CO2 Removal in Packed-Bed Columns and Hollow-Fiber Membrane

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CO2 removal in packed-bed columns and hollow-fiber membrane reactors - Investigation of reactor performance I. Iliuta, and Maria C. Iliuta Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.5b03454 • Publication Date (Web): 19 Nov 2015 Downloaded from http://pubs.acs.org on November 24, 2015

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Industrial & Engineering Chemistry Research

CO2 removal in packed-bed columns and hollow-fiber membrane reactors Investigation of reactor performance Ion Iliuta and Maria C. Iliuta* Department of Chemical Engineering, Laval University, Québec, Canada, G1V 0A6

Abstract CO2 absorption performance of hollow-fiber membrane and conventional packed-bed column reactors under similar operating conditions was evaluated. Two-scale, non-isothermal, steady-state models were used to simulate the reactors behaviour. The membrane reactor model accounts for CO2 diffusion in gas-filled membrane pores, CO2 and amine diffusion accompanied by chemical reaction in liquid-filled membrane pores and CO2 and amine diffusion accompanied by chemical reaction in the liquid film zone surrounding the inside membrane wall. The packedbed column reactor model interconnects a two-fluid 2D hydrodynamic platform with 2D mass and energy transport equations in the gas and liquid phases and nonlinear differential equations governing diffusion and reaction in the liquid film. In the absence of membrane wetting, the hollow-fiber membrane reactor outperforms the packed-bed column reactor with similar volume and specific surface area. This is not the case under membrane wetting conditions, when at low specific surface areas the packed-bed column reactor can outperform the membrane reactor. However, as the hollow-fiber membrane reactors can be stacked with very high specific surface areas in the same reactor volume, the performance of this type of reactor remains better even under partial wetted membrane conditions. Keywords: CO2 absorption; hollow-fiber membrane reactor, packed-bed column reactor, performance *corresponding author ([email protected])

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1. Introduction The energy needs primarily derived from economic development and population growth are rising significantly. According to 2013 Intergovernmental Panel on Climate Change (IPCC) report,1 the concentration of CO2 (the main contributor to global warming) in the atmosphere has increased by 40% since the beginning of industrial era, mainly on the account of the combustion of fossil fuels and various industrial activities, especially oil refineries, cement, aluminum and steel production. Besides potential strategies like the reduction of energy consumptions, improvement of energy efficiency, and implementation of substitute fuels (e.g., hydrogen) or renewable energy sources, CO2 capture from different industrial gas mixtures has lately generated a large interest in the GHG (greenhouse gas) mitigation. Among the large variety of CO2 capture tools (absorption, adsorption and membrane techniques),25

the CO2 absorption process in countercurrent packed-bed columns is a mature technology for

CO2 capture with very high removal efficiency (up to 90%).5 Although conventional packed-bed columns are traditionally the most extensively used in industry, they have several drawbacks including elevated capital cost, significant space occupancy, important tendency for corrosion, sensitivity to motion, and a diversity of operational problems such as flooding, entrainment, liquid channeling and foaming.6-8 One promising alternative to packed-bed columns is the gas–liquid membrane contactors9,10 which involves the transfer of CO2 within a porous membrane and the absorption in the liquid phase. The gas-liquid membrane contactors combine therefore the advantages of absorption (large selectivity) and membrane separation (compact configuration and modularity)11 and allow the independent control of liquid and gas flowrates, thereby eliminating the operational problems seen in packed columns. Also, the modular nature of membrane systems


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permits high specific surface area, reduced size of equipment, large and stable gas/liquid interfaces, high mass transfer rates, operational flexibility, easy design and scale-up or scaledown.12 However, the diffusion in the membrane pores enhances the mass transfer resistance, which can be significant when the membrane pores are wetted by the liquid phase,10,13 which leads to the decline of the efficiency of CO2 absorption process. The comparison of CO2 removal performance of membrane reactors and conventional packed-bed column reactors has been reported in the literature by several investigators. Nii et al.14 studied the absorption of CO2 in an absorption membrane system using NaOH (sodium hydroxide) and K2CO3 (potassium carbonate) aqueous solutions. Their results indicate comparable absorption rates in the membrane contactor and a random packed-bed column with Raschig rings. CO2 removal performance of a polypropylene microporous hollow-fiber membrane module was investigated by Karoor and Sirkar8 for CO2 absorption in water. The values of the mass transfer coefficients attained in the membrane contactor were approximately 5 times higher compared to literature data for different random packed-bed columns. Nishikawa et al.15 studied CO2 absorption in aqueous monoethanolamine (MEA) solutions in a membrane contactor with polytetrafluoroethylene (PTFE) and polyethylene (PP) membranes and obtained comparable results: mass transfer values 5 times higher than those available for random packed-bed columns with ceramic Berl saddles and steel rings.16 Rangwala17 studied the CO2 removal in water, NaOH and diethanolamine solutions in a membrane module containing polyethylene membranes. Their results reported mass transfer values from 3 to 9 times larger in comparison with literature data for a random packed-bed column with Raschig rings. deMontigny et al.18 used monoethanolamine and 2-amino-2-methyl-1-propanol aqueous solutions for CO2 capture and evaluated the performance of an absorption membrane module with microporous polytetrafluoroethylene and polyethylene 3 ACS Paragon Plus Environment


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membranes and a structured packed column containing Sulzer DX via the overall mass transfer coefficient. Compared to the structured packed-bed column, the absorption membrane systems tested produced overall volumetric gas-phase mass transfer coefficients up to 4 times higher. Polytetrafluoroethylene membranes outperformed the polyethylene membranes which are known to be more easily wetted by amine solutions. Although the advancement accomplished up till now in the comparison of CO2 removal performance in membrane and conventional packed-bed column reactors are noticeable, these comparisons should be considered prudently because the two systems were not operated under similar operating conditions. The aim of this work is to attempt to eliminate some of the ambiguity by investigating CO2 removal performance of hollow-fiber membrane reactors and conventional packed-bed column reactors with similar reactor volume and specific surface area and operated under similar conditions, and allows a more consistent and appropriate performance comparison which permits to establish the more attractive reactor from economic point of view. CO2 removal by MEA was chosen in the assessment and CO2 removal performance was compared via the exit CO2 mole fraction. The packed-bed column contained metal and ceramic Pall rings, ceramic Berl saddles and ceramic Raschig rings and the hollow-fiber membrane reactor tested polytetrafluoroethylene membranes. The influence of membrane wetting fraction on the hollowfiber membrane reactor performance was considered. 2. Modelling framework 2.1. Membrane contactor model The two-scale, non-isothermal, steady-state model developed by Iliuta et al.19 accounting for CO2 diffusion in the gas-filled membrane pores, CO2 and amine diffusion accompanied by the chemical reaction in the possible liquid-filled membrane pores and CO2 and amine diffusion 4 ACS Paragon Plus Environment

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accompanied by the chemical reaction in the liquid film zone surrounding the inside membrane wall was applied to investigate CO2 removal in the hollow-fiber membrane contactor. The mathematical model has been presented elsewhere (Iliuta et al.19) and will only be described briefly in the following sections. 2.1.1. Porous membrane scale model. The porous membrane scale model describes the diffusion of CO2 (A) in the gas-filled membrane pores and CO2 diffusion accompanied by the chemical reaction in the liquid-filled membrane pores (Figure 1):19  1 ∂  ∂C Ag,m DAeff, g   r r ∂ r  ∂r 

   = 0  

(1)

 1 ∂  ∂C Al ,m DAeff,l   r r ∂ r  ∂r 

 l   − RA ( C j , m ) = 0  

(2)

The boundary conditions are (gas-liquid interface is located inside the membrane): r = Rmout

r=R

gl m

(

k g C A, g − C Ag, m

D

eff A, g

C Ag,m

r = Rmin

DAeff,l

r = Rmout

∂C Ag, m ∂r

r = Rmgl

=D

eff A, l

= C Al , m

r = Rmg l

∂r

(3) r = Rmout

∂r

(4) r = Rmg l

1 m

= D A, l r = Rmin

∂C Ag, m

∂C Al ,m

r = Rmgl

∂C Al ,m ∂r

)

= − DAeff, g

(5)

∂C A,lf ∂r

(6) r = Rmin

The amine (B) diffusion accompanied by the chemical reaction in the wetted membrane pores is described by the following steady-state mass balance equation:19



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 1 ∂  ∂CBl ,m DBeff,l   r  r ∂r  ∂r r=R

gl m

r=R

in m

 l   − RB ( C j , m ) = 0   eff B ,l

D

eff B ,l

D

∂CBl ,m ∂r

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(7)

=0

(8)

r = Rmg l

∂CBl ,m ∂r

= DB ,l

∂CB ,lf ∂r

r = Rmin

(9) r = Rmin

The heat transport in the dried membrane pores and the heat transport and reaction in the wetted membrane pores are described by the following equations:20

λ

eff m, g

 1 ∂  ∂Tmg  r  r ∂r  ∂r

  = 0 

(10)

 1 ∂  ∂Tml   l r   + RA ( C j , m ) ( −∆H RA ) = 0 r ∂ r ∂ r   

λmeff,l 

r = Rmout

r=R

gl m

(

α g Tg − Tmg

λ

eff m, g

Tmg

r=R

in m

λ

∂Tmg ∂r

r = Rmgl

eff m ,l

∂Tml ∂r

r = Rmout

)

= −λmeff, g

=λ

eff m ,l

r = Rmgl

= Tml

∂Tmg ∂r

(12) r = Rmcat

(13) r = Rmgl

(14)

r = Rmgl

= λl r = Rmin

∂Tml ∂r

(11)

∂Tlf ∂r

(15) r = Rmin

Under membrane gas-filled pores conditions (Figure 2) the mathematical model is reduced to the Eqs.1 and 10.19 2.1.2. Liquid film scale model. The model for the liquid film zone surrounding the inside membrane wall was developed via the film theory:21 6 ACS Paragon Plus Environment

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 1 ∂  ∂C A,lf   DA , l  r   − RA ( C j ,lf ) = 0 ∂r    r ∂r 

(16)

 1 ∂  ∂CB ,lf   DB ,l  r   − RB ( C j ,lf ) = 0 r ∂ r ∂ r   

(17)

The heat transport equation in the liquid film is:19  1 ∂  ∂Tlf  r  r ∂r  ∂r

λl 

   + RA ( C j ,lf  

) ( −∆H ) = 0

(18)

RA

The boundary conditions for the liquid film model are: (i)

membrane pores partially filled with liquid

r=R

in m

D

eff j ,l

λ

eff m ,l

r = R lf (ii)

r = Rmin

∂C lj ,m ∂r ∂Tml ∂r

C j ,lf

∂C j ,lf ∂r

r = Rmin

= λl r = Rmin

where j=A,B

(19)

r = Rmin

∂Tlf

(20)

∂r

= C j ,l

r = Rlf

r = Rmin

Tlf

r = Rlf

= Tl

where j=A,B

(21)

membrane pores totally filled with gas (gas-liquid interface is located at membrane wall)

C A , lf

r = Rmin

= C Ag,m

∂CB ,lf ∂r

λmeff, g r = R lf

= D j ,l

C j ,lf

r = Rmin

m

(22)

=0

(23)

r = Rmin

∂Tmg ∂r

r = Rlf

= λl r = Rmin

= C j ,l

∂Tlf ∂r

Tlf

(24) r = Rmin

r = Rlf

= Tl

where j=A,B

(25)



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2.1.3. Membrane contactor scale model. The steady state mass balance equations for CO2 and amine in the liquid phase, assuming “concentration plug flow” under laminar flow conditions,19,22,23 are:

vsl

vsl

∂C A,l

= − DA, l

∂z ∂CB ,l

= − DB ,l

∂z

∂C A,lf ∂r

r =R

lf

r =R

lf

∂CB ,lf ∂r

av ,out − RA ( C j ,l )

(26)

av ,out − RB ( C j ,l )

(27)

CO2 mass balance equation in the gas phase (shell side) is,

±

∂ ( vsg C A, g ) ∂z

=D

eff A, g

∂C Ag,m

av ,out

∂r

(28)

r = Rmout

where gas velocity axial gradient was evaluated from the overall mass balance equation in gas phase:

±

(

Pg  ∂  g  vsg  = − k g C A, g − C A,m ∂z  RTg 

r = Rmout

)a

(29)

v , out

Heat balance equations in the liquid and gas phases are:19

ρl vsl c pl

∂T ∂Tl = −λl lf ∂z ∂r

± ρ g vsg c pg

∂Tg ∂z

= λmeff, g

av ,out + RA ( C j ,l ) ( −∆H rA )

(30)

r = R lf

∂Tmg ∂r

av ,out

(31)

r = Rmout

The corresponding boundary conditions for mass and heat balance equations are: (i) z=0

cocurrent flow (“+” in eqs. 28, 29, 31)

C A, g vsg

z =0

z =0

= C Ain, g

= vsgin

C B ,l

Tl

z =0

z =0

= CBin,l

= Tlin

(32)

Tg

z =0

= Tgin

(33)


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(ii)

countercurrent flow (“-” in eqs. 28, 29, 31)

z=H

C A, g

z =H

z=0

C B ,l

z =0

= C Ain, g = CBin,l

vsg

Tl

z=H

z =0

= vsgin

Tg

= Tlin

z=H

= Tgin

(34) (35)

2.1.4. Reaction mechanism. CO2-MEA reaction kinetics was described by the mechanism proposed by Caplow24 and Danckwerts:25

RCO2 =

k2, B C ACB 1 1+  k H 2O  k  C H 2O  +  B C B    k−1   k −1 

(36)

This mechanism supposes the creation of a zwitterion followed by the subtraction of a proton by all bases being in solution: k2  → RNH 2+ COO − RNH 2 ( B ) + CO2 ( A) ← 

(37)

kb → RNH 2+ COO − + Base ← RNHCOO − + BaseH +

(38)

k −1

k− b

The kinetic constants and model parameters estimation are given in Liao and Li,26 and Iliuta et al.19

2.2. Packed-bed column model 2.2.1. Hydrodynamic model. A two-fluid 2-D steady-state hydrodynamic model was used to simulate two-phase countercurrent flow in the packed-bed column. Both flowing fluids were considered continuous viscous Newtonian phases. The liquid was incompressible and the gas phase was ideal. The hydrodynamic model are based on the volume-averaged forms of mass and momentum balance equations.27 The discontinuity in pressure between the gas and liquid phases was described using the equilibrium constitutive equation developed by Attou and Ferschneider.28 9 ACS Paragon Plus Environment


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The interaction forces exerted on the liquid and gas phases were evaluated via the model developed by Iliuta et al.29 The continuity and momentum balance equations (z-axis oriented downwards in the direction of liquid flow) are:

u gr

∂P ∂ ∂ ρ g ε g u gz ) + u gz ( ρ g ε g u gz ) = −ε g g + µ gef ( ∂r ∂z ∂z

 1 ∂  ∂ ( ε g u gz )  ∂ 2 (ε u )  g gz   r +   ∂r ∂z 2   r ∂r   

(39)

+ε g ρ g g + fint, g , z

u gr

 ∂  1 ∂ ( rε g u gr )  ∂ 2 (ε u )  ∂Pg ∂ ∂ g gr ef   + f int, g ,r (40)  + u + u u = − + ρ ε ρ ε ε µ ( ( g g gr ) gz g g gr ) g g   ∂r ∂z ∂r ∂r ∂z 2   ∂r  r  

ρ lulr

∂P ∂ ∂ ( ε lulz ) + ρlulz (ε lulz ) = −ε l l + µlef ∂r ∂z ∂z

 1 ∂  ∂ ( ε lulz )  ∂ 2 (ε lulz )    r + ∂r  ∂z 2   r ∂r 

(41)

+ε l ρ l g + f int,l, z

ρ lulr

∂P ∂ ∂ ( ε lulr ) + ρlulz (ε lulr ) = −ε l l + µlef ∂r ∂z ∂r

 ∂  1 ∂ ( rε lulr )  ∂ 2 (ε l ulr )     + fint,l ,r + ∂r ∂z 2    ∂r  r

(42)

∂ 1 ∂ ρ g ε g u gz ) + ( ( r ρ g ε g ugr ) = 0 ∂z r ∂r

(43)

∂ 1 ∂ ( ρlε lulz ) + ( r ρlε lulr ) = 0 ∂z r ∂r

(44)

εl + εg = ε (r )

(45)

At the column inlet the liquid and gas velocities are quantified in Dirichlet-type boundary conditions. An open boundary condition was applied at the column outlet,30 and no-slip conditions were forced at the column wall. The boundary conditions are: z=0

z=H

∂u g , z ∂z ∂ul ,z ∂z

=

=

∂u g ,r ∂z ∂ul ,r ∂z

=0

ul, z = ulin

ul , r = 0

ε l = ε lin

Pg = Pl = P 0

=0

ug, z = ugin

ug, r = 0

∂ε l =0 ∂z

∂Pg ∂z

=

∂Pl =0 ∂z

(46)

(47)


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∂u g , z

r=0

∂r

r=R

=

∂ul , z ∂r

=0

u g , z = ul , z = 0

u g , r = ul , r = 0

∂ε l =0 ∂r

∂Pg

u g , r = ul , r = 0

∂ε l =0 ∂r

∂Pg

∂r

∂r

=

∂Pl =0 ∂r

(48)

=

∂Pl =0 ∂r

(49)

The volume-averaged liquid-solid and gas-liquid interfacial forces were estimated via the model developed by Iliuta et al.29 The assumption of completely packing wetting entrains that the interaction forces exerted on the gas phase will have only contributions due to effects localized at the gas-liquid interface. Thus, the interfacial force exerted on the gas phase ( f int, g , z ( r ) = Fg l , z ( r ) ) is given by the drag force exerted on the gas phase because of the relative motion between the two fluids

( F gl ).

(f

= − Fg l , z ( r ) − Fls , z ( r ) ) the drag force experienced by the liquid due to the shear stress

int,l , z ( r )

The

interaction

force

exerted

on

the

liquid

phase

includes

nearby the liquid-solid interface ( Fls ) and the gas-liquid interfacial drag due to the slip between fluids ( F gl ).

Fls , z

 E1 as2 µl E2 as 2 2  = + ρ u + u l l l z r  ul z ε l 2 6 εl  36 ε l 

(50)

Fls ,r

 E1 as2 µl E2 as 2 2  = + ρ u + u l l z l r  u lr ε l 2 6 εl  36 ε l 

(51)

Fg l , z

 E1 a 2s µ g E2 a s = + ρg 2 6 εg  36 ε g

 jrz2 + jrr2  jrz ε g 

2  E a s µ g E2 a s Fg l , r =  1 + ρg 2 6 εg  36 ε g

 jrz2 + jrr2  jrr ε g 

were jrz ( r ) = u gz ( r ) − ulz ( r )

(52)

(53)

The drag forces are generated for positive axial velocity defined in the gravity direction. For countercurrent flow, the gas velocity is negative in these relationships, and in the continuity and



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momentum balance equations. Finally, the difference between the liquid and gas phase pressures was modeled via the capillary pressure correlation developed by Attou and Ferschneider:28  1− ε Pg − Pl = 2σ l   1− ε g 

  

0.33

 1 1  +  d p d min

  

(54)

2.2.2. Mass and heat transport equations. The reaction zone between CO2 and MEA is confined

near the gas-liquid interface and the saturation of the bulk liquid with CO2 can be neglected.19 Consequently, the steady state mass transfer model coupled with chemical reaction is reduced to species balance equations for CO2 in the gas phase and amine in the liquid phase. The temperature gradients in the packed-bed column reactor are given by an average heat balance equation (only convection, conduction and reaction heat were considered).20 Mass balance equations for the various species and energy equation are two-dimensional and account for the non-uniform velocities profile: u gr

1 ∂  ∂PA, g ∂ ∂ ε g PA, g ) + u gz ( ε g PA, g ) = Dgr  (  rε g ∂r ∂z ∂r  r ∂r 

 ∂2 + D (ε g PA,g ) − ( N Aa ) RT  gz ∂z 2 

1 ∂  ∂CB ,l   ∂ ∂ ∂2 ulr ( ε l CB ,l ) + ulz ( ε l CB ,l ) = Dlr   rε l   + Dlz 2 ( ε lCB ,l ) −ν B ( N A a ) ∂r ∂z ∂r   ∂z  r ∂r 

(ρ

l

c plε lulr + ρ g c pg ε g ugr )

∂T ∂T  1 ∂  ef ∂T   + ( ρ l c plε lulz + ρ g c pg ε g u gz ) =  rε l λr  ∂r ∂z  r ∂r  ∂r  

(55)

(56)

(57)

+ ( −∆H rA )( N A a ) The boundary conditions are the followings: z=0

C B ,l

z =0

= CBin,l

z=H

PA, g

z=H

= PAin, g

r=0

∂PA, g ∂r

=0

T

z =0

= Tlin

(58) (59)

∂CB ,l ∂r

=0

∂T =0 ∂r

(60) 12

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∂PA, g

r=R

∂r

∂CB ,l

=0

∂r

=0

T = Tw

(61)

The model for the liquid film (considered isotherm) surrounding the gas-liquid interface was developed via the film theory:21 DA , l DB ,l

dC A,lf dx dCB ,lf

x=0

x = δl

dx

− RA ( C j ,lf ) = 0

(62)

− RB ( C j ,lf ) = 0

(63)

kg

(P RT

− PA, g ,i ) = − DA,l

C A,lf

=0

A, g

x =δ l

∂C A,lf

∂CB ,lf

∂x

∂x

x =0

CB ,lf

=0

(64)

= CB ,l

(65)

x =0

x =δ l

2.2.3. Model parameters. The molecular diffusion coefficients in the liquid phase were evaluated

using the Wilke-Chang method.31 CO2 solubility in the liquid phase was taken from Versteeg and van Swaaij.32 MEA solution thermal conductivity was evaluated using Filippov equation and Sastri method.31 The heat capacity of MEA solutions was evaluated using data by Chiu and Li,33 and the reaction enthalpy was evaluated using data from Kim and Svendsen.34 The effective thermal conductivity was evaluated with the model developed by Weekman and Myers.35 Gasliquid mass transfer parameters were estimated via Onda et al.36 correlations. The liquid and gas axial dispersion coefficients were taken from Otake and Kunugita37 and Sater and Levenspiel,38 and the ratio between the axial and radial dispersion coefficients was assumed 10.39,40 The kinetic constants are given in Liao and Li.26

3. Results and discussion 3.1. Numerical implementation



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Aspen Custom Modeler from Aspen Tech Company was used to generate the numerical platforms able to solve the gas-liquid hollow-fiber membrane contactor and packed-bed column models. In the membrane contactor a 2nd -order central finite difference method was used for the discretization in the radial direction and a 1st-order backward finite difference method was used for the discretization in the axial direction. In the packed-bed column reactor, the second-order orthogonal collocation on finite elements method was employed for the discretization in the axial and radial directions. Newton method was used to solve the two steady-state models.

3.2. Models validation Analysis of Figures 3-5 shows a good agreement between the model predictions and the experimental absorption data (axial concentration and temperature profiles) obtained in a packed-bed column at different liquid and gas flow rates, CO2 partial pressure in the gas phase and amine concentration in the liquid phase. The experimental data were taken from Tontiwachwuthikul et al.41 for the atmospheric pressure absorption of CO2 using aqueous solutions of MEA into a packed-bed column (0.1 m internal diameter and total packing height of 6.55 m) filled with 12.7 mm ceramic Berl saddles. The concentration of CO2 in the gas was varied between 11.5 and 19.5 vol % and the MEA concentration was 2.0 and 3.0 mol/L. Figures 3-5 show that there is almost no difference in exit gas CO2 concentration for the conditions of the study because of the large reactor height. Figure 4 shows that even if the liquid flow rate decreases, the local reactor performance is higher because of the higher MEA concentration (run22). On the other side, Figure 5 shows that even if the gas flow rate decreases, the local reactor performance is inferior because of the lower MEA concentration (run21). Figures 6 and 7 show that the hollow-fiber membrane reactor model predicts very well the experimental data from deMontigny et al.42 obtained for absorption of CO2 into aqueous MEA 14 ACS Paragon Plus Environment

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solutions at atmospheric pressure at different liquid flow rates, CO2 partial pressure in the gas phase and MEA concentration in the liquid phase (see also Table 1). The absorption rate is marginally affected by the liquid velocity (Figures 6 and 7) and the temperature amplification in the membrane reactor decreases insignificantly with the increase of liquid flow rate (Figure 8). So, the raise of the temperature in the membrane module (experimental: 0.3-0.4 K; model: 0.10.25K) remains trivial for superficial liquid and gas velocities (0.02-0.05 m/s and, respectively, 0.25 m/s) comparable to velocities usually attained in packed-bed columns and the membrane module operation can be considered as quasi-isothermal.

3.3. Hollow-fiber membrane reactor vs. packed-bed column reactor CO2 absorption performance of the hollow-fiber membrane reactor in terms of exit CO2 mole fraction is compared with the performance of a conventional pilot countercurrent packed-bed column reactor operated under similar operating conditions. The operating conditions and the characteristics of membrane, membrane reactor and packed-bed column reactor are given in Tables 2 and 3. The membrane reactor contains microporous PTFE hollow fibers and the packed-bed column is filled with 15 mm ceramic Raschig rings (packing specific area = 310 m2/m3), 12 mm porcelain Berl saddles (packing specific area = 220 m2/m3), 35 mm ceramic Pall rings (packing specific area = 155 m2/m3) and 50 mm metal Pall rings (packing specific area = 105 m2/m3). CO2 removal by MEA was chosen as a case study. The membrane reactor is operated under countercurrent flow conditions with the gas phase flowing within the shell side and the liquid flowing in the lumen side. Also, the packed-bed column is operated under countercurrent flow conditions with the gas phase flowing toward the column top and the liquid phase flowing toward the column bottom.



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The isothermal performance of the membrane and packed-bed column reactors as a function of membrane/packing specific surface area is expressed in Figure 9a. In the absence of membrane wetting, the hollow-fiber membrane reactor outperforms the packed-bed reactor with similar volume and specific surface area due to reduced values of gas-liquid interfacial area (Figure 9b) and overall volumetric mass-transfer coefficient (Figure 10) attained in packed-bed reactor. The degree of improvement of CO2 absorption performance in the membrane reactor can be more important when the packing density/specific surface area, in the same reactor volume, is increased. In theory, specific surface area can be larger than 1000 m2/m3 and consequently, the membrane reactor performance will be considerably enhanced (not shown), but both gas and liquid side pressure drops become very important. Therefore, the packing density of fibers must be limited.43 Figure 9a shows a high performance of the membrane reactor when the inside specific surface area has reasonable values: 400-535 m2/m3 (the maximum number of fibers was evaluated with the correlation presented in Taborek et al.44). The performance of the membrane reactors is proportional to the value of the specific surface area available for mass transfer: Figure 11 shows that CO2 absorption process is amplified when specific surface area increases, even if the mass transfer flux diminishes. Figure 12 shows an important impact of membrane wetting on the membrane reactor performance. Partial wetting of membrane pores gives an important intensification of the mass transfer resistance, even if the pores wetted fraction is small,19,45 and the result is a significant diminution of CO2 absorption flux. Under partial wetted membrane conditions, the performance of hollow-fiber membrane reactor deteriorates significantly at low specific surface areas and the packed-bed column reactor can perform better than the membrane reactor (Figure 13). However,


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because the hollow-fiber membrane reactor allows high specific surface areas, the performance of this type of reactor remains better even under partial wetted membrane conditions.

4. Conclusion CO2 absorption performance of hollow-fiber membrane reactors and conventional packed-bed column reactors under similar operating conditions was examined. Two-scale, non-isothermal, steady-state models were used to simulate the comportment of the membrane and packed-bed column reactors. The membrane reactor model accounts for CO2 diffusion in the gas-filled membrane pores, CO2 and amine diffusion accompanied by the chemical reaction within the liquid-filled membrane pores and CO2 and amine diffusion accompanied by the chemical reaction in the liquid film zone surrounding the inside membrane. The packed-bed column reactor model interconnects a two-fluid 2D hydrodynamic platform with 2D mass and energy transport equations in the gas and liquid phases and nonlinear differential equations governing diffusion and reaction in the liquid film. In the absence of wetting, the hollow-fiber membrane reactor outperforms the packed-bed column reactor with similar volume and specific surface area. The membrane reactor absorption performance is improved significantly when the membrane specific surface area, in the same reactor volume, is increased. The membrane wetting significantly reduces the hollow-fiber membrane reactor performance, even at low wetting fractions: at low specific surface areas the packed-bed column reactor can outperform the membrane reactor. However, as the hollow-fiber membrane reactor allows very high specific surface areas, the performance of this type of reactor remains better even under partial wetted membrane conditions.



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Nomenclature a

gas-liquid interfacial area, m2/m3

as

packing specific surface area (surface packing area/column volume), m2/m3

avin

inside specific area, m2/m3

avout

outside specific area, m2/m3

cp,α

specific heat capacity of α-phase ( α = g , l ), J/kgK

Cj

concentration of species j, kmol/m3

d min

inner diameter of hollow fiber membrane, m

d mout

outer diameter of hollow fiber membrane, m

dmin

minimum equivalent diameter of the area between three contacting spheres, d min

 3 1 =  −  π 2 

1/ 2

dp

dp

particle diameter, m

D

reactor diameter, m

Dj,α

molecular diffusivity coefficient of species j in α phase ( α = g , l ), m2/s

D eff j ,α

effective diffusivity of species j inside membrane ( α = g , l ), m2/s

Dkj

Knudsen diffusion coefficient of species j, m2/s

E1 , E2

two-fluid model parameters, -

Fg l

gas-liquid drag force, N/m3

Fls

liquid-solid drag force, N/m3

g

gravitational acceleration, m/s2

H

reactor, membrane height, m 18 ACS Paragon Plus Environment

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kg

gas-phase mass transfer coefficient, m/s

kl

liquid-phase mass transfer coefficient, m/s

m

distribution coefficient

NA

mass transfer flux, N A = − D

NA

eff j, A

mass transfer flux, N A = − DA,l

∂C Al , m ∂r

= − DA , l r = Rmin

∂C A,lf ∂x

∂C A,lf ∂r

(membrane reactor) r = Rmin

(packed-bed column reactor) x =0

P

reactor pressure, Pa

Pj

partial pressure of species j, Pa

r

radial coordinate, m

rmout

outer hollow fiber membrane radius, m

R

ideal-gas constant or reactor radius, m

Rj

reaction rate of the component j, kmol/m3s

R lf

radius of liquid film surrounding the hollow fiber membrane, m

Rmg l

radius of gas-liquid interface in hollow fiber membrane, m

Rmin

inner radius of hollow fiber membrane, m

Rmout

outer radius of hollow fiber membrane, m

T

temperature, K

uα

interstitial velocity of α-fluid, m/s

vsα

α-phase superficial velocity, m/s

x

liquid film coordinate, m

z

axial coordinate, m 19 ACS Paragon Plus Environment


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Page 20 of 44

Greek Letters

αg

gas-phase heat transfer coefficient, kJ/m2sK

δ

membrane wall thickness, m

δl

liquid film thickness, m

∆H r

reaction enthalpy, KJ/Kmol

ε

packed bed porosity

εα

α-phase holdup

εb

porosity in the bulk region of the packed bed, -

ϕ

fiber volume fraction

λα

conductivity of α-phase, kJ/msK

λref

radial effective thermal conductivity, J/msK

µα

dynamic viscosity of α-phase, kg/m s

µαef

α-phase effective viscosity (combination of bulk and shear terms), kg/m s

νB

stoichiometric coefficient of amine in reaction

ρα

α-phase density, kg/m3

σℓ

surface tension, N/m

Subscripts/Superscripts g

gas phase

i

gas-liquid interface

in

inlet, inside

l

liquid phase


Page 21 of 44

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lf

liquid film

m

membrane

r

radial direction

s

solid phase

z

axial direction

out

outer, outside

w

reactor wall



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Page 22 of 44

Acknowledgments Financial support from Natural Sciences and Engineering Research Council of Canada (NSERC) is gratefully acknowledged.


Page 23 of 44

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Table 1. Hollow-fiber membrane module used in the experimental study of deMontigny et al.42 Operating conditions

Data

Fiber external diameter

0.002 m

Fiber internal diameter

0.001 m

Membrane wall thickness

5×10-4 m

Fiber height

3*0.122 m

Membrane porosity

0.5

Number of fibers per module

57

External specific area

581.6 m2/m3

Internal specific area

290.8 m2/m3

Module void fraction

0.709

Reactor pressure

0.1 MPa

Inert gas flow rate

31.6 kmol/m2h

Liquid flow rate

73.5-173.2 m3/m2h

Inlet liquid amine concentration

2 kmol/m3

Inlet gas CO2 mole fraction

0.095 & 0.1518



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Table 2. Hollow-fiber membrane reactor used in simulations Operating conditions

Data

Fiber outer diameter

0.002 m

Fiber inside diameter

0.001 m

Membrane wall thickness

5×10-4 m

Reactor height

1.6 m

Reactor diameter

0.3 m

Membrane porosity

0.5

Number of fibers per module

2400-12000

Outside specific area

210-1070 m2/m3

Inside specific area

105-535 m2/m3


0.92-0.6

Reactor pressure

0.1 MPa

Superficial gas velocity

0.4 m/s

Superficial liquid velocity

0.015 m/s


2 mol/L


0.2


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Table 3. Countercurrent packed-bed column reactor used in simulations Operating conditions

Data

Reactor height

1.6 m

Reactor diameter

0.3

Packing specific area

105-310 m2/m3


0.956-0.735

Reactor pressure

0.1 MPa

Superficial gas velocity

0.4 m/s

Superficial liquid velocity

0.015 m/s


2 mol/L


0.2



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Figure captions

Figure 1. Schematic diagram of absorption process in hollow-fiber membrane - membrane partially liquid-filled pores.

Figure 2. Schematic diagram of absorption process in hollow-fiber membrane - membrane gasfilled pores.

Figure 3. Experimental and theoretical axial CO2 mole fraction (a) and liquid temperature (b) profiles for CO2 absorption in MEA amine solutions at different inlet CO2 mole fractions (packed-bed column reactor; CO2-2 mol/L MEA system; liquid flow rate = 13.5 m3/m2h; air flow rate = 14.8 mol/m2s; inlet liquid temperature = 19oC; CO2 mole fraction = 0.115 (run19), 0.153 (run13), 0.195 (run15)).

Figure 4. Experimental and theoretical axial CO2 mole fraction profiles for CO2 absorption in MEA amine solutions (packed-bed column reactor; 19.5% CO2-MEA system; liquid flow rate=13.5 m3/m2h (run15), 9.5 m3/m2h (run22); air flow rate = 14.8 mol/m2s; inlet liquid temperature = 19oC; MEA concentration = 2 mol/L (run15), 3 mol/L (run22)).

Figure 5. Experimental and theoretical axial CO2 mole fraction profiles for CO2 absorption in MEA amine solutions (packed-bed column reactor; 19.1% CO2-MEA system; liquid flow rate = 9.5 m3/m2h; air flow rate = 11.1 mol/m2s (run21), 14.8 mol/m2s (run22); inlet liquid temperature = 19oC; MEA concentration = 2 mol/L (run21), 3 mol/L (run22)).

Figure 6. Experimental and theoretical axial CO2 mole fraction profiles for CO2 absorption in MEA amine solutions (hollow-fiber membrane module; CO2-2 mol/L MEA system; inert gas flow rate = 31.6 mol/m2s; CO2 mole fraction = 0.095 (run1 and run2), 0.1518 (run6 and run7)): (a) liquid flow rate = 73.5 m3/m2h; (b) liquid flow rate = 92.5 m3/m2h.


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Figure 7. Experimental and theoretical axial CO2 mole fraction profiles for CO2 absorption in MEA amine solutions (hollow-fiber membrane module; CO2-2 mol/L MEA system; inert gas flow rate = 31.6 mol/m2s; CO2 mole fraction = 0.095 (run3, run4 & run5), 0.1518 (run8, run 9 & run 10)): (a) liquid flow rate = 115 m3/m2h; (b) liquid flow rate = 146.1 m3/m2h; (c) liquid flow rate = 173.2 m3/m2h.

Figure 8. Theoretical axial temperature profiles for CO2 absorption in MEA amine solutions (hollow-fiber membrane module; 9.5% CO2-2 mol/L MEA system; inert gas flow rate = 31.6 mol/m2s; liquid flow rate = 73.5 m3/m2h (run1), 92.5 m3/m2h (run2), 115 m3/m2h (run3), 146.1 m3/m2h (run4), 173.2 m3/m2h (run5)).

Figure 9. Exit CO2 mole fraction in gas phase (a) and ratio between gas-liquid interfacial area and specific area of packing (b) vs specific area (CO2-2 mol/L MEA system; non-wetted membrane).

Figure 10. Axial overall volumetric gas-phase mass transfer coefficient profiles for CO2 absorption in membrane and packed-bed column reactors (CO2-2 mol/L MEA system; nonwetted membrane).

Figure 11. Axial CO2 mole fraction (a) and axial CO2 absorption flux (b) profiles for CO2 absorption in membrane reactor (CO2-2 mol/L MEA system; non-wetted membrane).

Figure 12. Influence of membrane wetting on the axial CO2 mole fraction profiles (a) and exit CO2 mole fraction in gas phase (b) (CO2-2 mol/L MEA system).

Figure 13. Exit CO2 mole fraction in gas phase vs. inside specific area of membrane reactor at different values of membrane wetting (CO2-2 mol/L MEA system).



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Figure 1.

Liquid-filled pores

Gas-filled pores

1/ k g 

External gas film Liquid Gas

[1/ kl ] Liquid film C A, g

C

CB ,l

g A,m

C B , lf

C Al , m

C Bl , m

C A , lf

C A,l

Rmout Rmin Rmgl R lf


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Figure 2.

Gas-filled pores

1/ k g 

External gas film Liquid Gas

[1/ kl ] Liquid film C A, g

C

C B ,l

g A, m

C B ,lf C A , lf

C A ,l Rmout

Rmin R lf



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Figure 3.


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Figure 4.



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Figure 5.


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Figure 6.



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Figure 7.


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Figure 8.



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Figure 9.


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Figure 10.



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Figure 11.


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Figure 12.



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Figure 13


CO2 Removal in Packed-Bed Columns and Hollow-Fiber Membrane

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