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Separation and Purification Technology 59 (2008) 85–90
CO2 removal from a gas stream by membrane contactor Aldo Bottino a , Gustavo Capannelli a , Antonio Comite a , Renzo Di Felice b,∗ , Raffaella Firpo a a
b
Dipartimento di Chimica e Chimica Industriale, Universit`a degli Studi di Genova, Via Dodecaneso 31, 16146 Genova, Italy Dipartimento di Ingegneria Chimica e di Processo “G. Bonino”, Universit`a degli Studi di Genova, Via Opera Pia 15, 16145 Genova, Italy Received 1 February 2007; received in revised form 28 May 2007; accepted 28 May 2007
Abstract Gas–liquid membrane contactors were applied to remove carbon dioxide from a gas stream by using an aqueous monoethanolamine (MEA) solution as absorbent. Various modules composed by different numbers of commercial polypropylene capillary membranes were constructed and tested in a laboratory-scale plant fed with a N2 –CO2 gas mixture. Attention was especially focused on the CO2 removal efficiency of the different membrane modules when gas flow rate was increased from 5 up to 360 L/h. A mathematical model was developed to simulate the absorption process in order to predict gas removal efficiency from the knowledge of the system physical parameters. The overall membrane mass transfer coefficient kM was determined and used to compare experimental and predicted removal efficiencies. A good agreement between the developed model and experimental results was found. © 2007 Elsevier B.V. All rights reserved. Keywords: Carbon dioxide; Membrane contactor; Mathematical modelling
1. Introduction Gas/liquid and liquid/liquid contacting operations are traditionally carried out by using towers or columns. The main effort in designing and operating these absorbers is to maximize the mass transfer rate by producing as much interfacial area as possible. In the case of packed columns, this is achieved by proper selection of packing material and uniform distribution of fluids fed to the packed bed [1,2]. Conventional packed bed absorbers present several disadvantages such as flooding at high flow rates, unloading at low flow rates, and channeling and foaming, that hinder the mass transfer between gas and liquid. An alternative technology that overcomes these disadvantages is a non-dispersive contactor through a microporous membrane, combining the advantages of selective absorption with membrane modularity, compact equipment and no need of a washing section [3]. Membrane contactors have been investigated extensively since the mid-1980s for a wide range of applications, including racemic leucine separation [4], removal of ethanol from fermentation broth [5,6], extraction of mevinolinic acid from fermentation broth for pharmaceutical use [7], extraction of metal ions from industrial waste and hydromet-
∗
Corresponding author. Tel.: +39 0103532924; fax: +39 0103532586. E-mail address:
[email protected] (R. Di Felice).
1383-5866/$ – see front matter © 2007 Elsevier B.V. All rights reserved. doi:10.1016/j.seppur.2007.05.030
allurgical process streams [8–11], recovery of sulfur aroma compounds from food industry wastewaters [12–14], and gas absorption [15–17]. More specifically, typical absorbers for removing carbon dioxide include bubble columns, packed towers, venturi scrubbers and sieve trays. An aqueous solution of sodium hydroxide, sodium carbonate, monoethanolamine or diethanolamine is often employed as absorbent. When a suitable membrane configuration such as the capillary one is used, the fluids to be contacted flow on the opposite side of the capillary (lumen and shell), and the gas–liquid interface forms at the mouth of each membrane pore. The available contact area remains undisturbed even at a high or low flow rate because the two fluid flows are independent from each other. This type of membrane represents a highly efficient alternative to conventional packed towers [18], because of the following reasons: no flooding at high flow rates, no unloading at low flow rates, no emulsions, no need for a density difference between fluids. In addition, scaling-up is more straightforward with the membrane contactor. Membrane operations usually scale linearly, so that a predictable increase in capacity is achieved simply by adding new membrane modules. A modular design also allows a membrane plant to operate over a wide range of capacities. The interfacial area is known and constant, allowing the performance to be predicted more easily than the conventional absorption system like packed columns, where the interfacial area per unit
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2. Mathematical modelling of the membrane contactor Nomenclature C d Da DAB DK k kG kM L n R s Sh v0 z
concentration (kmol/m3 ) mean pore size (m) Damkholer number, defined by Eq. (3) molecular diffusion coefficient (m2 /s) Knudsen diffusion coefficient (m2 /s) numerical parameter, defined by Eq. (8) gas phase mass transfer coefficient (m/s) membrane mass transfer coefficient (m/s) capillary length (m) number of capillaries capillary radius (m) membrane thickness (m) Sherwood number volumetric flow rate (m3 /s) axial distance (m)
Greek letters α dimensionless parameter ε dimensionless membrane porosity η removal efficiency, defined by Eq. (10) τ dimensionless membrane tortuosity ϕ dimensionless velocity
volume may be known, but it is often difficult to determine the loading, which represents the fraction of the available interfacial area actually used [19]. The successful use of the membrane contactor process over the conventional absorption processes will largely depend on the gas–liquid system, the types of membranes used and the operating conditions. Many authors developed mathematical models using the resistance in series concept to predict overall membrane contactor performance, combining process conditions, membrane properties and module geometric characteristics. Among them, Mavroudi et al. [20] simulated the CO2 removal in various absorbents, focusing on the possibility of partial membrane-wetting; Wang et al. [21] investigated the CO2 absorption flux by employing three typical amine solutions, flowing through the membrane hollow fiber lumen. In a previous work [22] we reported the results of screening tests aiming at selecting the more appropriate membrane material and configuration for an efficient carbon dioxide removal from flue gases. In the present study the results aimed at assessing the performance of capillary membrane previously selected for carbon dioxide removal from a CO2 –N2 gas mixture, as a function of module design (i.e., number of capillaries assembled in a given module volume) and operating parameters (i.e., gas flow rate) are presented. A mathematical one-dimensional model was developed to relate the membrane contactor efficiency to the mass transfer coefficient. The experimental parameters obtained for each membrane module were compared with the mathematical model, so that an agreement between theory and experiments could be found.
By considering a small cylindrical tube with porous wall, for fully developed laminar flow and in steady-state condition, when no reaction takes place in the fluid phase the mass balance for the generic component A can be written in dimensionless term: 2 ∗ ∗ ∂ CA 1 ∂CA ∂C∗ (1) + ϕ ∗A = α ∂z ∂r ∗2 r∗ ∂r ∗ where CA , 0 CA
∗ CA =
z∗ =
z , L
r∗ =
r , R
ϕ=
πR2 u and v0
πDAB L v0 The boundary condition at the wall can be expressed by equating the flux of the component A at the fluid–membrane interface to that through the membrane. When concentration of component A outside the membrane is negligible, its flux in the membrane can be expressed, introducing an overall mass transfer coefficient kM , by the product between this coefficient and gas concentration at the gas–membrane interface, so that we have (again in dimensionless terms) α=
∗ ∂CA ∗ = −DaCA,w ∂r ∗
at r∗ = 1
(2)
where Da is the Damkholer number, defined, for our particular system: kM R DAB
Da =
(3)
Eq. (1) represents a partial differential equation that can be solved numerically to yield concentration distribution function of axial and radial position. However, it is much easier to tackle the same problem by using a one-dimensional approach as suggested, for example, by Tronconi et al. [23]. By introducing the average component concentration over the tube section as ϕC∗ dS ∗ CA,m = A (4) ϕ dS we can write the mass balance as ∗ dCA,m
dz∗
∗ ∗ = −2α Sh(CA,m − CA,w )
(5)
with Sh =
kG R DAB
The condition at the boundary wall is this time: ∗ ∗ ∗ − CA,w ) = DaCA,w Sh(CA,m
(6)
As demonstrated by Tronconi et al. [23], this simpler onedimensional approach is practically numerically equivalent to the more complex two-dimensional approach if a correct value for the Sh is utilised: they suggested that for the whole tube
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∗ /C ∗ Fig. 1. CA,w A,m as a function of Damkholer number Da.
Fig. 2. Expected removal efficiency at various geometrical and transport properties of the membrane.
length, with the exception of a small portion at the entrance, Sh is constant and equal to 1.85. The magnitude of the Damkholer number Da (relative to the ∗ : Sherwood number, Sh = 1.85) will determine the value of CA,w ∗ it will approach CA,w for small values of Da (i.e., all the resistance to the mass transfer is concentrated on the membrane) and will approach zero for very large values of Da (i.e., the resistance is concentrated on the gas phase) as depicted also in Fig. 1. The knowledge of Da is therefore of fundamental importance and requires the estimation of the overall membrane mass transfer coefficient kM . With the simplifying assumption of a homogenous porous membrane kM is given by the ratio between the effective component diffusivity in the membrane and the membrane thickness. As it will be quantified later, this leads to a value of the Da number for our case of the order of 10−1 : ∗ from Fig. 1 we can estimate the wall concentration CA,w to be ∗ and practically coincident with the average concentration CA,m Eq. (5) therefore reduces to: ∗ dCA,m
dz∗
∗ ∗ = −2α Da CA,m = −k CA,m
(7)
with k =
87
removal efficiency of the membrane as η=
∗ ∗ CA,in − CA,out ∗ CA,in
= 1 − e−k
(10)
In practice, once the system has been determined, only gas flow rates can be adjusted in order to optimise performances. The efficiency of the system can then be obtained from: η = 1 − e−k
/ν 0
where k = 2πLRkM Fig. 2 depicts, for operating values typical of this work (maximum gas flow rate in a channel equal to 0.0001 m3 /s), expected removal efficiency at various geometrical and transport properties of the membrane. As expected, large values of k (i.e., good membrane transport properties) favour large efficiencies; at the same time increasing gas flow rate efficiency declines. This decline, however, is very sharp in a very specific range of flow rate which should, consequently, be avoided. 3. Experimental
2πLRkM v0
(8)
∗ =1 Simple integration of Eq. (7), between inlet (where CA by definition) and outlet conditions, yields expression for the dimensionless concentration at the exit of the membrane module function of the system physical characteristics: ∗ CA,out = e−k
(9)
As in this work our specific interest is in the removal of the component A from the gaseous stream, it is useful to define a
Table 1 lists the main characteristics of the capillary membranes Accurel S6/2 (Membrana, Germany), used in this study. Fig. 3 shows the scanning electron micrographs of the membrane cross-section (Fig. 3A), external surface (Fig. 3B) and internal surface (Fig. 3C). The external surface represents the contact area between the gas stream and the absorbent solution. Four modules were prepared using different numbers of capillary membranes (1, 3, 10, 18) with the same length of 17 cm, in order to obtain different contact areas. The capillary membrane ends were inserted into small plastic tubes and sealed using an epoxy resin to form a module.
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Table 1 Main characteristics of Accurel S6/2 capillary membrane supplied by the manufacturer Material Structure Water wettability Nominal pore size (m) Thickness (m)a Inner radius (m) Porosity (%) a
Polypropylene Asymmetric Hydrophobic membrane 0.2 400 900 60
Measured by SEM micrograph shown in Fig. 3A.
Fig. 4. Glass housing for removable membrane module: (A) glass housing; (B) O-ring; (C) metal ring; (D) union joint.
Fig. 4 shows the photo of the glass housing (A). The hydraulic seal between the module and the glass housing was provided by the O-ring (B), which was installed on the plastic tube outer surface and squeezed by the metal ring (C) by tightening a nipple in the female half of the union joint (D). An important feature of the glass housing was its reusable nature, since at the end of a given test the module could be simply removed from the housing by unscrewing the nipple connected to the union joint. Tests were performed using the laboratory-scale plant, whose scheme is shown in Fig. 5. A gas mixture (supplied by SIAD, Italy) containing 15% (v/v) of CO2 and 85% of N2, which is the typical composition of a flue gas from a coal combustion plant, was fed to the membrane lumen. The gas flow rate was carefully regulated by a fine metering valve (5) and measured at the inlet and outlet of the modules (G1 and G2) by soap bubble flow meters (F). A monoethanolamine (MEA) 3 M aqueous solution (V = 800 mL) was used as CO2 absorbent. Its temperature was
Fig. 3. SEM micrographs of Accurel S6/2 capillary membrane: (A) crosssection; (B) external surface; (C) internal surface.
Fig. 5. Scheme of the lab-scale plant used for CO2 absorption experiments. (1) Cryostatic bath; (2) recirculation pump; (3) absorption solution tank; (4) glass housing and membrane module; (5) fine metering gas valve; F: soap bubble flow meter; M: manometer; T: thermometer; L1: absorption solution inlet; L2: absorption solution outlet; G1: gas mixture inlet; G2: gas mixture outlet.
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kept at 20 ◦ C by a stainless steel coil immersed in the absorbent solution tank (3) and run through by water cooled by the bath (1). The liquid phase was fed on the shell side of the membrane module with a recirculation pump (2) at a flow rate of about 100 L/h. Its pressure was kept at a proper value, in order to avoid any possible penetration of the solution into the membrane pores as well as any gas bubble formation in the absorbent solution. The gas mixture flow rate was increased from 0.14 × 10−5 to 10.0 × 10−5 m3 /s. CO2 removal was estimated calculating the difference in CO2 concentration in the gas phase between membrane module inlet and outlet. The CO2 concentration was measured employing a gas chromatograph (mod. Autosystem, Perkin Elmer) equipped with a packed Porapak Q column and a TCD. 4. Results and discussion As Fig. 6 shows the percentage of CO2 at the outlet of each type of membrane module increases by increasing the gas flow rate, ranging from 0 up to 15% (v/v), i.e. the whole amount of CO2 in gas mixture. The rate of CO2 percentage increase depends on the number of capillaries: the higher the number of capillaries the lower the CO2 percentage. Fig. 7 shows the removal efficiency as a function of gas flow rate for two specific runs carried out with 1 and 10 capillaries modules, respectively. The behaviour shown in Fig. 7 is in qualitative agreement with the theoretical prediction depicted in Fig. 2. Obviously, it would be of great interest if we could predict gas removal efficiency from the knowledge of the system physical parameters. The parameter to be determined is the membrane overall mass transfer coefficient kM , which is represented by the ratio between the effective diffusion coefficient (Deff ) and the membrane thickness (s). This latter value (experimentally
Fig. 7. Experimental CO2 removal efficiency as a function of gas flow rate.
evaluated by a scanning electron micrograph of the membrane cross-section, shown in Fig. 3A) is reported in Table 1 along with the values of porosity (60%) and the nominal pore size (0.2 m) supplied by the manufacturer. In the SEM micrographs of Fig. 3B and C the contrast between the external and internal surfaces is striking. On the external surface pores some tens of nanometers large can be clearly observed, while on the internal one much smaller pores are present, whose size is often lower than the one (0.2 m) declared by the manufacturer. Inspection of the cross-section micrograph (Fig. 3A) also reveals a tortuous pore structure. On the basis of these experimental evidences and keeping into consideration the mean free path of CO2 molecules, it is possible to consider CO2 diffusion as mainly controlled by a combined bulk-Knudsen diffusion mechanism. Knudsen diffusion coefficient DK was determined by using the well-known correlation [24]: T DK = 48.5d (12) M where d is the mean pore size, T the operating temperature and M is the CO2 molecular mass. Calculations lead to a Knudsen diffusion coefficient of 25.2 × 10−6 m2 /s. A molecular diffusion coefficient taken from the literature [25] for CO2 bulk diffusion (DAB ), equal to 16.9 × 10−6 m2 /s was also considered. Both the coefficients were corrected for the effect of porosity, ε = 0.6 (Table 1), and tortuosity, τ, by using the relations: DK,eff =
εDK εDAB and DAB,eff = τ τ
(13)
where the tortuosity was calculated through the correlation: Fig. 6. Percentage of CO2 in gas samples taken at the membrane modules outlet, as a function of gas flow rate.
τ=
1 ε2
(14)
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References
Fig. 8. ln(1 − ηexp )] as a function of n/v0 compared with theoretical predictions.
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