Also, they modified the above mass transfer model for chemical absorption of CO2 using monoethanolamine (MEA) and NaOH liquid solutions.(16) Shirazian...
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Aug 21, 2014 - ... and Systems (Chongqing University), Ministry of Education of PRC, ... Department of Chemical and Biomolecular Engineering, The Ohio ...
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Theoretical study on CO2 absorption from biogas by membrane contactors: Effect of operating parameters Zhien Zhang, Yunfei YAN, Li Zhang, Yuanxin Chen, Jingyu Ran, Ge Pu, and Changlei Qin Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/ie502830k • Publication Date (Web): 21 Aug 2014 Downloaded from http://pubs.acs.org on August 26, 2014
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Theoretical study on CO2 absorption from biogas by membrane contactors: Effect of operating parameters
Zhien Zhang†, Yunfei Yan †,*, Li Zhang†, Yuanxin Chen‡, Jingyu Ran†, Ge Pu† and Changlei Qin† †
Key Laboratory of Low-grade Energy Utilization Technologies and Systems (Chongqing
University), Ministry of Education of PRC, Chongqing 400044, China ‡
William G. Lowrie Department of Chemical and Biomolecular Engineering, Ohio State University,
Columbus, OH 43210, USA
Supporting Information ABSTRACT: Biogas upgrading and utilization is a novel technology to obtain resource-efficient vehicle fuel. In this study, a mass transfer model for CO2 absorption from biogas into potassium argininate (PA) solutions was developed. The computational fluid dynamics (CFD) methods were employed to solve the differential equations in three domains of the membrane contactor. The simulations were focused on the characteristics of both gas and absorbent phases to demonstrate the concentration distributions in axial and radial directions in the module. The simulated results were in excellent agreement with experimental data when considering the effect of initial CO2 concentration and gas velocity. Furthermore, the effect of operating pressure, flow pattern, flow condition, and modules in series on the membrane performance was investigated. The results showed the purity of CH4 reached 95 % with the operating pressure of 0.9 MPa. It was found that a fluid in the turbulent condition or counter-current configuration had a significant effect on improving the contactor performance. The simulation results also indicated that the use of two modules could increase CO2
removal and obtain high CH4 purity. Finally, the results confirmed that the developed 2D model was able to predict the behavior of CO2 separation in the membrane contactors.
1. Introduction With the growing crisis in energy, biogas is regarded as a kind of green and renewable energy which could replace the traditional fossil fuels. However, the initial biogas contains a large amount of CO2. In order to reuse biogas, the purity of CH4 used in vehicle fuel should be above 95 %. Thus, it is crucial to improve the utilization efficiency of biogas and make it to be a substitute of natural gas. Currently, there are a variety of methods for CO2 capture from biogas including physical and chemical absorption, pressure and temperature swing adsorption, cryogenic refrigeration, and membrane technology.1-3 Recently, the membrane absorption method which combines the advantages of chemical absorption and membrane technique has drawn researchers’ attention. This novel technology provides no channeling, flooding, or foaming issues; small size, lightweight, and it is easy to operate in comparison with conventional separation methods.4 Thus, a large number of scholars have investigated acid gas removal from gas mixture using a membrane module experimentally and theoretically. Some studies have been investigated the effects of fluid and membrane module properties on the contactor performance.5-9 Ghadiri et al.10 proposed a 2D mathematical model for CO2 stripping in nano porous membrane contactors. It revealed that the system with a higher liquid velocity or slower gas velocity provided better CO2 stripping performance. The similar phenomenon was also reported by
porosity-to-tortuosity ratio on CO2 removal in the non-wetting mode. Increasing the liquid concentration leaded to increasing the percentage of CO2 capture. Meanwhile, the membrane with a higher porosity-to-tortuosity ratio improved the absorption process due to the decrease of membrane mass transfer resistance. Wang et al.12 investigated the influences of membrane length, membrane diameter, and wetting ratio on CO2 absorption performance. They recommended that a shorter membrane module was preferred for thinner membranes to avoid the ineffective part of the membrane contactor. Meanwhile, it was suggested that the operation without wetting was good for the CO2 stripping process. So far there are some work focusing on the experiments of CO2 membrane absorption from biogas considering the effect of operating conditions. But the biogas upgrading in a hollow fiber membrane contactor (HFMC) using CFD approaches was rarely reported. Previous modeling investigations are mainly concentrated on acid gas separation from flue gas or natural gas.13,14 Al-Marzouqi et al.15 developed a comprehensive 2D mathematical mass transfer model for CO2 physical absorption from CO2/CH4 gas mixture using a HFMC. Also, they modified the above mass transfer model for chemical absorption of CO2 using monoethanolamine (MEA) and NaOH liquid solutions.16 Shirazian et al.17 carried out the numerical simulation CO2 removal from CO2/N2 gas mixture in membrane contactors. Faiz et al.18 used a mathematical model for CO2 capture from natural gas at high pressure up to 5 MPa. These works revealed that the proposed models could predict well for CO2 separation by hollow fiber membrane modules. Some researchers also concentrated on the mathematical modeling of H2S or SO2 absorption using gas-liquid hollow fiber membrane contactors.19-21 There is no study regarding CO2 separation from biogas using various absorbent solutions in membrane contactors.
In addition, absorbent is a very significant factor for the gas-liquid membrane system. A variety of physical and chemical solutions, and mixed aqueous solvents have been performed in membrane gas absorption fields. In the case of physical absorption, water and methanol are the common absorbents which are widely used in research.22,23 But using the physical method could not get high CO2 removal efficiency. Although the chemical solvents provide high CO2 absorption efficiency, mature technology, and easy operation, the regeneration performance of the absorbent is required to be improved. In order to make use of an absorption-desorption cycle to capture CO2, Abanades et al.24 compared the absorbent cost and performance in a range of systems. In our previous publication, the absorption and regeneration performance of various absorbent solutions were systematically concluded.4 Regarding single absorbent, the order of CO2 absorption performance was MEA > diethanolamine (DEA) > 2-Amino-2-methyl-1-propanol (AMP) > triethanolamine (TEA) > N-methyldiethanolamine (MDEA), while the generation performance order was TEA > MDEA > DEA > AMP > MEA. Recently, hot potassium carbonated (K2CO3) solution has been proved to be a promising alternative of conventional amine solutions for CO2 separation.25,26 It was reported that the performance of different absorbent solutions for CO2 absorption was blend of DEA and MDEA < MDEA < DEA < AMP < K2CO3. Currently, there is a few studies reporting the amino salts used as the liquid solutions in the gas-liquid membrane absorbers.27 In the present study, a comprehensive mass transfer model is proposed for CO2 capture from biogas. PA solutions are used as the absorbent for the non-wetted conditions. Because arginine with an amino group in carbonate solution shows excellent characteristics like low volatility, ionic property, and resistance to oxidative and thermal degradation. The effects of gas velocity and composition, operating pressure, flow pattern, flow condition, and module connection form are
systematically observed. The aim of this work is to theoretically capture CO2 from biogas containing 40 % CO2 and 60 % CH4, and to obtain high purity of CH4 production. The model predictions are then validated with the experimental data obtained from the literature using TEA, DEA, and MEA as the absorbent solutions.
2. Mass transfer theory A material balance was implemented on the gas-liquid membrane contactor for developing and solving the governing equations. The model was divided into three domains, i.e., the tube side, the membrane, and the shell side. Figure 1 demonstrates the model domains of the membrane module in this study. Regarding the transport process of CO2 in the HFMC, biogas containing of CO2 and CH4 flows into the shell side of the module, and passes through the membrane section in the non-wetted mode. Then CO2 reacts with absorbents in the lumen side, and flows out at the outlet of tube. The model is developed considering the following assumptions: (1) steady state and isothermal condition; (2) Henry’s law for the gas-liquid interface; (3) ideal gas behavior; (4) operated in the non-wetting mode. The set of governing equations are then numerically solved in these simulations.
2.1. Mass transfer equations with chemical reaction in the tube
The material balance for any species in the tube domain considering the chemical reaction can be written as , 1 , , , , + + + = , ( )
(1)
The reaction mechanisms and kinetic parameters between CO2 and chemical solutions are shown
in Tables 1 and S1 (Supporting Information). The velocity profile of the fluid in the fibers is assumed to follow Newtonian laminar flow33
r 2 Uz,tube (r)=2Uav,tube 1- " # r1
(2)
2.2. Mass transfer equations through the membrane
In the case of the non-wetted mode, the steady-state continuity equation for gas diffusion in the membrane is ∂CCO2,mem 1 ∂CCO2 ,mem ∂2 CCO2,mem DCO2,mem + + =0 ∂r ∂r r ∂z2
(3)
2.3. Mass transfer equations in the shell
The steady-state continuity equation for gas transport in the shell side can be expressed as follows ∂CCO2 ,shell 1 ∂CCO2,shell ∂2 CCO2 ,shell ∂CCO2,shell DCO2,shell + + =Uz-shell ( ) 2 ∂r r ∂r ∂z ∂z
(4)
Assuming Happel’s free surface model, the laminar parabolic velocity distribution in the shell side may be written as34 (r⁄r3 )2 -(r2 ⁄r3 )2 +2ln(r2 ⁄r) r2 2 Uz,shell =2av,shell 1- " # r3 3+(r2 ⁄r3 )4 -4(r2 ⁄r3 )2 +4ln(r2 ⁄r3 )
(5)
3. Numerical solution procedure The governing equations for CO2 absorption in a HFMC related to the tube, membrane, and shell region with the appropriate boundary conditions (listed in Table S2 in the Supporting Information)
were solved numerically using the software of COMSOL Multiphysics. This software ran finite element analysis which was combined with error control and adaptive meshing using the numerical solver of PARDISO. The solver is thread-safe, high-efficiency, robust, memory efficient and easy to use for solving the differential equations. A computer (Intel® CoreTM i5-4200U CPU @ 1.60 GHz and 4.00 GB RAM) was employed to solve the set of equations. The computational time for solving the mathematical model was about 3 minutes. It should be noted that the software generated triangular meshes in the domain. A large amount of elements (52938 elements) were then created with scaling. Table 2 lists the transport properties of the gas-liquid system used in this work.
4. Results and discussion
4.1. Model validation
The developed mass transfer model was validated with the experimental data reported in the literature50. The membrane module specifications and operating conditions used in this study are illustrated in Table 3. Three aqueous liquid solutions of 20 wt% TEA, DEA, and MEA were used as the absorbents for the validation. As listed in Table S3 (Supporting Information), the membrane performances of different values of CO2 content in the biogas are presented. The CO2 content in the gas mixture was investigated from 30 to 50 %. The error variations for CO2 absorption efficiency, flux, and CH4 recovery are from 0.12 to 6.14 %, from 0.63 to 9.68 %, and from 1.63 to 9.84 %, respectively. Since the maximum error is blow 10 % in these simulations, the proposed model is reliable for the simulation of biogas upgrading in a HFMC. It is also noted that increasing the CO2
content adversely affects the module performance. This is due to the fact that the mass transfer resistance at the boundary layer of gas phase decreases with an increase in CO2 content. As described, it is noted that the three indexes have slight change since the driving force for the mass transfer process is mainly provided by liquid phase. Thus, CO2 content has an insignificant effect on the biogas upgrading technique in actual operations. The impacts of this factor in actual production could be negligible. The previous work51 confirms the current study.
In addition, the model validation for CO2 removal efficiency, CO2 flux, and CH4 recovery as a function of gas phase velocity is plotted in Figure 2. The liquid concentration was 10 wt%. It should be pointed out that with an increase of gas velocity the CO2 removal efficiency and CH4 recovery noticeably decrease. The decreasing rates of MEA and DEA are higher than that of TEA due to higher reaction rate with CO2. However, the CO2 flux dramatically increases when the gas velocity increases. This is due to the fact that the increment of gas velocity enhances the mass transfer process in the membrane contactor. It can be concluded that lower velocity of the gas provides a high-purity CH4 production which is good to be a substitute of conventional fossil fuels. According to this figure, the mass transfer model could well predict the biogas upgrading process in the membrane module. It is also found that the model results are higher than the corresponding values in the experiments. This is caused by the long-term operation and membrane wettability issues in the experimental tests which deteriorate the membrane performance.
Figure 3 illustrates the concentration gradient of gas phase outside fibers. The dimensionless concentration of CO2 along the membrane length has a downward trend from 1.0 to 0.14. The flux vectors also show the flow process of biogas in the shell side of the membrane contactor. As observed, the binary gas mixture enters into the module from the inlet of the shell (z = L) and then diffuses into the fibers via the pores. The gas mixture after the removal process emits at the outlet of the shell for the following regeneration procedure. As shown in Figure 4, the contours of liquid concentration using 10 wt% PA solution as the absorbent in the tube side are plotted. The absorbent is fed into the tube from inlet of the membrane module (z = 0). It is noted that the PA concentration decreases along the module length in the tube side due to the increment in the amount of absorbed CO2. Additionally, the concentration of absorbent in the center of the module is the highest, and then it gradually decreases to the tube-membrane interface.
4.3. Concentration profile in the axial direction
Figure 5a depicts the dimensionless concentration profiles of the absorbents in axial direction at the gas-liquid contact interface. It can be seen from the figure that PA has the best absorption performance of CO2 among them, and the outlet concentration ratio reaches nearly zero due to more solutions reacting in the reaction. Instead, the outlet dimensionless TEA concentration is around 0.8 which shows poor absorption performance. Figure 5b shows the axial dimensionless concentration of gas at the membrane-shell interface using the four liquid solutions. All of these four curves show a similar downward trend along the membrane module length. The outlet CO2 dimensionless
concentrations with TEA, DEA, MEA, and PA are 0.86, 0.39, 0.21, and 0.14, respectively.
4.4. Concentration profile of CO2 in the radial direction
As illustrated in Figure 6, the CO2 concentration profiles for various solvents at z/L=0.5 in r direction of the membrane contactor are plotted. With respect to this figure, the concentration value drops from the shell to the tube due to the concentration difference of CO2. The reason is due to the fact that the diffusivity of CO2 in the shell is much larger than that in the tube or membrane section. In particular, the CO2 concentration has a significant drop cross the membrane side, and then reaches the minimum at the center of the tube due to the chemical reaction.
4.5. Effect of operating pressure
In this section, the operating pressures in the range of 0.1 and 3 MPa were investigated. 5 wt% PA solution was used as the absorbent. When considering the operation at high pressures, the equation for CO2 diffusivity in the shell side may be written as52 %&' ,()** = 10
8
. 0.-2 31⁄4%&' + 1⁄4%56 7' ,- / 8
8
9 :3;%&' 7< + 3;%56 7< =
(6)
where the diffusion volumes of CO2 and CH4 are 26.9 and 24, respectively53. It can be clearly seen from Figure 7 that a high pressure provides better separation performance due to the enhancement in the gas mass transfer process18. The flux of CO2 considerably increases to 20.4 from 12.3 mol m-3 h-1. At the same time, the CO2 removal efficiency and CH4 recovery increase from 69.1 to 100 %, and from 84.6 to 100 %, respectively. When the pressure is 0.9 MPa, the percentage of CH4 recovery
reaches higher than 95 %. It indicates that the biogas upgrading process operated at 0.9 MPa could meet the requirements of the CH4 production as using 5 wt% aqueous solvent. Generally, other parameters like absorbent concentration and liquid velocity could be changed to improve the contactor performance under atmospheric conditions.
4.6. Effect of fluid condition
In terms of turbulent condition, the steady state material balance for CO2 absorption in the tube side is given by54 , 1 , , , > (, + , ) + + + = , ( ) > %& ' ,
8
8 '
0 0.08, 0.3164H ,6 = × D I > ? 8 BC%& '
(7)
(8)
The comparisons of separation performance between the laminar and turbulent flows are depicted in Figure 8. As illustrated from the series of bar charts, it is found that the membrane performance of a turbulent flow is better than that of a laminar flow due to the increment in the diffusion process of liquid in the tube. For physical absorption, the CO2 absorption efficiency, CO2 flux and CH4 recovery in a turbulent flow dramatically increase by 14.6 %, 6.6 % and 3.0 mol m-3 h-1 in comparison with a laminar flow, respectively. In terms of chemical absorption, DEA, MEA and PA reach the complete removal when considering turbulent fluid conditions. Thus, higher absorption efficiency could be obtained when operating under turbulent flow conditions.
Figure 9 shows the influence of fluid direction on the membrane performance at various absorbent velocities. It is clearly seen that the counter-current flow gives better contactor performance with the same liquid velocity than the co-current flow configuration. The co-current flow pattern indicates that with an increase in absorbent velocity, the CO2 removal efficiency increases from 65.3 to 85.1 %, and the percentage of CH4 recovery increases from 86.1 to 94.0 %. For the counter-current model, the CO2 removal efficiency increases to 95 from 77.4 %, while the recovery of CH4 increases to 98.0 from 91.0 %. This is due to the reason that the counter-current flow condition provides greater driving force for the CO2 transport in the HFMC.55
4.8. Effect of serial modules Very few studies focusing on the serial modules have been conducted in previous literature55,56. As depicted in Figure 10, CO2 dimensionless concentration along the module length for single and two modules is plotted. The outlet CO2 dimensionless concentration for the first module is 0.38, whereas this value decreases to 0.14 at the outlet of the second module. This can be explained by the fact that the gas-liquid contact area increases, which further increases the amount of absorbed CO2 in the liquid phase55. Figure 11 shows the modeling results from different absorbent solutions using one module or two modules in series. It should be noted that the percentages of CO2 removal and CH4 recovery moderately increase when using two modules. Particularly, the solution which has a high reaction rate with CO2 gives superior module performance. Regarding MEA and PA solutions, the CO2 removal efficiency and CH4 recovery increase about 25 and 10 % compared to the use of one
5. Conclusions In the present study, the simulations of biogas purification in a HFMC were studied. A wide-ranging 2D mass transfer model for CO2 removal and CH4 recovery was proposed in the non-wetted conditions. PA aqueous solutions with different concentrations were used as the absorbents in the gas-liquid system. The model predictions showed good consistency with the experimental data obtained from previous literature. The effects of operating pressure, flow pattern, flow condition, and serial module on the removal of CO2 were investigated. The simulated results showed that operation under high pressure could benefit the biogas upgrading process. In the case of 5 wt% PA solution, the percentage of CH4 recovery was higher than 95 % when the operating pressure reached 0.9 MPa. The turbulent flow gave higher CO2 removal efficiency and CH4 production purity than the laminar flow. Theoretical results also indicated that the membrane performance in the counter-current condition was better than that in the co-current condition. Additionally, when two modules in series were employed, the absorption performance was obviously improved. Thus, in this work the established mathematical model was able to simulate the gas removal process using the membrane technology. It was also shown to be an effective method instead of traditional separation techniques applied in the biogas purification field.
Author Contributions The manuscript was written through contributions of all authors. All authors have given approval to the final version of the manuscript.
Funding Sources The authors would like to thank for financial support from Fundamental Research Funds for the Central Universities (No. CDJZR14145501), National Natural Science Foundation of China (No. 50906103), Chongqing Science and Technology Talent Training Plan (No. cstc2013kjrc-qnrc90002) and China National Tobacco Corp Chongqing Branch (No. NY20130501010010). Notes The authors declare no competing financial interest.
= overall reaction rate of any species (mol m-3 s-1)
Re
= Reynolds number
Sc
= Schmidt number
T
= temperature (K)
U
= velocity (m s-1)
V
= diffusion volume
Greek Letters ε
= membrane porosity
τ
= membrane tortuosity
θ
= packing density
δ
= membrane thickness (m)
η
= removal efficiency or recovery (%)
Subscripts av
= average
G
= gas
in
= inlet
L
= liquid
mem
= membrane
T
= Turbulence
out
= outlet
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Tables and Figures Table 1. Kinetic Parameters for the Reaction between CO2 and Absorbent Table 2. Diffusion, Solubility Coefficients and Other Constants Used in Calculations Table 3. Parameters of the Membrane Contactor and Operating Conditions Used in this Simulation50 Figure 1. A schematic diagram of the membrane module for non-wetting. Figure 2. Effect of gas velocity on the module performance: (a) CO2 removal efficiency, (b) CO2 flux, and (c) CH4 recovery (line represents simulation results and symbol denotes experimental data; initial CO2 concentration: 40%; absorbent concentration: 10 wt%; UL: 0.06 m s-1). Figure 3. A presentation of dimensionless concentration distribution of CO2 inside the shell (initial CO2 concentration: 40%; absorbent concentration: 10 wt%; UG: 0.32 m s-1; UL: 0.06 m s-1). Figure 4. A presentation of the contours of PA solution inside the tube (initial CO2 concentration: 40%; absorbent concentration: 10 wt%; UG: 0.32 m s-1; UL: 0.06 m s-1). Figure 5. Dimensionless concentration of (a) absorbent and (b) CO2 along the length of membrane contactor (initial CO2 concentration: 40%; absorbent concentration: 10 wt%; UG: 0.32 m s-1; UL: 0.06 m s-1). Figure 6. Concentration profile of CO2 in the radial direction at z/L=0.5 (initial CO2 concentration: 40%; absorbent concentration: 10 wt%; UG: 0.32 m s-1; UL: 0.06 m s-1). Figure 7. Effect of operating pressure on the HFMC performance for various absorbents (initial CO2 concentration: 40%; PA concentration: 5 wt%; UG: 0.32 m s-1; UL: 0.06 m s-1). Figure 8. Effect of flow condition on the HFMC performance for various absorbents (initial CO2 concentration: 40%; absorbent concentration: 10 wt%; UG: 0.32 m s-1; UL: 0.06 m s-1).
Figure 9. Effect of flow configuration on the HFMC performance for various absorbents (initial CO2 concentration: 40%; absorbent concentration: 10 wt%; UG: 0.32 m s-1). Figure 10. Dimensionless concentration distribution of CO2 along the module length in the shell side of HFMCs (initial CO2 concentration: 40%; absorbent concentration: 5 wt%; UG: 0.32 m s-1; UL: 0.06 m s-1). Figure 11. Effect of serial module on the HFMC performance for various absorbents (initial CO2 concentration: 40%; absorbent concentration: 5 wt%; UG: 0.32 m s-1; UL: 0.06 m s-1).
Figure 2. Effect of gas velocity on the module performance: (a) CO2 removal efficiency, (b) CO2 flux, and (c) CH4 recovery (line represents simulation results and symbol denotes experimental data; initial CO2 concentration: 40%; absorbent concentration: 10 wt%; UL: 0.06 m s-1).
r Figure 3. A presentation of dimensionless concentration distribution of CO2 inside the shell (initial CO2 concentration: 40%; absorbent concentration: 10 wt%; UG: 0.32 m s-1; UL: 0.06 m s-1).
r Figure 4. A presentation of the contours of PA solution inside the tube (initial CO2 concentration: 40%; absorbent concentration: 10 wt%; UG: 0.32 m s-1; UL: 0.06 m s-1).
Figure 5. Dimensionless concentration of (a) absorbent and (b) CO2 along the length of membrane contactor (initial CO2 concentration: 40%; absorbent concentration: 10 wt%; UG: 0.32 m s-1; UL: 0.06 m s-1).
Figure 6. Concentration profile of CO2 in the radial direction at z/L=0.5 (initial CO2 concentration: 40%; absorbent concentration: 10 wt%; UG: 0.32 m s-1; UL: 0.06 m s-1).
Figure 7. Effect of operating pressure on the HFMC performance for various absorbents (initial CO2 concentration: 40%; PA concentration: 5 wt%; UG: 0.32 m s-1; UL: 0.06 m s-1).
Figure 8. Effect of flow condition on the HFMC performance for various absorbents (initial CO2 concentration: 40%; absorbent concentration: 10 wt%; UG: 0.32 m s-1; UL: 0.06 m s-1).
Figure 9. Effect of flow configuration on the HFMC performance for various absorbents (initial CO2 concentration: 40%; absorbent concentration: 10 wt%; UG: 0.32 m s-1).
Figure 10. Dimensionless concentration distribution of CO2 along the module length in the shell side of HFMCs (initial CO2 concentration: 40%; absorbent concentration: 5 wt%; UG: 0.32 m s-1; UL: 0.06 m s-1).
Figure 11. Effect of serial module on the HFMC performance for various absorbents (initial CO2 concentration: 40%; absorbent concentration: 5 wt%; UG: 0.32 m s-1; UL: 0.06 m s-1).
