i.e., a modified Wicke–Kallenbach cell consisting of a body, disks, and rods. The AMH module (a modified Wicke–Kallenbach cell) was activated at 473 K under vacuum. Prior to each experiment, the AMH module was saturated with helium gas to prevent contamination and sorption of water from the air. The binary mixture was fed into the adsorbent bed (inside), and the separated gas permeated through the membrane (outside). The retained gas was then ejected from the system, and the product was collected through the system outlet by He gas sweeping. Permeation and separation of the CO2/CH4 binary mixture (50/50 vol.%) was empirically evaluated in stage-cut mode by controlling the feed and retentate flow rates. Two mass flow controllers (MFCs, F-201C-FAC-11-V, Bronkhost, Netherands) were used to control the retentate flow rate to fix the stage-cut. Pre-testing by varying the retentate flow rate according to the permeated flow rate was carried out to fix the two MFCs for the sweep gas (He) and retentate to set the stage-cut for each experiment; the experiments were started after system initialization. The permeation flux and composition were measured by using a mass flow meter and an online mass spectrometer (HPR 20, Hiden Analytical Ltd., UK). The permeation and separation experiments employing the AMH system were performed at temperatures between 313 and 333 K and pressures of 100 to 600 kPa. The FAU-zeolite membrane only was evaluated under the same conditions for comparison. In this study, the transient (unsteady) permeation and pseudo-steady-state separation performances were evaluated. The term “pseudosteady-state” implies constant permeation flux in the AMH system, while “transient (or unsteadystate)” indicates the intermediate state before reaching the pseudo-steady-state. (10)

(9)

6

5

(6) 1

Sweeping gas (7)

2

(1)

3

(5)

(2)

50.00

(1) (2) (3) (4) (5)

(3)

Permeated gas (11)

50.00

(8)

Retentate gas

(4)

Feed gas Feed pressure transducer RTD & Temperature controller Membrane cell MFC(Sweeping gas control)

(6) Sweeping gas (He) (7) MFM(Permeat flow) (8) Quadruple Mass Spectrometer (9), (10) Data interfacing computermeter (11) MFC(Stage cut control)

4

1 2 3 4 5 6

: : : : : :

Adsorbent FAU-zeolite membrane AMH module Mixed gas(Feed gas) in Mixed gas out Permeate gas out

Figure 2. Schematic diagram of gas permeation and separation measurement system (modified Wicke-Kallenbach AMH cell) 10


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4. Results and Discussion 4.1. Permeation and separation in the FAU-zeolite membrane Prior to the experiments employing the AMH system, the permeation flux and separation factor of the CO2/CH4 mixture (50:50 vol.%) in the FAU-zeolite membrane-only system were measured as functions of the pressure and temperature under fixed feed flow rate, stage-cut, and sweep gas flow rate conditions. Figures 3 and 4 respectively show the experimental transient permeation and pseudo-steady-state variation along with the simulated results. For the membrane alone (Fig. 3), since a greater amount of strong adsorptive CO2 penetrated the FAU zeolite membrane via surface diffusion, permeation of the relatively weak adsorbate, CH4, was hindered by competitive adsorption. Therefore, the permeation flux of CH4 was much smaller than that of CO2, but the time required to reach pseudo-steady-state permeation flux for CH4 was slightly faster than that of CO2. As shown in Fig. 3, the permeation flux reached pseudo-steady-state more quickly with an increase in the temperature due to activation of molecular motion. At 313 K and 333 K, the flux reached pseudo-steady-state at 150 and 100 s, respectively. The transient permeation curves for CO2 and CH4 followed similar trends to those documented.35 Increasing the temperature from 313 K to 323 K caused the relative ratio of the CH4 permeation flux to increase to a greater extent than that of CO2 because the 'hindrance effect' of adsorbed CO2 was reduced. Therefore, it is expected that the hindrance effect derived from surface diffusion of the adsorbed molecules was dominant to pore diffusion and adsorption equilibrium when separation was performed at lower temperatures. Further, the variation of the flux in the temperature range of 323 K to 333 K was relatively small. The changes in the permeability and selectivity were also small given that the mechanisms of separation and permeation were primarily controlled by the adsorption equilibrium difference. The kinetic contribution from the tetrahedral structure and the kinetic diameter of CH4 were relatively small in the FAU-zeolite membrane system, which is different from the case of the kinetic control membrane.22, 23

11



0.012

Permeation Flux [mol/m2sec]

CO2 CH4

0.010

Simulation 0.008

0.006

0.004

0.002

(a)

0.000 0

25

50

75

100

125

150

175

200

Time [sec]

0.012


CO2 CH4

0.010

Simulation 0.008

0.006

0.004

0.002

(b)

0.000 0

25

50

75

100

125

150

175

200

Time [sec]

0.012

CO2 Permeation Flux [mol/m2sec]

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

CH4

0.010

Simulation 0.008

0.006

0.004

0.002

(c)

0.000 0

25

50

75

100

125

150

175

200

Time [sec]

Figure 3. Transient permeation flux of CO2/CH4 mixture (50/50 vol%) on FAU zeolite membrane at 150 kPa, sweeping gas flow rate (=50 sccm) and stage cut (=0.60): (a) 313 K, (b) 323 K, and (c) 333 K.

12


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The permeation flux curves (Fig. 4) show favorable curvature with pressure at a fixed temperature because the amount of CO2 and CH4 adsorbed on the FAU zeolite increased with pressure. At all temperatures, a steep increase of the permeation flux was observed up to 200 kPa. The permeation flux subsequently increased asymptotically and the separation factor decreased slightly with pressure. At pressures exceeding 200 kPa, the difference in the permeation flux and separation factor at various temperatures remained approximately the same with variation of the pressure. In addition, as the temperature increased from 313 to 323 K, the permeation flux increased significantly and the separation factor decreased significantly, as expected from Figure 3. However, the difference in the permeation flux and separation factor at temperatures in the range of 323 and 333 K was very small. The highest permeability was achieved at 333 K, with maximum permeabilities of CO2 and CH4 of 9.48 × 10−8 mol·(m2·sec·Pa)−1 and 3.04 × 10−8 mol·(m2·sec·Pa)−1, respectively. The maximum separation factor of the CO2/CH4 mixture was 4.48 at 313 K. The amount of CO2 and CH4 adsorbed on the FAU-zeolite membrane increased with pressure, with a consequent increase in the permeation flux; however, the relative increase of the CH4 flux led to a decrease of the separation factor. On the other hand, the adsorption amount and affinity decreased at higher temperature because adsorption is exothermic. In addition, since the contribution of surface diffusion to CO2 permeation was much higher than that to CH4 permeation, the surface diffusion effect decreased and the molecules were activated with increasing temperature. Therefore, a decline of the separation factor was observed in Figure 4.

0.025

CO2 Permeation Flux [mol/m2sec]

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


313 K 323 K 333 K Simulation

0.020

0.015

0.010

0.005

0.000

(a) 100

150

200

250

300

350

Pressure [kPa]

13


400


5.0

Simulation CO2/CH4 Separation Factor [ - ]

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

4.5

313 K 323 K 333 K

4.0

3.5

3.0

2.5

2.0

(b) 1.5 100

150

200

250

300

350

400

Pressure [kPa]

Figure 4. Pseudo-steady-state permeation of CO2/CH4 mixture (50/50 vol.%) on FAU-zeolite membrane at 313 – 333 K, sweeping gas flow rate (=50sccm) and stage cut (=0.60): (a) CO2 permeation flux and (b) separation factor.

4.2. Permeation and separation in the AMH system In the AMH system, the mixed gas was fed into the adsorbent bed (inside the membrane) for bulk separation; the gases were then purified by the membrane. Figure 5 shows the effect of temperature on the transient CO2/CH4 mixture at 150 kPa in the AMH system. The time required to reach pseudo-steady-state and the slope of the transient permeation profile were slightly slower and wider, respectively, than those of the FAU-zeolite membrane (Fig. 4) because the pseudo-steady-state must be achieved in both the packed zeolite 5A pellets and the FAU-zeolite membrane. Compared to the FAU-zeolite membrane system, at all temperatures, the CH4 permeation flux declined more significantly than the CO2 permeation flux in the AMH system. The CO2 molecules with strong adsorption affinity are selectively captured by the zeolite 5A pellets and the FAU zeolite membrane. According to the adsorption breakthrough phenomenon in a packed bed,29, 30 the mass transfer zones of both components are separated and the concentration wavefront of CH4 propagates axially faster than that of CO2 in the zeolite packed bed. However, in the AMH system, the feed flowed in the axial and radial directions. Due to the permeation of CO2 through the FAU-zeolite membrane and the improved interfacial contact of the gas phase with zeolite 5A and the membrane, the CO2 permeation flux in the membrane declined to a relatively lower extent than the CH4 permeation flux. The strong adsorption affinity of CO2 and the increased CO2 concentration in the membrane pore led to a greater difference in the permeation flux of CO2 versus CH4. Furthermore, the permeation flux increased with increasing temperature in the temperature range of 313 to 333 K because of the variation of the adsorption amount with temperature. However, the faster propagation 14


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of CH4 in the packed zeolite 5A led to a decrease of the slope of the CO2 permeation transient curve and extension of the range corresponding to the small difference between the two permeation fluxes during the initial period, compared to the results for the FAU-zeolite membrane in Figure 3.

0.0035

CH4

0.0030

Simulation

2

Permeation Flux [mol/m sec]

CO2

0.0025 0.0020 0.0015 0.0010 0.0005

(a)

0.0000 0

25

50

75

100

125

150

175

200

Time [sec]

0.0035


CO2 CH4

0.0030

Simulation 0.0025 0.0020 0.0015 0.0010 0.0005

(b)

0.0000 0

25

50

75

100

125

150

175

200

Time [sec]

0.0035

CO2 CH4

0.0030

Simulation

2


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


0.0025 0.0020 0.0015 0.0010 0.0005

(c)

0.0000 0

25

50

75

100

125

150

Time [sec]

15


175

200


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

Figure 5. Transient permeation flux of CO2/CH4 binary mixture (50/50 vol%) on zeolite 5A/FAU-zeolite membrane hybrid system at 150 kPa, sweeping gas flow rate (=50 sccm) and stage cut (=0.60): (a) 313 K, (b) 323 K, and (c) 333 K.

Figure 6 shows the variation of the permeation flux and the separation factor of the CO2/CH4 mixture in the AMH system at 313, 323, and 333 K. As the feed pressure increased, the permeation flux increased, but the separation factor decreased, which is similar to the results obtained with the FAU-zeolite membrane (Fig. 4). At 333 K, the maximum permeability of CO2 and CH4 was 2.88 × 10−8 and 5.04 × 10−9 mol·(m2·sec·Pa)−1, respectively, and the maximum separation factor was 5.72. The permeation flux increased asymptotically with pressure for the integrated system, which was the same as the results for the FAU-zeolite membrane only (Fig. 4(a)); however, the variation of the interval pattern with temperature was similar for both systems due to the strong dependence of the adsorption capacities of zeolite 5A on temperature and pressure. Compared to the FAU-zeolite membrane, the permeation flux was small for the integrated system under equivalent conditions because of gas adsorption in the packed adsorbents. Figure 6(b) illustrates that the separation factor in the AMH system was greater than that in the FAU-zeolite membrane, which mainly stemmed from the more significant decrease in the CH4 flux, as shown in Figure 5. Notably, for the integrated system, the slope of the plot of the variation of the separation factor with pressure was steeper than that for the FAU-zeolite membrane (see Fig. 4). The adsorption isotherm of zeolite 5A showed a similar variation of the adsorption amount with increasing temperature and decreasing pressure. However, the separation factor showed a weaker temperature-dependence for the AMH system than in the case of the FAU-zeolite membrane. Furthermore, at 333 K, the separation factor for the AMH system was greater than that at the other temperatures, which is opposite to the results obtained with the FAU-zeolite membrane. As the temperature increased, less CO2 was adsorbed on the packed zeolite 5A as well as on the FAU-zeolite membrane. On the other hand, since the zeolite 5A pellets could still strongly adsorb CO2, the zeolite 5A pellets contributed to improving the separation in the FAU-zeolite membrane. As a result, the separation factor in the AMH system was greater than that in the FAU-zeolite membrane-only system under the same operating conditions. The decrease of separation factor with increasing pressure in the AMH system (Fig. 6(b)) was greater than that in the FAU-zeolite membrane (Fig. 4). And the separation factor at the highest pressure in the study (600 kPa) became almost the same as the FAU-zeolite membrane-only system, implying the loss of all the advantages of the AMH system. Since the adsorption isotherm for CO2 on zeolite 5A has a highly favorable shape (Type I), CO2 adsorption approaches saturation with increasing pressure. Therefore, with increasing pressure, more CO2 was adsorbed on the zeolite; CH4enrichment may become too much in a gas phase, thereby reducing the advantage of the contact 16


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efficiency of CO2 with the membrane and zeolite in the radial direction. This indicates that an operating pressure near the saturation of adsorbents is not desirable in the AMH system. 313 K 323 K 333 K Simulation

2

CO2 Permeation Flux [mol/m sec]

0.010

0.008

0.006

0.004

0.002

(a)

0.000

100

200

300

400

500

600

700

Pressure [kPa] 6.0

CO2/CH4 Separation Factor [ - ]

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


313 K 323 K 333 K Simulation

5.5

5.0

4.5

4.0

3.5

(b) 3.0 100

200

300

400

500

600

700

Pressure [kPa]

Figure 6. Pseudo-steady-state permeation of CO2/CH4 mixture (50/50 vol%) on zeolite 5A/FAU zeolite membrane hybrid system at 313.15 – 333.15 K, sweeping gas flow rate (=50sccm) and stage cut (=0.60): (a) CO2 permeation flux and (b) separation factor

4.3 Effects of sweeping gas and retentate flow rates on the AMH system The flow of the sweeping gas (helium) contributed to improving the separation of the CO2/CH4 mixture because low CO2 coverage on the permeate side enhances the adsorption selectivity of the membrane. As shown in Figure 7, with an increase in the flow rate of the sweeping gas, the permeation flux in the AMH system increased steeply then decreased. In contrast, the separation factor increased gradually with the sweeping gas flow. If viscous flow or Knudsen flow is operative in the pores, back diffusion of the sweeping 17



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

gas will cause the separation factor to increase by pushing the weak adsorbate from the adsorbed pores to the retentate side. However, as shown in Figure 7, with an increase in the sweeping gas flow rate, the separation factor increased only slightly compared to the decrease in the permeation flux. Thus, surface diffusion was dominant in the separation of the CO2/CH4 mixture in the FAU zeolite membrane system due to the strong affinity for CO2. The contribution of the sweeping gas to the permeation driving force derived from decreasing the CO2 concentration on the permeate side was relatively small.36 The effect of the retentate flow rate on the gas separation was studied under fixed pressure, temperature, and sweeping gas flow rate conditions. Figure 8 shows the transient permeation behavior in the AMH system with variation of the retentate flow rate at 150 kPa and 323 K. The experimental data at a retentate flow rate of 200 sccm was compared to the simulated results at 50, 100, 150, 200, and 250 sccm. Interestingly, the shape and the pattern of variation of the transient permeation curves with the retentate flow rate were very similar to that of the breakthrough curves for a typical adsorption bed.36 The time required to reach pseudo-steady-state permeation flux decreased with an increase in the retentate flow rate. Increasing the retentate flow rate from 50 to 100 sccm induced a significant change in the transient permeation. The difference between the two retentate flow rates became smaller with an increase in the retentate flow rate. Moreover, the time required to reach pseudo-steady-state permeation flux of CO2 under lower retentate flow conditions was greater than that for CH4 because CH4 propagation in the packed zeolite pellets was faster than CO2 propagation. Therefore, a reduction of the permeation flux and separation factor was expected at lower retentate flow rates. Figure 9 shows the effects of the retentate flow rate on the permeation flux and separation factor. At retentate flow rates from 200 to 350 sccm, the variation of the permeation flux was nearly constant, and the separation factor was similarly maintained (5.2−5.3). The difference in the transient permeation between 200 and 250 sccm was very small, as shown in Fig. 8. The AMH system can be operated at a retentate flow rate of 200 sccm without a significant decrease in the separation factor or permeation flux. However, below 150 sccm, the permeation flux of CO2 varied somewhat with the retentate flow rate, whereas the variation of the permeation flux of CH4 was trivial. Therefore, the separation factor declined remarkably despite the small decrease in the permeation flux of CO2.

18


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6 0.004

5

0.003

0.002

4

0.001

Permeation Flux CO2/CH4 Separation Factor


2

CO2 Permeation Flux [mol/m sec]

0.005

3

0.000 0

50

100

150

200

Sweeping Gas Flow Rate [ccm]

Figure 7. Effect of sweeping gas flow rate on zeolite 5A/FAU-zeolite membrane hybrid system at 150 kPa, 323K and constant stage cut (=0.60).

0.0018

CO2 (200 ccm) CH4 (200 ccm)

0.0016 2


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


0.0014 0.0012

50 ccm 100 ccm 150 ccm 200 ccm 250 ccm

0.0010 0.0008 0.0006 0.0004 0.0002 0.0000 0

50

100

150

200

250

300

Time [sec]

Figure 8. Effect of retentate flow rate in transient permeation on zeolite 5A/FAU-zeolite membrane hybrid system at 150 kPa, 323K and constant sweeping gas flow rate (=200 sccm).

19



0.0030

0.0028

2

5.4

5.2 0.0026 5.0 0.0024 4.8

CO2/CH4 Separation Factor Simulation Permeation Flux SImulation

0.0022

4.6


5.6


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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4.4

0.0020 0

50

100

150

200

250

300

350

Retentate Flow Rate [ccm]

Figure 9. Effect of retentate flow rate in pseudo-steady-state permeation on zeolite 5A/FAUzeolite membrane hybrid system at 150 kPa, 323 K and constant sweeping gas flow rate (=200 sccm).

5. Conclusions The separation of a CO2/CH4 mixture (50/50 vol%) by the adsorbent/membrane hybrid (AMH) system, which consisted of a FAU zeolite membrane packed with zeolite 5A pellets, was compared to that of the FAU zeolite membrane alone based on experimental and theoretical analyses. The permeation and separation in the AMH system were well-predicted by the integrated model comprising the adsorbent-packed bed model and tubular-type membrane model. In the packed zeolite 5A pellets, the axial propagation of CH4 was faster than that of CO2. However, in the AMH system, the feed also flowed radially through the membrane and strongly adsorbed CO2 on the membrane surface hindered the permeation of CH4. In addition, the packed zeolite 5A pellets contributed to enhancing the interfacial contact between the membrane and gas phase in the radial direction. Therefore, the reduction of the permeation flux of the less strongly adsorbed molecule (CH4) was more significant than that of the more strongly adsorbed molecule (CO2). The separation factor of the CO2/CH4 mixture was greater than that of the FAU-zeolite membrane alone. However, compared to the FAU zeolite membrane, the permeability of the AMH system was reduced due to gas adsorption in the packed adsorbents. Due to the gas adsorption of the zeolite 5A pellets, the dependence of the permeation flux and separator factor on temperature and pressure for the integrated system was weaker than that in the FAU-zeolite membrane-only system. The maximum permeability of the FAU-zeolite membrane alone 20


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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


was 1.25 × 10−7 mol·m−2·sec−1·pa−1 (CO2: 9.48 × 10−8 mol·m−2·sec−1·pa-1 and CH4: 3.04 × 10−8 mol·m−2·sec−1·pa−1) at 333 K, and the maximum separation factor was 4.48 at 313 K. For the AMH system, the maximum permeability was 3.38 × 10−8 mol·m-2·sec-1·pa-1 (CO2: 2.88 × 10−8 mol·m−2sec−1 Pa and CH4: 5.04 × 10−9 mol·m−2sec−1) and the maximum separation factor was 5.72. For the AMH system, the gas permeation and separation were also significantly affected by the operating variables such as the sweeping gas flow rate and the stage-cut. The permeation flux was reduced by increasing the flow rate of the sweeping gas after an initial steep increase in the low sweep gas range. However, the separation factor subsequently showed a sharp increase with an increase of the sweep gas flow rate, followed by a smooth increase. Similarly, as the retentate flow rate increased, the selectivity for more strongly adsorbed molecules was drastically enhanced. The AMH system integrating the two equilibrium separation agents (the FAU-zeolite membrane and the zeolite 5A pellets) exhibited enhanced separation efficiency. In addition, the separation performance of this system was less affected by fluctuation of the feed temperature and pressure than that of the membrane alone. However, since the packed zeolite pellets may become saturated after certain period, a regeneration step such as purging or depressurization, which are general techniques used in the pressure swing adsorption process, may be required. For a continuous production process, the multi-module system should be operated in a cyclic sequence, similar to that employed in the PSA and TSA processes.

Acknowledgement We would like to acknowledge the financial support from the Ministry of Environment and Korea Research Institute of Chemical Technology (KRICT : E616-00169-0605-1).

Nomenclature

Roman letters

bi

parameter in the Langmuir isotherm model (Pa-1)

b0,i

parameter of species i in the Langmuir isotherm model at zero loading (Pa-1)

Cp

heat capacity (J g−1 K−1)

Ð0,i

Maxwell-Stefan surface diffusivity at infinite temperature (m2 sec-1)

Ði

Maxwell-Stefan surface diffusivity (m2 sec-1) 21



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

∆H ads

heat of adsorption (J mol-1)

k1

temperature dependent parameter for saturated adsorption capacity (mol/g)

k2

temperature dependent parameter for saturated adsorption capacity (-)

KL

effective axial thermal conductivity (J cm-1s-1K-1)

KR

effective radial thermal conductivity (J cm-1s-1K-1)

Ni

molar flux of species i (mol m-2 sec-1)

N iP

molar flux of pore diffusion of species i (mol m-2 sec-1)

N iS

molar flux of surface diffusion of species i (mol m-2 sec-1)

Nitot

total molar flux of species of i (mol m-2 sec-1)

Pi

partial pressure of species i, (Pa)

qi

adsorbed species concentration within layer pores, (mol g-1)

qi

sat

saturated capacity of adsorbed species i (mol g-1)

R

gas constant, 8.314 (J mol-1 K-1)

T

absolute temperature, (K)

t

time, (sec)

Greek letters

ε

porosity of membrane (-)

εt

total void fraction

ρ

membrane density (g m-3)

τ

tortuosity factor (-)

Subscripts/superscripts i,j,1,2

component i,j,1,2

tot

total flow

sat

saturated

B

bed

P

adsorbent pellet

g

gas phase

s

solid phase

w

wall

22


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Separation

characteristics

of

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tetrapropylammoniumbromide templating silica/alumina composite membrane in CO2/N2, CO2/H2 and CH4/H2 systems. Korean J. Chem. Eng. 2004, 21, (2), 477-487. (33)

Kosinov, N.; Gascon, J.; Kapteijn, F.; Hensen, E. J. Recent developments in zeolite

membranes for gas separation. J. Membr. Sci. 2016, 499, 65-79. (34)

Cho, C. H.; Yeo, J. G.; Ahn, Y. S.; Han, M. H.; Kim, Y. H.; Hyun, S. H. Secondary Growth of

Sodium Type Faujasite Zeolite Layers on a Porous α-Al₂O₃ Tube and the CO₂/N₂ Separation.

Membr. J. 2007, 17, (3), 254-268. (35)

Zhang, J.; Burke, N.; Zhang, S.; Liu, K.; Pervukhina, M. Thermodynamic analysis of

molecular simulations of CO2 and CH4 adsorption in FAU zeolites. Chem. Eng. Sci. 2014, 113, 5461. (36)

Skoulidas, A. I.; Sholl, D. S. Multiscale models of sweep gas and porous support effects on

zeolite membranes. AICHE J. 2005, 51, (3), 867-877.

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Figure List Figure 1.

Adsorption isotherms for CH4 and CO2 on zeolite 5A(symbol for experimental data24, lines for Langmuir model)

Figure 2.

Schematic diagram of gas permeation and separation measurement system (modified Wicke-Kallenbach AMH cell)

Figure 3.

Transient permeation flux of CO2/CH4 mixture (50/50 vol%) on FAU zeolite membrane at 150 kPa, sweeping gas flow rate (=50 sccm) and stage cut (=0.60): (a) 313 K, (b) 323 K, and (c) 333 K.

Figure 4.

Pseudo-steady-state permeation of CO2/CH4 mixture (50/50 vol.%) on FAU-zeolite membrane at 313 – 333 K, sweeping gas flow rate (=50sccm) and stage cut (=0.60): (a) CO2 permeation flux and (b) separation factor.

Figure 5.

Transient permeation flux of CO2/CH4 binary mixture (50/50 vol%) on zeolite 5A/FAU-zeolite membrane hybrid system at 150 kPa, sweeping gas flow rate (=50 sccm) and stage cut (=0.60): (a) 313 K, (b) 323 K, and (c) 333 K.

Figure 6.

Pseudo-steady-state permeation of CO2/CH4 mixture (50/50 vol%) on zeolite 5A/FAU zeolite membrane hybrid system at 313.15 – 333.15 K, sweeping gas flow rate (=50sccm) and stage cut (=0.60): (a) CO2 permeation flux and (b) separation factor

Figure 7

Effect of sweeping gas flow rate on zeolite 5A/FAU-zeolite membrane hybrid system at 150 kPa, 323K and constant stage cut (=0.60).

Figure 8.

Effect of retentate flow rate in transient permeation on zeolite 5A/FAU-zeolite membrane hybrid system at 150 kPa, 323K and constant sweeping gas flow rate (=200 sccm).

Figure 9

Effect of retentate flow rate in pseudo-steady-state permeation on zeolite 5A/FAUzeolite membrane hybrid system at 150 kPa, 323 K and constant sweeping gas flow rate (=200 sccm).

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Table Table 1.

Boundary and initial conditions of AMH model

Table 2.

Diffusivities and adsorption isotherm parameters of CO2 and CH4 for AMH system

Table 3.

Characteristics of zeolite 5A pellet and FAU-zeolite membrane

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For Table of Contents Only

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For Table of Contents Only 84x47mm (300 x 300 DPI)



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Figure 1. Adsorption isotherms for CH4 and CO2 on zeolite 5A (symbol for experimental data24, lines for Langmuir model) Figure 1 222x190mm (300 x 300 DPI)


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Figure 2. Schematic diagram of gas permeation and separation measurement system (modified WickeKallenbach AMH cell) Figure 2 338x172mm (300 x 300 DPI)



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Figure 3. Transient permeation flux of CO2/CH4 mixture (50/50 vol%) on FAU zeolite membrane at 150 kPa, sweeping gas flow rate (=50 sccm) and stage cut (=0.60): (a) 313 K, (b) 323 K, and (c) 333 K. Figure 3 259x190mm (300 x 300 DPI)


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Figure 4. Pseudo-steady-state permeation of CO2/CH4 mixture (50/50 vol.%) on FAU-zeolite membrane at 313 – 333 K, sweeping gas flow rate (=50sccm) and stage cut (=0.60): (a) CO2 permeation flux and (b) separation factor. Figure 4 262x190mm (300 x 300 DPI)



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Figure 4. Pseudo-steady-state permeation of CO2/CH4 mixture (50/50 vol.%) on FAU-zeolite membrane at 313 – 333 K, sweeping gas flow rate (=50sccm) and stage cut (=0.60): (a) CO2 permeation flux and (b) separation factor. Figure 4 255x190mm (300 x 300 DPI)


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Figure 5. Transient permeation flux of CO2/CH4 binary mixture (50/50 vol%) on zeolite 5A/FAU-zeolite membrane hybrid system at 150 kPa, sweeping gas flow rate (=50 sccm) and stage cut (=0.60): (a) 313 K, (b) 323 K, and (c) 333 K. Figure 5 266x190mm (300 x 300 DPI)



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Figure 6. Pseudo-steady-state permeation of CO2/CH4 mixture (50/50 vol%) on zeolite 5A/FAU zeolite membrane hybrid system at 313.15 – 333.15 K, sweeping gas flow rate (=50sccm) and stage cut (=0.60): (a) CO2 permeation flux and (b) separation factor Figure 6 263x190mm (300 x 300 DPI)


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Figure 6. Pseudo-steady-state permeation of CO2/CH4 mixture (50/50 vol%) on zeolite 5A/FAU zeolite membrane hybrid system at 313.15 – 333.15 K, sweeping gas flow rate (=50sccm) and stage cut (=0.60): (a) CO2 permeation flux and (b) separation factor Figure 6 255x190mm (300 x 300 DPI)



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Figure 7. Effect of sweeping gas flow rate on zeolite 5A/FAU-zeolite membrane hybrid system at 150 kPa, 323K and constant stage cut (=0.60). Figure 7 280x190mm (300 x 300 DPI)


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Figure 8. Effect of retentate flow rate in transient permeation on zeolite 5A/FAU-zeolite membrane hybrid system at 150 kPa, 323K and constant sweeping gas flow rate (=200 sccm). Figure 8 259x190mm (300 x 300 DPI)



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Figure 9. Effect of retentate flow rate in pseudo-steady-state permeation on zeolite 5A/FAU-zeolite membrane hybrid system at 150 kPa, 323 K and constant sweeping gas flow rate (=200 sccm). Figure 9 286x190mm (300 x 300 DPI)


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Table 1. Boundary and initial conditions of AMH model Boundary

axial direction

conditions at z = 0,

 ∂c  − DL  i  = u(ci  ∂z  z =0

z = 0−

− ci

z = 0+

 ∂T  ) , u z =0 = u0 , −KL   =ερgCpg u(T z=0− −T z=0+ )  ∂z  z=0

at z = L,

 ∂ ci  =0    ∂z  z = L

 ∂u  =0    ∂z  z = L

 ∂T  =0    ∂z  z = L

radial direction at r = 0,

 ∂ ci   ∂u   ∂T  = 0,  =0,  =0    ∂r   ∂r  r =0  ∂r  r = 0   r =0 at r = Rin ,

 ∂u   ∂ci  = 0 , hi (Tw − T = 0 ,    ∂r  ∂ r   r =Rin   z =Rin Initial condition

ci ( z , r , t0 ) = 0 , qi ( z , r , t0 ) = 0 ,

r = Rin

) = KR

∂T ∂r

T = Tatm , Twall = Tatm


r = Rin


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Table 2. Diffusivities and adsorption isotherm parameters of CO2 and CH4 for AMH system FAU-zeolite membrane

Ði [m2/sec]

[mol/g]

[1/Pa]

CO2 313 K

8.90×10-11

6.11×10-3

1.54×10-5

323 K

1.05×10-10

5.99×10-3

1.52×10-5

333 K

1.06×10-10

5.73×10-3

1.50×10-5

313 K

1.20×10-10

1.10×10-3

1.40×10-5

323 K

1.32×10-10

9.80×10-4

1.38×10-5

333 K

1.33×10-10

9.71×10-4

1.32×10-5

CH4

zeolite 5A

[mol/g]

[1/Pa]

CO2 313 K

4.421×10-3

3.059

323 K

4.026×10-3

2.995

333 K

3.840×10-3

2.938

313 K

2.102×10-3

1.641×10-1

323 K

1.983×10-3

1.384×10-1

333 K

1.864×10-3

1.178×10-1

CH4


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Table 3. Characteristics of zeolite 5A pellet and FAU-zeolite membrane Adsorbent (zeolite 5A) Type

Sphere

Pellet size [mesh]

4-8

Pellet density [g/cm3]

1.16

Bulk density [g/cm3]

0.746

Heat capacity [cal/g K]

0.22

Average macropore diameter [Ǻ]

1972

Micropore diameter [Ǻ]

4.2~4.4 FAU-zeolite membrane

Average area of membrane [m2]

2.357×10-3

Effective thickness of the zeolite layer [μm]

0.15

Porosity of the zeolite layer [-]

0.20

Porosity of the support [-]

0.35

Tortuosity of the zeolite layer [-]

2

Density of the zeolite layer [g/m3]

1.0×106

Thickness of the membrane [mm]

0.8

Inner diameter of the membrane [mm]

9.1

Outer diameter of the membrane [mm]

10.7

Bed porosity [-]

0.357

Total void fraction [-]

0.77


Separation of Carbon Dioxide and Methane Mixture by an Adsorbent

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