7668
Ind. Eng. Chem. Res. 2005, 44, 7668-7675
CO2 Capture from Natural Gas Fired Power Plants by Using Membrane Technology May-Britt Ha1 gg* and Arne Lindbråthen Department of Chemical Engineering, Norwegian University of Science and Technology (NTNU), N-7491 Trondheim, Norway
Simulations of membrane separation were performed for two alternative processes for CO2 removal from an exhaust gas stream at a natural gas fired power plant. One process considered is an integrated membrane separation with pressurized gas (16.5 bar) rich in CO2 (10 vol %), while the other alternative solution is a process for CO2 capture from tail-end flue gas where the content of CO2 is low (4 vol %) and gas is released at 1 bar. A gas stream of 700 000 Nm3/h is considered for both cases. Flux data obtained for a fixed-site-carrier poly(vinylamine) membrane developed within the research group at NTNU were used. This membrane has moderate CO2 flux but very high selectivity in favor of CO2. A favorable process design with a single-stage unit without the extra need for compression, and also use of a sweep gas to meet the process specifications, was recommended for the pressurized gas. For the flue gas lean in CO2, the driving forces are too low for separation unless both a compression of the feed gas (f4 bar) and a vacuum pump on the permeate side are used. Also, in this case a small fraction of the sweep gas on the permeate side is recommended. Introduction To meet the challenges of manmade climate changes, the Kyoto protocol under the United Nations Framework Convention on Climate Changes (UNFCCC1) drafts the necessary reduction in the emissions of greenhouse gases that has to take place worldwide. The targets for reduced emissions are set on different levels for various groups of countries, and the protocol would enter into force when a sufficient number of countries producing around 50% of the manmade emissions had signed. This has now been done, and the Kyoto protocol will enter into force in February 2005. It states that within years 2008-2012 the emissions should be reduced by 5.2% of those in 1990. This is a very ambitious target that currently seems to be difficult to reach for many countries, including Norway among others. The production at the oil and gas facilities offshore, as well as the planning of natural gas fired power plants, has lead to national commitments for research and development to focus on efficient methods for CO2 capture and storage of CO2 in geological formations or use for enhanced oil recovery. Today, Norway basically relies on hydroelectric power (which is an environmentally friendly but a limited energy resource). The capacity of storage in identified aquifers off the coast is judged to be almost unlimited for the time scale expected of using fossil fuels as the energy source (150 years?), and CO2 from all of Europe can be received. Several concepts are being proposed by the industry for CO2 capture, and in the current paper, the focus is on the potential of using membrane separation for two different cases: (1) by CO2 capture as an end of process solution from the flue gas of a natural gas fired power plant and (2) by CO2 capture from pressurized gas by placing the membrane separation unit between the compressor and the turbine. * To whom correspondence should be addressed. Tel.: +47 73594033. Fax: +47 73594080. E-mail: may-britt.hagg@ chemeng.ntnu.no.
If CO2 is going to be reinjected into the reservoir or aquifer, there will be specifications for pressure and composition that have to be met for the gas, also depending on how it is to be transported to the point of injection. On the basis of available information,2 the following criteria were set for the membrane separation: (i) The CO2 capture should be at least 80% to be able to compete with similar concepts. (ii) The O2 concentration in the permeate should preferably not be higher than 1%. (iii) The pressure drop on the feed side of the membrane, loss of mass (other than CO2) over the membrane, and use of energy-consuming units must be held at a minimum to avoid large reductions in the plant’s performance and efficiency. (iv) Minimizing the membrane area will result in lower investment costs and a smaller footprint. In the current paper, the membrane process has been evaluated by optimized simulations with respect to the criteria above. The results are presented in the current paper. Description of Two Process Solutions for a Natural Gas Fired Power Plant There are different technologies to produce power based on natural gas. The technology on which the process discussed in the paper is based is a combined gas turbine-steam turbine process. This process will typically be used in countries where natural gas is easily available. In brief, the gas turbine consists of a compressor and a combustion chamber in addition to the turbine. The compressor will, under controlled conditions, suck air from the surroundings and compress this to a pressure in the range of 10-30 bar (depending on the type of turbine). In the combustion chamber, the combustion of fuel (CnHm) and air is strictly controlled in order to optimize the energy output and minimize the formation of pollutants, especially the amount of
10.1021/ie050174v CCC: $30.25 © 2005 American Chemical Society Published on Web 07/08/2005
Ind. Eng. Chem. Res., Vol. 44, No. 20, 2005 7669
Figure 1. Integrated CO2-capture process as suggested by Statoil (the Combicap cycle).4
NOx. The temperature of the flue gas out of the combustion chamber can be up to 1500 °C, but this may be restricted depending on the materials used and the cooling system. In a gas turbine, there will be almost complete oxidation of the fuel, so the chances of forming CO or having residues of hydrocarbons are very small. The main components in the flue gas from the gas turbine are CO2, H2O, N2, and excess O2. However, because the fuel is burned in air rather than in pure oxygen, the nitrogen in the air may participate in the combustion process to produce nitrogen oxides (NOx; the amount will depend on the system). In addition, argon (Ar) from the air will be present.3 The concentration of SO2 is minimal in the natural gas considered here and is therefore neglected in our separation process. The exhaust gas from a gas turbine will typically have a temperature of around 450-650 °C and has cooled to about 80-100 °C when it is released at approximately atmospheric pressure. The general reaction for the complete oxidation can be written as
(
C nH m + n +
m m O f nCO2 + H2O 4 2 2
)
(1)
From the brief description of the process given above, it can be understood that removal of CO2 can be done at two points in the system: either after the combustion chamber, while still at high pressure, or after the turbine at atmospheric pressure. The state-of-the-art solution today would be using a tail-end amine absorption plant. The major challenge with this concept is the relatively low CO2 partial pressure, which will make such a plant large in size and the CO2 removal expensive. Such plants are well-known for high-pressure natural gas processing and are then substantially smaller. The low partial pressure is also a challenge for tail-end membrane separation. The possibilities of alternatively using membrane separation for the capture of CO2 in the two possible processes are being discussed in the current paper.
Table 1. Permeances [m3(STP)/m2‚bar‚h] and Selectivities for Gases Used in the Simulations CO2
O2 10-3
N2 10-4
H2O
0.107 2 × 8× 1.067 0.107 3 × 10-4 10-4 0.140 3 × 10-4 10-4
CO2/O2 CO2/N2 membrane 51 356 466
134 >1000 .1000
A B C
Table 2. Typical Compositions of Exhaust Gas and Process Conditions2,3 process
XCO2 XO2
XN2 XH2O P [bar] T [°C] comment
pressurized gas 0.11 0.02 0.85 0.02 tail-end gas 0.04 0.13 0.76 0.07
16.5 1
50 90
ref 2 ref 3
(I) CO2 Capture from Pressurized Exhaust Gas; an Integrated Membrane Process. There are many ways of obtaining increased CO2 partial pressure in a gas turbine based power plant. Some concepts involve the connection of the CO2-capture plant to the highpressure part of the standard gas turbine cycle; other concepts involve partial recirculation of the gas turbine exhaust.3 A concept developed by Statoil, Stavanger, Norway, represents a new approach for an integrated process because it is aimed at providing a high partial pressure in order to use cost-effective CO2-capture technology. The concept involves a semiclosed gas turbine cycle where the separation plant is located after the compressor part of the gas turbine. The concept is called “Combicap” and is illustrated in Figure 1.4 In the figure, a CO2-capture plant is indicated as an integrated unit in the process. Removal of CO2 by an integrated membrane separation process at this point is evaluated in the current paper. Although not shown in the figure, a sweep gas may be introduced on the permeate side of the membrane for enhanced CO2 capture. Five different cases of membrane separation were simulated; the results are reported in Table 3 and discussed in paragraph 5. (II) Tail-End CO2 Capture; Membrane Purification of the Flue Gas. The standard gas turbine cycle is briefly described below. A schematic illustration of the cycle is shown in Figure 2. The exhaust gas from the cycle leaves the process where the flue gas is
7670
Ind. Eng. Chem. Res., Vol. 44, No. 20, 2005
Table 3. Simulations of the Pressurized CO2-Capture Processa case/ membrane
power consumption [MW]
sweep gas [vol % of the feed]
XCO2 in retentate
YO2 in permeate
CO2 removed [vol %]
membrane area [m2]; module II
total membrane area [m2]
1/A 2/B 3/B 4/C 5/C
12 11 9 0 0
10 no sweep 5 5 5
0.02 0.013 0.007 0.013 0.014
0.005 0.02 0.02 0.017 0.018
83.1 89 94.3 88.8 84.4
7.1 × 104 106 106 0 0
17 × 105 17 × 106 3 × 106 106 0.8 × 106
a
Gas volume ) 700 000 Nm3/h. Feed pressure ) 16.5 bar. Temperature ) 35 °C. CO2 in the Feed ) 11 vol %.
where D0 and S0 are temperature-independent constants, Ed is the activation energy for diffusion [J/mol], ∆Hs is the heat of solution [J/mol], R is the gas constant, and T is the temperature [K]. Depending on the type of membrane material and type of gas component, the relative effect of the temperature will vary; details about this may be found in the literature on transport mechanisms.5,6 The intrinsic separation factor, R, of a membrane is defined as the ratio between the pure gas permeabilities, P
R ) PA/PB Figure 2. Standard gas turbine-steam turbine cycle.
indicated. This is where the tail-end CO2-capture plant will be installed. It may be installed by using amine absorption or, with some modifications, by using membrane separation as described in the current paper. The results from six different cases are illustrated in Tables 4 and 5 and evaluated in the Discussion section. CO2 Capture by Membrane Separation General Procedures. The most attractive feature of a membrane separation process is the simplicity of the process; there is no need for the addition of chemicals or for the regeneration of any absorbent. The governing flux equation for the gas permeation (eq 2) is based on Fick’s law, and as can be seen, the driving force is then the difference in partial pressures over the membrane. In a simple way, the flux, J [m3(STP)/m2‚ h], may be expressed as
PA qpyA ) JA ) (phxA - plyA) Am l
(2)
where qp ) volume of the permeating gas [m3(STP)/h], PA ) permeability of gas component A [m3(STP)m/(m2‚ h‚bar)], l ) thickness of the membrane [m], ph and pl ) pressure on the feed and permeate sides [bar], xA and yA ) fractions of component A on the feed and permeate sides, respectively, and Am [m2] is the required membrane permeation area. The permeability, P, is also expressed as the product of diffusion, D [m2/s], and solubility, S [m3(STP)/m3‚bar], coefficients for the gas in the membrane material.
P ) DS
(3)
Both coefficients are temperature-dependent and may be expressed by Arrhenius types of equations (eqs 4 and 5). In general, the diffusion will increase with temperature and the solubility will decrease.
D ) D0 exp(-Ed/RT)
(4)
(6)
while the selectivity, R, between two components, A and B, in a mixed gas stream may be expressed as in eq 7 (where y and x are the compositions of A and B on the permeate and feed sides, respectively):
R)
yA/yB xA/xB
(7)
Different gas transport mechanisms may govern the membrane separation. Today, basically all commercial membranes for gas separation are based on the solutiondiffusion mechanism through a dense polymeric membrane (see eq 3). As can be seen from eq 2, the flux will greatly depend on the thickness of the membrane and the difference in partial pressures. The required membrane area for permeation is inversely proportional to the flux and directly proportional to the volume of the gas; hence, in order to be an economically viable choice in combination with the power plant, the driving force, the flux, and the selectivity of CO2 need to be enhanced by optimizing these factors. The current paper describes a membrane separation where the membrane itself is a polymer with a “carrier component” introduced into the polymer structure itself. This type of membrane is referred to as a “fixed-sitecarrier” (FSC) membrane. In the current paper, the material is a FSC poly(vinylamine) (PVAm) membrane, where the amine group will contribute to the transport of CO2 through the membrane as a bicarbonate ion (HCO3-) when the membrane is wet (swollen with water). We then have a carrier effect in addition to the Fickian diffusion, as illustrated in eq 8. Standard polymeric membranes without carriers will loose their separation efficiency when swollen, while for this specific FSC PVAm membrane, these properties will be enhanced. Details about this membrane may be found in the paper published by Kim et al.7
JA )
DA DAC (c - cA,l) + (c - cAC,l) l A,0 l AC,0 cA ) SApA
(8) (9)
where the first term on the right-hand side of eq 8 is
Ind. Eng. Chem. Res., Vol. 44, No. 20, 2005 7671 Table 4. Simulations of “Tail-End CO2 Capture”a case; based on membrane C 6 7 8
pressure ratio, ph/pl
sweep gas [vol % of the feed]
XCO2 in retentate
YO2 in permeate
CO2 removed [vol %]
membrane area [m2]
power consumption [MW]b
4/0.10 4/0.10 1.5/0.2
no sweep 5 10
0.019 0.002 0.006
0.198 0.013 0.038
54.5 95.2 87
35 × 106 5 × 106 25 × 106
52 + 2 52 + 9 13 + 4
a Gas volume ) 700 000 Nm3/h. Feed and permeate pressures vary. Temperature ) 35 °C. CO in the Feed ) 4 vol %. b The sum 2 indicates a compressor for feed + vacuum pump for the permeate.
Table 5. Simulations of “Tail-End” CO2 Capturea case; based on membrane C
pressure ratio, ph/pl
sweep gas [vol % of the feed]
XCO2 in retentate
YO2 in permeate
CO2 removed [vol %]
membrane area [m2]
power consumption [MW]b
9 10 11
4/0.10 4/0.10 1.5/0.2
no sweep no sweep 10% of the feed
0.008 0.011 0.022
0.07 0.010 0.003
80.3 90 80
5 × 106 7 × 106 107
52 + 8 52 + 8 13 + 7
a
Temperature ) 50 °C. CO2 in the feed ) 10 vol %. b The sum indicates a compressor for feed + vacuum pump for the permeate.
S ) S0 exp(-∆Hs/RT)
(5)
the Fickian diffusion (DA) and the second term represents the carrier-mediated diffusion (DAC). l is the thickness of the membrane, while c is the concentration of component A and its complex AC at the interfaces of the membrane 0 and l on the feed and permeate sides, respectively. The concentration difference of the complex AC in eq 8 must be further expressed by an equilibrium constant of the complexing reaction and a distribution coefficient. This is explained in detail by Cussler.8 Referring to eqs 2 and 8, it can be understood that the temperature and pressure of the CO2-capture process have a major impact on the membrane separation. The most favorable process conditions for using polymeric membranes are, in general, at low temperature ( 7. The design of the membrane module will also be important, both for maximizing the driving force and for reducing the required number of modules needed for separation. The process may also be optimized by various cascade solutions. This is more closely looked at in the Results and Discussion sections. Membrane Separation Process. The input data to the membrane simulations are based on measured gas permeation values at 35 °C and pressures in the 2-5 bar range. Initial measurements indicate that temperature variations in the range of 35-50 °C are negligible. The driving force, defined as the difference in the partial pressures over the membrane, is, however, important when the Fickian diffusion is considered (eq 2) and will have a large impact on the separation. If a ∆T over the membrane could be provided, an enhanced carrier effect would be expected. This is, however, not taken into account here; the carrier effect is only documented for equal temperature on both sides of the membrane. Two different processes were considered (described above) using the FSC PVAm membrane referred to: an integrated membrane process for pressurized gas and tail-end capture of CO2. The membrane process simulations were carried out using the input data given in Tables 1 and 2. Table 1 reflects how the separation properties of the FSC membrane have been improved during the development. The membranes are identified as membranes A-C. The results obtained with membrane A were carried out before the effect of water on the membrane was fully investigated. In the simulations
performed with data for membrane B and C, the transport of water through the membrane is neglected because of the fact that the membrane itself is kept continuously wet in the experimental setup and gases on both sides are saturated with water. The very high selectivities obtained in favor of CO2 are believed to be caused by the controlled cross-linking of the membrane and the effect of the fluoride ion in the membrane structure (cross-linking agent). The fluoride enhances the transport of the polar CO2 molecule in addition to the effect introduced by the amine carrier; details about this are found in ref 6. The thickness of the membrane was assumed to be 2 µm in the current simulations, which is set as a realistic goal for the membrane development. The volume of the feed gas to the membrane is set to be equal in both cases: ∼700 000 Nm3/h. This compares approximately to an exhaust gas stream from a 200MW power plant. The composition of the feed gas is varied for the two different processes as well as process conditions; see Table 2. With reference to Table 2 and eq 2, it can easily be seen that it will be difficult to obtain a driving force over the membrane for tail-end removal of CO2 unless (1) the exhaust gas is compressed before it is fed to the membrane, (2) the sweep gas is introduced on the permeate side in order to reduce the partial pressure of CO2, and/or (3) a vacuum pump is used on the permeate side. For tail-end removal, the concentration of CO2 is usually quite low; hence, the volume of the permeating CO2 may be acceptable for the use of a vacuum pump. On the basis of these general considerations, the simulations were performed for the two different processes described. Simulation Program The simulations were performed with an in-house membrane program interfaced to the process simulation program HYSYS; hence, it has the possibility of utilizing HYSYS’s capacity to calculate fugacities and also of incorporating energy balance into the models. The simulations were based on the Soave-Redlich-Kwong equations of state. The code is developed for cross-flow and cocurrent and countercurrent flow and with the possibility of introducing a sweep gas on the permeate side. This means that the user has the possibility of choosing a spiral-wound module or hollow-fiber module with or without a sweep gas. Typical input data are the feed composition, permeances, and process conditions
7672
Ind. Eng. Chem. Res., Vol. 44, No. 20, 2005
Figure 3. Illustration of a membrane countercurrent-flow model.9
such as pressure and temperature, and if the needed membrane area is assumed, the program will calculate the resulting retentate and permeate streams with volumes and compositions as well as give the “stage cut” (θ). The stage cut is defined as
θ ) qp/qf
(10)
Also, power consumption may be calculated. For the separation in the current paper, a countercurrent hollow-fiber model with the possibility of a sweep was chosen. The flow is then treated as plug flow on both the feed and permeate sides, basic equations are combined with the overall mass balances, and equations are derived over the volume element, as illustrated in Figure 3. It has been concluded by many parametric studies that, at the same operating conditions, the countercurrent-flow pattern yields the best separation and requires the lowest membrane area.9 Results Pressurized Gas Cycle. The point for an integrated membrane separation is indicated in Figure 1. The CO2lean exhaust gas may after treatment be recycled back into the process (heat exchanger and combustor), while the permeate, the CO2-rich stream, may go to reinjection into the reservoir and increased oil recovery (IOR). Five alternative cases were simulated. All results from these simulations are presented in Table 3 and commented on below. Figures 4 and 5 illustrate the recommended solutions. The adiabatic efficiency, η, of the compressor (when needed) was set to 0.75. For some of the simulated cases, a sweep gas is introduced in order to increase the partial pressure difference driving force. The sweep gas is available on site as steam, and the amount of sweep gas is given in the table as vol % of the feed. Case 1 (Results Based on Data for Membrane A; Table 1). From previous work,10 it was shown that the most efficient membrane process solution for CO2 capture from a pressurized gas stream was a cascade solution with recovery of CO2 in two steps in order to meet the criteria on both the O2 concentration in the permeate stream (80%). The recommended solution is illustrated in Figure 4. With reference to Table 1, it should be noted that this membrane was at that time not optimized with respect to flux (for CO2) and selectivity. Cases 2 and 3 (Results Based on Data for Membrane B; Table 1). The selectivity with respect to CO2 is now increased, but the flux is the same as that for data 1 and the process solution as shown in Figure 4. As can be seen, the goals for CO2 capture and the O2 concentration in the permeate can easily be reached both with and without using a sweep gas, but there is a major difference in the required membrane area when a sweep gas is introduced.
Cases 4 and 5 (Results Based on Data for Membrane C; Table 1). The CO2 flux has now been increased further, and the selectivity is very high. A second membrane unit with additional treatment of part of the permeate from the first stage is no longer necessary, and a single stage can be used. However, to increase the driving force, a sweep gas should be introduced on the permeate side at 1 bar. This is illustrated in Figure 5. The water vapor (sweep) may afterward be removed by a condenser as indicated. A “single stage” may be understood as several membrane modules in parallel. A two-stage cascade could also be pictured using a sweep gas only at the second stage. This was, however, not simulated. Tail-End CO2 Capture. With reference to Table 2, where a standard composition of an exhaust gas stream from a combined gas turbine-steam turbine process is shown, and given the general considerations in the previous section, it can be concluded that, in order to recover CO2 from the tail gas, compressors and/or vacuum pumps need be installed in order to increase the driving force (difference in CO2 partial pressures) for the membrane separation. A compressor may, however, in most cases be judged to be too expensive. The best solution is, therefore, believed to be separation by pulling a vacuum on the permeate side and using modules in series, as illustrated in Figure 6. Depending on the fraction of CO2 in the feed stream and the volume of gas to be removed (volume of the permeate), a small volume of sweep gas (water vapor) should be used. Six different cases were simulated, three with a CO2 fraction in the feed equal to 4 vol % (Table 4) and three with 10 vol % (Table 5), to illustrate possible solutions. The volume of the gas feed to the membrane is also simulated here as 700 000 Nm3/h; the composition for the tail-end gas is given in Table 2. The composition of the tail-end gas may vary depending on the type of power plant; hence, the fraction of CO2 in the gas stream may easily be much higher. Simulations were, therefore, also performed to illustrate how this may affect the membrane separation. The same pressure ratios and sweep/no sweep as those illustrated in Table 4 were used, but now for the composition XCO2 ) 0.10, XO2 ) 0.07, XN2 ) 0.75, and XH2O ) 0.08. Discussion General Procedures. When processes for CO2 capture are to be evaluated, it may be important to bring into the discussion why it is at all considered to recover the CO2 from flue gas streams. Because international agreements such as the Kyoto protocol restrict the emissions of greenhouse gases, it brings about the need for new technology development to bring down the costs for CO2-capture plants. An important question is further what to do with the captured CO2. Depending on where the power plant is situated, it has to be transported either by tankers or in pipelines for reinjection into reservoirs or aquifers. For some places in the world, dumping on large ocean depth or depositing the CO2 as carbonates is considered. The most attractive solution however, at least in Europe, seems to be reinjection into reservoirs or aquifers. This means that the permeate stream will need to be recompressed in order to be transported in bulk or by pipeline and to meet certain specifications. The pressure should be around 60 bar or even higher
Ind. Eng. Chem. Res., Vol. 44, No. 20, 2005 7673
Figure 4. Recommended cascade solutions for the low-selectivity membrane.9
Figure 7. Effect of the separation factor and pressure ratio on the permeate purity11 (feed xf ) 0.3). Figure 5. Recommended single-stage solution for the highselectivity membrane.
Figure 6. Membrane modules in series.
(dense phase). It will also be important to remove water from the gas in order to avoid the formation of hydrates and corrosion. Depending then on how to dispose of the CO2, these aspects may influence the choice of technology for CO2 capture. Effect of Process Variables on the Separation and Area. For the FSC membrane considered, a very high selectivity in favor of CO2 but only moderate CO2 flux [∼0.14 m3(STP)/m2‚h‚bar] has been documented (>1000). This means that the stage cut, θ, in general will be low and a very pure permeate will be obtained (which is good), but large permeation areas may be expected in order to recover the amount of CO2 needed from these large feed streams (>80%). As can be understood from eq 2, there will also be a problem with the driving force over the membrane for the Fickian diffusion as more and more CO2 is being removed. Although this will be less dramatic when countercurrent
flow is being used, it can be understood that, for gas streams with low fractions of CO2, the use of a sweep gas is needed to increase the partial pressure difference over the membrane. Increased temperature will have a positive effect on the carrier transport but may, on the other hand, be negative for the transport by solution diffusion of CO2 (refer to eq 8). Recent results indicate, however, that the net temperature effect on CO2 transport through the membrane is positive in the pressure range investigated. When water vapor is introduced as a sweep gas, the condensation temperature of the resulting permeate stream should be checked; this is automatically calculated by the simulation program. A general relationship between the pressure ratio, ph/ pl, separation factor, R, and purity of the permeate for a given feed may be considered. This has been done by Stookey et al.11 for a complete mixing model and with xf ) 0.3 in the feed gas (Figure 7). From this diagram, which shows a general trend, it can be seen that only for low separation factors the purity is greatly affected, while a high-pressure ratio always will be more favorable. For the separation illustrated in the figure, it can be seen that, for selectivities higher than 20, the purity is not greatly affected. It will, therefore, be useful to look at how this fact may affect our separation. It may be useful to introduce the term “separation power”, SP, which is understood as the flux times selectivity (for the most permeable component, eq 11):
SP ) JiR
(11)
Currently, the separation power with respect to our membrane C is ∼140 m3(STP)/m2‚h‚bar (see Table 1). This SP could then also be achieved by accepting a selectivity of ∼500 and aiming for a doubled permeance for CO2 in the same membrane at the same tempera-
7674
Ind. Eng. Chem. Res., Vol. 44, No. 20, 2005
ture-pressure conditions (0.14 × 2 ≈ 0.3 m3(STP)/m2‚ h‚bar); hence, the permeation area needed will be reduced by half. For this special membrane, this would mean that cross-linking of the selective layer will have to be adjusted (see ref 7). This is considered to be achievable and the economics of the membrane separation, hence, more attractive. A sample membrane process using a “high-performance CO2 membrane” was calculated by Baker.12 These results may be compared with the current highselectivity FSC PVAm membrane (membrane C, cases 4 and 5). In his example, natural gas at high feed pressure (∼70 bar) with the same amount of CO2 (10 vol %) is to be separated. The same CO2 permeance as that aimed for here [0.3 m3(STP)/m2‚h‚bar] and a selectivity for CO2/CH4 ) 40 are being used. Updating the numbers in this reference12 to the same volume gas stream as here (700 000 Nm3/h) would require a permeation area of ∼107 m2 and a power consumption of 16 MW for a configuration with recompression of the permeate and some recirculation). The current FSC membrane for a pressurized flue gas at 10 bar and a permeance of 0.14 m3(STP)/m2‚h‚bar requires ∼106 m2 and no compression power (Table 3); the doubled permeance (0.3) would require only half the permeation area (0.5 × 106 m2). It should be pointed out that CH4 and N2 are usually very much alike in the way they permeate and are therefore difficult to separate; this makes the comparison here interesting, although the feed gases are very different. The results confirm the promising potential of the current FSC PVAm membrane. Discussion on the Integrated Process for a Pressurized Gas. Referring to Table 3 for the results obtained for the pressurized gas with a high CO2 content (11 vol %), it can be stated that the most favorable is case 5. No extra compressor is needed, but a sweep gas is needed in order to reach the specifications. The process was also simulated without a sweep gas, but it was then impossible to meet the criteria for CO2 recovery. Using a combination of two stages, as illustrated in Figure 6, with a sweep gas only on the second stage, could possibly be a solution, but this would mean an increased membrane area. Discussion on the Tail-End Recovery of CO2. For recovery of CO2 from tail-end exhaust gas, Table 4 clearly shows that the only possibility for meeting the specifications here is a combination of a vacuum pump and a sweep gas. The power consumption is the sum of both pressurizing the gas (to 4 or 1.5 bar) and using a vacuum pump for the permeate (0.1 or 0.2 bar). For these large gas streams, the required membrane area is huge, and a membrane process is probably not a viable option for CO2-lean gas streams (4 vol %). The low CO2 fraction is also a problem for alternative processes. However, the tail-end gas from power plants using other types of fuel (coke, heavy oil, etc.) may have a higher fraction of CO2, and in Table 5, simulations are performed with a composition assumed to be more like this situation and with the pressure ratios used for the CO2-lean flue gas (10 vol %). As can be seen, now the criteria can be met also without using a sweep gas, and the permeation areas are not dramatically high. In Figure 8, the dependency of CO2 recovery on the
Figure 8. Capture of CO2 (%) as a function of the membrane area for ph/pl ) 1.5/0.2 and the fraction of CO2 in the feed, x ) 0.10.
membrane area for pressure ratio ph/pl ) 1.5/0.2 is shown for illustration. Calculation of the Number of Modules. In the current simulations, the membranes were assumed to be hollow fibers. The hollow-fiber module has a very high packing density; up to 30 000 m2/m3 is stated.6 The diameter and length of the fibers as well as the inner diameter of the module can be given as typical data from the vendor; hence, the number of modules may easily be calculated for the required permeation area. If a packing density of 30 000 m2/m3 is assumed and the inner diameter of the module is set to 0.4 m and the inner length to 5 m, the volume of the module is calculated to be 0.63 m3. Hence, the permeation area in each module will be 18 840 m2/module, and the number of modules may be calculated. The optimal size module may vary among different vendors. For the best case of a pressurized gas (case 5), the number of modules will thus be ∼90, while for tail-end removal, case 8 (4 vol % CO2 in flue gas) has 1327 modules and for case 11 (10 vol % CO2 in flue gas) 530 modules. If the flux for the FSC membrane can be brought up to 0.3 as indicated in the discussion on process variables, the number of modules will be reduced by half. Summary Membrane separation may in the future be a viable option for CO2 capture from exhaust gas from a natural gas fired power plant. The best choice would then, however, be to use an integrated solution for a pressurized gas where the CO2 fraction has been increased; here a fraction of ∼10 vol % has been assumed. The driving forces over the membrane may further be increased by using water vapor as the sweep gas on the permeate side. To reduce the permeation area for these large gas streams, a CO2 flux of ∼0.3 m3(STP)/m2‚h‚bar should be aimed for. Membrane separation is not recommended for tail-end recovery unless a sweep gas or a vacuum pump is used on the permeate side. Acknowledgment The authors want to acknowledge Ph.D. student David Grainger for his quality work on the code for the membrane separation model and Statoil for supplying information on the Combicap process. Literature Cited (1) http://unfccc.int/essential_background/kyoto_protocol/items/ 2830.php.
Ind. Eng. Chem. Res., Vol. 44, No. 20, 2005 7675 (2) Knoph, I. Simulation of an integrated membrane process for CO2-capture at a natural gas fired power plant. Thesis (in Norwegian) IKP, Norwegian University of Science and Technology (NTNU), Trondheim, Norway, 2004. (3) Bolland, O. Natural gas fired power plants (in Norwegian). Energy in Norway A5171; Sintef: Trondheim, Norway, 2000; Chapter 2.5. (4) Lynghjem, A.; Jakobsen, J.; Kobro, H.; Lund, A.; Gjerset, M. The Combicap cyclesefficient combined cycle power plants with CO2 capture. The 2nd Trondheim Conference on CO2 Capture, Transport and Storage, Trondheim, Norway, Oct 2004. (5) Vieth, W. R. Diffusion in and through polymers; principles and applications; Hanser and Oxford University Press: Mu¨nchen, Germany, and Oxford, U.K., 1991. (6) Mulder, M. Basic principles of membrane technology, 2nd ed.; Kluwer Academic Publishers: Dordrecht, The Netherlands, 1996. (7) Kim, T.-J.; Li, B.; Ha¨gg, M.-B. Novel fixed-site-carrier polyvinylamine membrane for CO2 capture. J. Polym. Sci., Part B: Polym. Phys. 2004, 42, 426-4336.
(8) Cussler, E. El. Facilitated and active transport. In Polymeric gas separation membranes; Paul, D. R., Yampol’skii, Y. P., Eds.; CRC Press: Boca Raton, FL, 1993; Chapter 6. (9) Geankoplis, J. C. Transport processes and separation process principles, 4th ed.; Prentice Hall: Englewood Cliffs, NJ, 2003; Chapter 13. (10) Knoph, I.; Lindbråthen, A.; Ha¨gg, M.-B. Natural gas fired power plant with CO2 capture. GPA Conference, Dublin, Ireland, May 2004. (11) Stookey, D. J.; Patton, C. J.; Malcolm, G. L. Membranes separate gases selectively. Chem. Eng. Prog. 1986, 82 (11), 36. (12) Baker, R. W. Future directions of membrane gas separation technology. Ind. Eng. Chem. Res. 2002, 41, 1393-1411.
Received for review February 14, 2005 Revised manuscript received June 1, 2005 Accepted June 6, 2005 IE050174V
