Post-combustion Carbon Capture with a Gas Separation Membrane

Article pubs.acs.org/EF

Post-combustion Carbon Capture with a Gas Separation Membrane: Parametric Study, Capture Cost, and Exergy Analysis Xiangping Zhang,*,†,‡ Xuezhong He,§ and Truls Gundersen† †

Department of Energy and Process Engineering, and §Department of Chemical Engineering, Norwegian University of Science and Technology, NO-7491 Trondheim, Norway ‡ Beijing Key Laboratory of Ionic Liquids Clean Process, State Key Laboratory of Multiphase Complex System, Institute of Process Engineering, Chinese Academy of Sciences, Beijing 100190, People’s Republic of China ABSTRACT: A systematic method that integrates process simulation, capture cost estimation, and exergy analysis is applied to evaluate a gas separation membrane process for post-combustion carbon capture in a coal power plant. The influences of membrane performance and process configuration on the energy consumption, required membrane area, and capture cost have been studied. The results indicate that the energy consumption decreases with the increase of CO2/N2 selectivity, but a larger membrane area is required, while for a high CO2 permeance, the membrane area can be significantly reduced. The carbon capture ratio influences the specific energy consumption as well, which should be a trade off. For a two-stage membrane process, the capture load distribution between the first and second stage affects the separation performance greatly, and the optimal range varies with the other parameters. Under the assumptions in this work, the profiles of capture cost related to membrane parameters show that the optimal CO2/N2 selectivity is 70−90. The exergy analysis indicates that the main energy bottleneck of a membrane technology is located in the membrane unit operation, which has relatively low exergy efficiency. On the other hand, CO2 compression has less potential for energy savings because it has already had very high exergy efficiency. Department of Energy’s Carbon Sequestration Program.3 In fact, membranes for gas separation have emerged from a technical curiosity in the 1960s to initial commercialization in the 1970s, followed by intensive research activities and further commercialization in the 1980s.9 Many patents have been awarded in membranes for CO2 removal from natural gas and H2 purification in reforming processes (UOP and Air Products and Chemicals),5 which are the foundation of carbon capture from flue gases.10 The advantages of the gas separation membrane process can be summarized as less environmental impacts and small footprint, no steam load, ease of upscaling and operation, and ease of use both in grassroot power plants and retrofitting existing power plants.4,5 Moreover, membranes are also a low-cost means for gas separation, especially when high-purity gas streams are not required.11 In a membrane gas separation process, the membranes act as a filter to separate one or more gas components from a feed gas mixture and produce a specific gas-rich permeate and retentate stream. In a cross-flow filtration, a feed stream at high pressure can pass through a membrane to the permeate side at low pressure based on a positive pressure difference (i.e., driving force). Two parameters, permeability (permeance = permeability/thickness) and selectivity are typically used to characterize the membrane performance. Permeance is defined as the flux of a specific gas passing through a membrane at a given pressure and temperature, and selectivity is evaluated by the

1. INTRODUCTION Ensuring reliable, affordable energy while reducing carbon dioxide (CO2) emissions is a critical challenge, and carbon capture and storage (CCS) is likely to be an essential part of the solution. Amine-based technologies, such as the monoethanolamine (MEA) solvent system, have been proven with the highest feasibility in the commercialization of postcombustion carbon capture among the various capture technologies. Thus, most of the pilot plants established in the world now mainly use amine-based or advanced amine-based systems,1,2 and a few of them have been improved by Fluor, Mitsubishi Heavy Industries, and Cansolv.3 The main goals of the technological improvement are to reduce energy and solvent consumptions and eventually reduce the operation cost and investment. However, the inherent weaknesses of aminebased chemical absorption methods still hinder their economic applications, for example, solvent degradation, leading to high material costs and high disposal costs; operating limitations, such as the interdependence of the two fluid phases to be contacted, which may produce emulsions, foaming, unloading, and flooding; additional environmental pollution caused by the solvent emission and degradation; high energy consumption, i.e., the low-pressure steam from steam turbines, which would be employed for regenerating the rich solvent; and auxiliary retrofitting measures or costs in power plants, such as modification of the steam turbines.4−6 To overcome the disadvantages of amine-based systems, some new technologies, such as solid absorbents, membranes, and ionic liquids, have been studied and developed.3,7,8 Among them, the gas separation membrane is regarded as one of the most promising capture technologies, which might create cost reduction benefits in the near future, as reported by the U.S. © 2013 American Chemical Society

Special Issue: Accelerating Fossil Energy Technology Development through Integrated Computation and Experiment Received: December 28, 2012 Revised: February 28, 2013 Published: March 4, 2013 4137

dx.doi.org/10.1021/ef3021798 | Energy Fuels 2013, 27, 4137−4149

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Figure 1. Framework of this study.

potential research directions to accelerate its industrial and commercial applications. Most studies have been focused on the synthesis membrane materials, characterization, and development in the last 2 decades,10,11,26−28 while only a few studies on the systematic analysis of the membrane system have been reported.4,18,21−23,29,30 In 1992, Van Der Sluijs et al. evaluated the feasibility of polymer membranes for CO2 recovery from flue gas in a power plant, and several parameters were optimized to obtain the lowest specific CO2 mitigation costs.29 Bounaceur et al. conducted a systematic analysis of the separation performance and energy consumption for a singlestage membrane unit and optimized pressure ratio.21 Zhao et al. described a detailed parametric study of the mass and energy balance for a single-stage membrane22 and also the energetic and economic analyses of multi-stage membrane processes.4 Merkel et al. investigated the influences of membrane parameters and process configurations on the energy consumption and capture cost considering a real industrial process, and some important suggestions were proposed for the future directions.23 Hussain et al. investigated the feasibility of carbon capture from flue gases by facilitated transport membrane based on process simulation and economic cost estimations for the real flue gas.18 These works are important for understanding the feasibility of carbon capture with membrane technology; however, some results were inconsistent, and some were far from a real process. Thus, more systematic research should be conducted. Additionally, only a few studies focused on this energyintensive process with the thermodynamic principle. In 1991, Britan et al. investigated the feasibility of membrane technology for N2/H2 separation using exergic efficiency as the main criterion.31 Their results showed that a very high pressure ratio (661) and a high selectivity (261) are required to achieve the product with 90 and 95% H2 recoveries, respectively, from a feed gas containing 40% H2 using a single-stage membrane process, which indicated, in theory, that the process consumed a large amount of compression power, even though such a high pressure ratio is unrealistic.31 Merkel et al. pointed out that power consumption is a key issue of a membrane system for post-combustion carbon capture based on the comparison of the power consumptions in different modules.23 Thus, understanding the energy consumption status of the whole process is a key point for process design. Exergy analysis can be a powerful tool for energy analysis, which reflects the losses of potential work resulting from irreversibilities because of the second law of thermodynamics.32 It will be helpful to find where the exergy losses occur and how many losses and then calculate the exergy efficiencies. However, to the best of our

ratio of permeability values in different gas species. The separation principle is based on the difference in the transport properties of the gas molecules; for example, CO2 is a fast diffusing gas molecule compared to N2 and O2 in many membrane materials, such as glassy and rubbery polymers and carbon molecular sieve membranes.5 In addition, CO2 also has a relatively high molecular weight and a large quadruple moment, enabling it naturally to dissolve more easily into the membrane materials compared to some other gases, (e.g., N2 and CH4).5 In a post-combustion CO2 capture process, the purity of the captured CO2 in the permeate stream mainly depends upon the selectivity of CO2 over the other gas species, such as N2 and O2. Moreover, a high CO2 permeance is also required to reduce the membrane area. Gas separation membranes have many inherent benefits for carbon capture, such as less environmental impacts, no retrofit requirement on a power plant, and easy scale-up and operation,11,12 which show great potential to overcome the drawbacks of MEA-based systems. Recently, a lot of efforts have been put on the investigations of membrane systems for postcombustion carbon capture, in both membrane material development 1 0,11, 13−17 and techno-economic assessments.4,15,18−24 Brunetti et al. gave an overview on the polymeric membranes currently studied for their use in CO2 capture and their transport properties and provided a simple useful tool for an immediate and preliminary analysis on membrane technology for CO2 capture from flue gas.25 However, most of the studies are conducted in lab- and pilotscales because of the challenges related to the commercial applications as follows: (1) low CO2 concentration in flue gas results in large quantities of gases to be processed; (2) the gases should be cooled to below 100 °C in case the high temperature will rapidly damage the membranes; (3) some impurities, such as SO2 and NOx, in flue gas will pollute the membranes and reduce their performance over time; (4) creating a pressure difference across the membrane requires a significant amount of power, which will, in turn, increase the energy consumption;11 and (5) the CO2 permeabilities and selectivities of CO2/N2 for current membrane materials are not good enough to reach the economically available requirements for CO2 capture from flue gas. However, many studies have been focused on the development of new membrane materials, which could potentially overcome these challenges.4,12 Thus, from the viewpoints of industrial application and process design, it is very important to deeply understand the technical and economic benefits as well as the challenges of this new technology before it comes into commercialization along fundamental experimental results, which could provide the 4138


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knowledge, no literature has investigated this system with the exergy analysis method. In this work, a systematic method, which integrates process simulation, technological and economic analyses, and exergy analysis, is applied to evaluate a post-combustion CO2 capture with membrane systems. On the basis of the exergy analysis in a two-stage membrane process, the status of the energy consumption can be well-understood. The motivation of this work is to clarify the advantages and challenges of membrane systems for post-combustion carbon capture, understand how the key parameters influence the energy consumptions and capture costs, and indicate what the main obstacles are and how far from commercialization this technology is. The results can be used to identify the suitable membrane materials and processes for CO2 capture from flue gases in power plants in the near future. The overall framework of this study is shown in Figure 1.

Figure 2. Scheme of the exergy balance of a system.

consumption in specific separation requirements. The main energy consumption in a gas separation membrane process comes from the gas compressors and vacuum pumps, which is different from a MEAbased capture process, where the energy consumption is mostly used in a stripper. 2.2. Cost Model. The capture cost is calculated on the basis of Table 1. The cost model and parameters are mainly derived from the models reported by Merkel et al.23 and Hussain et al.18 The two key parts in the current model are annual capital-related cost (CRC) and annual variable operating and maintenance cost (VOM). Capital cost estimation in a gas separation membrane process is based on the evaluation of the major equipment, such as the compressors, vacuum pumps, expanders, and membrane units. The membrane price for post-combustion CO2 capture has been estimated to be U.S. $160/ m2,29 U.S. $50/m2,23 €50/m2,4 and U.S. $5/ft2,18 by different researchers, but it is still unknown today because there is no commercial membranes used as demonstration or industrial pilots. In this work, an assumed price of U.S. $50/m2 for membrane skid (including membrane modules, housings, frame, valves, and piping) was used for cost estimation. The operating and maintenance cost is mainly estimated on the basis of power cost, material maintenance cost, and membrane replacement in this process. 2.3. Exergy Analysis Method. The exergy calculation and exergy analysis method have been depicted in several publications.34−36 A schematic diagram of the exergy balance of a system is shown in Figure 2, where the illustrated system can be designated as different scales, such as a part of a unit, reactor, separator, exchanger, pump, section (consists of more than one unit, such as the gas compression section and gas purification section), process, plant, and industrial park. The exergy balance is a statement of the law of the degradation of energy, which is different from the energy balance, which is a statement of the law of the conservation of energy. Degradation of energy is equivalent to the irretrievable exergy loss as a result of all real processes being irreversible.36 Here, the total exergy losses (Bloss) were defined as two parts, i.e., external losses and internal losses.37,38 The external losses (Bexternal loss) are caused by the streams, which are not further used, such as the waste mass streams (gas, liquid, and solid discharges) and waste energy streams, typically the heat losses in flue gases and cooling water, while the internal losses (Binternal loss) are also known as thermodynamic irreversibilities, which are caused by friction, heattransfer, and spontaneous processes, such as chemical reaction and gas mixing. On the basis of Figure 2, the total exergy losses of a system can be expressed as eq 1

2. MATERIALS AND METHODS 2.1. Gas Separation Membrane Model. The models of the gas separation membrane and the detailed calculation procedure can be

Table 1. Economic and Process Parameters for CO2 Capture with Membrane category

value

units

Total Plant Investment (TPI) 50

membrane skid cost $/m2 (MC) compressor, vacuum 500 $/kW pump, and turboexpander cost (CC) fixed cost (FC) MC + CC $ base plant cost (BPC) 1.12 × FC $ project contingency 0.2 × BPC $ (PC) total facilities investBPC + PC $ ment (TFI) startup cost (SC) 0.1 × VOM TPI TFI + SC Annual Variable Operating and Maintenance Cost (VOM) contract and material 0.05 × TFI maintenance cost (CMC) local taxes and insur0.01 × TFI ance (LTI) labor cost (LC) 17 $/h membrane replacement 0.2 × 50 $/m2 costs (MRC) power cost (PC) 0.04 $/kWh VOM CMC + LTI + LC + MRC + PC annual capital-related 0.2 × TPI cost (CRC) CO2 capture cost (CRC + VOM)/CO2 captured $/ton of CO2 Other Assumptions compressor, turboex0.85 pander, and vacuum pump efficiency membrane lifetime 5 year operational time 7500 h/year

∑ Binternal loss = (∑ Bfeed − ∑ Bproduct + ∑ BQ − (∑ Bm loss + ∑ BQ loss )

found in the literature,22,33 and the membrane model is applicable to high flux asymmetric membranes in any flow patterns provided that it can meet the assumptions described in the literature.22 Process simulation is conducted by PRO/II software (Simulation Science, Inc.).4,22 The unit operation of the Flowsheet Optimizer in the software is used to optimize the operating parameters, such as the feed pressure and membrane area, based on the minimization of energy

−

∑ BW ) (1)

where Bfeed and Bproduct are the exergy of a feed stream and a product stream, respectively, BQ and BW are the exergy of heat and power, respectively, and Bm loss and BQ loss are defined as the exergy losses caused by a waste mass stream and waste heat stream, respectively. The exergy of mass stream consisting of chemical exergy and physical exergy can be calculated using the detailed equations in the literature35 4139


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based on the stream information (flow rate, temperature, pressure, and composition), which is from process simulation. The exergy of power/electricity is the same as its value. The exergy of heat can be calculated with Carnot efficiency.38 If an amount of heat Q (here, Q is only taken as an absolute value) is extracted from a hot stream (or unit) at temperature TH, the exergy BQH can be calculated as

⎛ T ⎞ BQH = ⎜1 − 0 ⎟Q TH ⎠ ⎝

Table 3. Comparison of Base Cases with and without MEABased Carbon Capture39 parameters gross electricity output auxiliary power consumption net electricity efficiency CO2 captured CO2 emitted specific stripper heat duty specific equivalent energy consumption

(2a)

If the heat is supplied to a cold stream (unit) at temperature TC, the exergy BQC can be calculated as

⎛T ⎞ BQC = ⎜ 0 − 1⎟Q ⎝ TC ⎠

(2b)

ΔTm =

⎛ T ⎞ BQCF = ⎜ 0 − 1⎟Q , ⎝ ΔTm ⎠

ΔTm =

T1 − T2 ln

T1 T2

(3a)

T1 − T2 ln

T1 T2

(3b)

where T1 and T2 are the heat flow temperatures of the inlet and outlet of the system, respectively, T 0 is the temperature at the thermodynamic standard state, ΔTm is the log mean temperature difference (K), and Q is the heat change of the heat flow from T1 to T2. In combination with Figure 2, two exergy efficiencies of the system are defined as eqs 4a and 4b. Bintotal is the total exergy feed into the system, including stream, heat, and power. ηex irre is the exergy efficiency that only considers the exergy losses caused by the irreversibilities, while ηex total is the exergy efficiency that considers the total exergy losses. ex ηirre =

ex ηtotal =

in Btotal − ∑ Binternal loss × 100% in Btotal

(4a)

in Btotal − (∑ Binternal loss + ∑ Bexternal loss ) × 100% in Btotal

(4b)

3. PROCESS DESCRIPTION 3.1. Gas Resource and Base Case. A power plant with an advanced supercritical boiler and turbine delivering 819 MWe Table 2. Characteristics of Flue Gas in a Post-combustion Coal Power Plant39 parameters flue gas flow rate temperature pressure composition O2 CO2 SO2 NOx H2O Ar N2 particulate

value

unit

781.8 50 1.016

kg/s °C bar

3.65 13.73 85 120 9.73 0.005 72.86 8

vol % wet vol % wet mg/Nm3 mg/Nm3 vol % wet vol % wet vol % wet mg/Nm3

with capture

unit

819 65 754 45.4

684.2 135 549.2 33.4 144 104.7 3.73 1.422

MWe MWe MWe % LHV kg/s kg/MWh MJth/kg of CO2 MJe/kg of CO2

763

(gross) without carbon capture was chosen as the simulation basis, where the final net power plant output is 754.3 MWe by excluding the auxiliary power and yielding a net cycle efficiency of 45.5%. The feedstock coal is South African Douglas Premium 2 with a flow rate of 65.8 kg/s. The flue gas composition of the coal power plant is shown in Table 2.39 Flue gas was pretreated to remove ash and acid gases, such as SO2 and NOx, before feeding into a membrane system,4,23 and only the main compositions of CO2, N2, O2, and H2O were considered to simplify the process simulation.4,23 If 90% of the total CO2 in the flue gas (capture ratio) is captured, the capacity of CO2 capture is about 4 million tons per year. Another important issue is the evaluation of the energy consumption of different technologies. Table 3 shows the comparison of the base case with and without the MEA-based capture system.39 Many researchers have compared different new technologies with the MEA-based process considering the energy consumption and capture cost.18,21,24,39 It has to be noted that two parts of energy consumption are involved in a MEA process:22 one is the steam extracted from steam turbines in a power plant, which is used for stripping CO2 from the rich solvent, and the other is the power demands for the gas blowers, solvent pumps, CO2 compressors, and other driving machines. However, in a gas separation membrane process, the main energy consumptions come from the power demands of the compressors and vacuum pumps. Thus, power consumption is usually employed for energy assessment in a membrane process. To easily compare to MEA-based technology, especially with the same benchmark, energy consumption in a MEA-based process was converted to an equivalent power (electricity) consumption based on electricity penalty for CO2 capture in the current work. The net electricity loss related to CO2 capture (including the CO2 compression to 110 bar) was estimated to be 204.8 MWe in the MEA solvent base case, as shown in Table 3, and the CO2 captured is about 144 kg/s;39 therefore, the specific equivalent power consumption is 1.422 MJe/kg of CO2 including the power for CO2 compression and 1.09 MJe/kg of CO2 excluding the CO2 compression. 3.2. Process Configurations. Figure 3 depicts different flow sheet configurations of gas separation membrane processes for post-combustion carbon capture. Figure 3a is a schematic diagram of a one-stage gas separation process with a turboexpander. The flue gas is fed into the membrane unit at a higher pressure relative to the permeate side. A proportion of CO2 can easily pass through the membrane, while the other gases (such as O2 and N2) will be mostly retained because of their lower permeabilities. The retentate stream is subsequently sent into an expander to recover power, which can reduce the

where eqs 2a and 2b can be used to calculate the exergy of the waste heat discharged from a system to the environment. Moreover, the exergy of hot and cold heat flows BQHF and BQCF can be calculated by eqs 3a and 3b

⎛ T ⎞ BQHF = ⎜1 − 0 ⎟Q , ΔTm ⎠ ⎝

without capture

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Figure 3. Process configurations of gas separation membrane processes for carbon capture.

Gas permeance (Nm3 m−2 bar−1 h−1 or GPU) was used for calculation instead of permeability (Barrer) in the current work because of the uncertainty of the thickness of the membrane materials. The following performance indicators related to the operating and capital costs as well as the thermodynamic efficiency are used to investigate the influences of the above parameters: power consumption for compressors and vacuums pumps and recovered by retentate stream expanders and CO2 compressors (MJe/kg of CO2), membrane area demand (million m2) or specific area demand per captured CO2 flow rate (m2 h kg−1 of CO2 captured), capture cost (U.S. $/ton of CO2 captured), exergy values of steams, units, and processes (MJ/h or MW), and thermodynamic efficiency (exergy efficiency) of units in the process (%). Among them, power consumption is expected to be the major contribution to the operating costs, while membrane area demand and some key power equipment, such as the compressors and vacuum pumps, are mainly used to estimate the capital costs. The thermodynamic efficiency reflects the status of energy use.

net power consumption. Most polymer membranes are quite challenging to achieve a high CO2 purity and recovery at the same time using a single-stage membrane unit. Therefore, stream recycling (or circulation) and multi-stage membrane units are applied to achieve CO2 streams amenable to geologic storage with lower energy consumption. Figure 3b shows a onestage process with both a turboexpander and stream circulation, and Figure 3c is a basic flowsheet of a two-stage membrane process. However, if a multi-stage process is used, the process will become much more complex and the operating cost may increase significantly. For example, recompression and suction will consume additional power. 3.3. Parameters. The key parameters related to the membrane materials and process conditions are listed as follows, which will be investigated by process simulation: CO2 permeance, 1−8 Nm3 m−2 bar−1 h−1; CO2/N2 selectivity, 50− 220; CO2 capture ratio, 55−90%; CO2 purity, 90 and 95%; pressure of the second stage, 2, 2.5, and 3 bar; flue gas compression pressure, varying depending upon the specific separation requirements and membrane parameters; stream circulation ratio, 0.2−1.0; capture load distribution between the first and second stage for a two-stage process (CO 2 concentration in permeate from the first stage), 50−80%; and process configurations, one-stage with and without circulation and two-stage processes. Moreover, to simplify the process simulation, the following parameters are kept as constant: O2 permeance, 0.02 Nm3 m−2 bar−1 h−1;18 H2O permeance, 2.25 × 10−9 Nm3 m−2 bar−1 h−1;18 permeate pressure, 0.2 bar;23 compression pressure of CO2 for transport, 110 bar; and efficiencies of power equipment (compressor, expander, and vacuum pump), 85%.

4. PARAMETRIC STUDY AND PERFORMANCE ASSESSMENT 4.1. Theoretical Energy Consumption for CO2/N2 Separation. On the basis of the thermodynamic principle, the theoretical energy consumption for CO2/N2 separation is calculated by eq 5 at an isobaric and isothermal process and ideal gas state. It represents the minimum energy demand from the gas mixture to pure CO2 and N2. 4141


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Figure 4. Specific energy consumption. Conditions: one-stage membrane separation with a turboexpander (Figure 3a). Specifications: capture ratio, 85%; CO2 purity, 90% (mol); and permeance of CO2 and N2, 1 and 0.005 Nm3 m−2 bar−1 h−1, respectively.18

Figure 6. Influence of the circulation ratio on the energy consumption and membrane area. Simulation conditions: CO2 permeance, 4 Nm3 m−2 bar−1 h−1; CO2/N2 selectivity, 100; pressure of permeate, 0.2 bar; CO2 capture ratio, 70%; CO2 purity, 90% (mol); and mole circulation ratio, mole flow of circulation stream/CO2 product stream.

difference between the feed and permeate sides, and the energy consumption will be extremely high because of a low CO2 feed concentration and a high volume of gas stream to be treated. Therefore, besides continually improving the performances of the membrane materials, permeate stream circulation to increase the CO2 concentration in the feed stream or multiple stage separation to decrease the irreversibility of the whole process could be two feasible and practical solutions to reduce the energy consumption. This will be discussed later. 4.2. Influence of Membrane Inherent Parameters on the Energy Consumption and Membrane Area. To investigate the influence of membrane parameters (permeance and selectivity) on the energy consumption and required membrane area, a one-stage membrane process is designed, as shown in Figure 3a. Process simulation is conducted to determine the feed gas pressure and required membrane area to achieve a 90% CO2 capture ratio and a 90% (mol) CO2 purity. CO2 permeance and CO2/N2 selectivity are two key parameters in post-combustion CO2 capture, which are widely used to characterize the membrane performance expressed on Robeson “upper bounds”.9 In general, the higher the permeability, the lower the selectivity. There is a big challenge in the development of membrane materials for gas separations. Luis et al. summarized systematically the progress of the CO2selective membrane materials and provided some useful information of the state-of-the-art membrane-based technology for CO2 separation. Generally, it is concluded that the maximum permeabilities are around 10 000 Barrer with maximum selectivities of 100. 6 However, the recently developed composite membranes made from poly(vinyl alcohol) or poly(vinyl amine) showed a high selectivity of >100.18,40 The simulation results in Figure 5 indicate the influences of CO2 permeace and CO2/N2 selectivity on the energy consumption and membrane area in a gas separation process. The higher CO2 permeance can reduce both the specific energy consumption and membrane area considerably. For CO2/N2 selectivity, the specific energy consumption decreases with an increase of the selectivity; however, more membrane area is required. Thus, continuously increasing selectivity will not affect the separation performance so much, which means that permeance could be the dominating parameter because of the opposite tendency of the energy

Figure 5. Influence of CO2/N2 selectivity on the energy consumption and membrane area. Simulation assumptions and specifications: capture ratio, 90%; CO2 purity, 90% (mol); one-stage process (Figure 3a); Ppermeate, 0.05 bar; and CO2 permeance value, 1−5 Nm3 m−2 bar−1 h−1. Simulation results: pressure of the flue gas compressor (bar); membrane area, A1, A2, A3, A4, and A5 (million m2); and specific energy consumption, E1, E2, E3, E4, and E5 (MJe/kg of CO2); In the figure, 1, 2, 3, 4, and 5 refer to CO2 permeance values of 1, 2, 3, 4, and 5 Nm3 m −2 bar−1 h−1, respectively, without considering CO2 compression.

⎛ 1 1 ⎞⎟ Etheor = nRT0⎜⎜xCO2 ln + x N2 ln xCO2 x N2 ⎟⎠ ⎝

(5)

The results in Figure 4 show that the theoretical specific energy consumption or minimum energy consumption for CO2 separation from a CO2 and N2 mixture is relatively lower, ranging from 0.08 to 0.26 MJe/kg of CO2, while the feed CO2 concentration varies from 0.5 to 0.025 (mol fraction). The basic principle is that the gas separation is an entropy process, while contrarily, gas mixing is an irreversible spontaneous entropy production process. For a membrane process, the specific energy consumption has a similar tendency to the theoretical curve but decreases significantly with the increase of the feed CO2 concentration. Hence, a large improvement space for CO2 separation is possible with membrane technologies, especially for lower CO2 partial pressure. The driving force for CO2 transport through a membrane is the CO2 partial pressure 4142


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Figure 7. Influence of the capture load distribution for a two-stage process.

and membrane area. Line AB in Figure 5 shows the specific equivalent energy consumption (1.09 MJe/kg of CO2) of the MEA-based case without a CO2 compressor. If the operational point of the membrane process is above line AB, the membrane system might not compete with MEA-based systems when the energy consumption is only considered. 4.3. Permeate Stream Circulation. The configuration with the circulating (or recycling) part of the permeate stream to the feed (shown in Figure 3b) can improve the feed CO2 concentration, which could be one of the potential options to decrease the energy consumption by increasing the CO2 partial pressure difference across the membrane. Figure 6 shows the influence of the permeate circulation ratio on the energy consumption and membrane area. It was found that the specific energy consumption (including the power consumption for CO2 compression to 110 bar) decreased significantly with the increase of the circulation ratio, but the tendency slowed when the ratio was higher than 0.6. Although recycling part of permeate gas will increase the load of the gas compressor to compress the additional feed gas, the increase of the CO2 concentration in feed gas can reduce the required pressure ratio for the membrane system, which can offset the total energy consumption. However, the predominance becomes weak with the increase of the circulation ratio. On the other hand, increasing the load of the gas processing also increases the required membrane area consequently; this will increase the capital costs. In conclusion, the prominent benefit of stream circulation is to improve the CO2 purity in the product to achieve the specific requirement for CO2 transport and storage (e.g., >90% in ENCAP project41). The basic principle of the permeate circulation is similar to the reflux operation in a distillation tower. However, using circulation operation should also balance the other additional costs, such as an increased feed gas flow rate and membrane area, and then an optimized circulation ratio can be obtained. 4.4. Influence of the Capture Load Distribution in a Two-Stage Membrane Process. Because of the limitation of the currently developed membranes, it is still hard to reach both 90% capture ratio and 95% CO2 purity using a one-stage membrane unit in a post-combustion carbon capture process. In Figure 5, we have assumed a very high CO2 permeance of 5 Nm3 m−2 bar−1 h−1 and a high selectivity of 250. However, most of the reported membrane materials cannot reach such high performances. For example, the properties of some

Figure 8. Influence of the membrane price on the capture cost. Simulation conditions: two-stage process (Figure 3c); CO2 permeance, 6 Nm3 m−2 bar−1 h−1; CO2/N2 selectivity, 70; pressure into the second membrane, 2.5 bar; pressure of the permeate, 0.2 bar; CO2 capture ratio, 90%; and CO2 purity, 95% (mol).

Figure 9. Influence of the capture ratio on the capture cost. Simulation conditions: two-stage process (Figure 3c); CO2 permeance, 4 Nm3 m−2 bar−1 h−1; CO2/N2 selectivity, 70; pressure into the second stage, 2.5 bar; pressure of the permeate, 0.2 bar; CO2 purity, 95% (mol); including CO2 compression; and membrane price, U.S. $50/m2 (capture cost 1) and U.S. $100/m2 (capture cost 2).

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Figure 10. Profiles of (left) energy consumption and (right) cost with membrane parameters. Simulation conditions: two-stage capture (Figure 3c); pressure of the second stage, 2−3 bar; pressure of the permeate, 0.2 bar; CO2 purity, 95% (mol); capture ratio, 90%; including CO2 compression; and membrane price, U.S. $50/m2.

polymer membranes are as follows: poly(acetylene) dense films (selectivity < 25), polyarylates (selectivity < 25), and 6FDAbased polyimides (selectivity < 30). Poly(ethylene oxide) (PEO) and PEO-based blends and co-polymers or cross-linked polymers exhibited a better performance for CO 2 /N 2 separation, where the CO2/N2 selectivity can reach 60 with the corresponding CO2 permeability up to 200−250 Barrer.11 Yave et al. reported novel PEO−poly(butylene terephthalate) (PBT)-based CO2-philic polymer membranes, which showed extremely high permeability at 400−750 Barrer and CO2/N2 selectivity at 40−50.28 Recently, Kim et al.42 reported a polyvinylamine (PVAm) fixed-site-carrier (FSC) composite membrane, and its permeance is as high as 5 Nm3 m−2 bar−1 h−1 with the selectivity of CO2/N2 reaching 500. Merkel et al.23 and Zhao et al.4 recommended that a suitable selectivity should be in the range of 20−120. Too high of a selectivity will increase the required membrane area tremendously, as shown in Figure 5. Stream circulations can achieve a relatively higher CO2 purity and capture ratio; however, an additional membrane area is required, and more feed gas is needed to be compressed. Alternatively, multi-stage membrane separation processes provide a feasible solution but certainly make the process more complex. A key point that cannot be overlooked is that the permeate stream comes out from the first stage needing to be recompressed before feeding into the second stage, as shown in Figure 3c, which will additionally increase the energy consumption and the capital cost of the power equipment. Thus, the process should be optimized to reach a minimum energy consumption and membrane area. For a two-stage membrane process, the capture load distribution between the first and second stages is found to be another important parameter. Some authors had reported the permeate CO2 concentration in the first stage, for example, 50,23 42−55,18 and 61%,4 but no detailed analysis has been conducted. Therefore, the influences of the capture load distribution will be investigated and analyzed in the current work. The simulation basis is chosen as follows: CO2 permeance is 2 and 6 Nm3 m−2 bar−1 h−1, and the CO2/N2 selectivity is 70. The CO2 capture ratio and CO2 purity are 90 and 95% (mol), respectively. The feed pressure of the first stage is determined by minimizing energy consumption, while the feed pressure of the second stage is assumed as 2.5 bar. Moreover, the permeate pressure in both stages is set to 0.2 bar

based on the industrial consideration of vacuum operation.23 The relatively pure CO2 is compressed to be in the supercritical state (∼73.8 bar) and then pumped to 110 bar for transport and storage.43 The main aim here is to investigate the influence of the process configuration on the energy consumption and capture cost. The simulation results are shown in Figure 7, where the abscissa is the permeate CO2 concentration in the first stage, xfirst stage. It was found that the capture load distribution between the first and second stages influences the total energy consumption and total membrane area significantly. In a lower xfirst stage, because of a high gas flow pass through the membrane, a larger membrane area is required but a lower compression pressure is needed. With the increase of xfirst stage, the required membrane area decreases but the required feed pressure of the flue gas compressor increases gradually, which results in an increase in the specific energy consumption. It indicates that the lowest energy consumption is 1.8−1.9 MJe/ kg of CO2 at the xfirst stage of 60% in Figure 7a and 65% in Figure 7b. On account of the synergistic effect by power consumption and membrane area required, the lowest capture cost is found to be U.S. $50/ton of CO2, including the CO2 compressor from Figure 7a, and the optimal xfirst stage is about 72%. Below 72%, the capture cost rises because of the excessive membrane area, and above 72%, the capture cost rises because of an excessive energy consumption caused by the higher pressure in flue gas compression. In Figure 7b, the lowest capture cost is U.S. $43/ ton of CO2 with a higher CO2 permeance. The results indicated that the capture cost can be greatly reduced with U.S. $7/ton of CO2 by improving the CO2 permeance from 2 to 6 Nm3 m−2 bar−1 h−1 at the same CO2/N2 selectivity of 70. It is worth noting that the optimal xfirst stage depends upon the membrane parameters, the price of the membrane and electricity, but the tendency is similar. With regard to the capture cost, the gas membrane separation technology has shown predominance compared to MEA technology, because the typical costs of CO2 capture with a MEA-based solvent in power plants are in the range from 20 to U.S. $80/ton of CO2 based on the International Energy Agency (IEA) data44 and from 40 to U.S. $100/ton of CO2 based on the estimation by Merkel et al.23 However, by only considering the specific energy consumption, membrane technology still shows a higher 4144


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Figure 11. Schematic diagram of a two-stage membrane process. Simulation conditions: CO2 permeance, 6 Nm3 m−2 bar−1 h−1; CO2/N2 selectivity, 70; pressure into the second stage, 2.5 bar; pressure of the permeate, 0.2 bar; CO2 purity, 95% (mol); capture ratio, 90%; and including CO2 compression.

compression pressure, while the membrane area increases slightly and then decreases rapidly. As a result of the trade-off between the energy consumption and membrane area, the capture cost decreases slightly at first to an optimal range of 65−70% with the lowest cost of U.S. $40/ton of CO2. After that range, the cost increases rapidly, even though the membrane area still drops continuously, especially after a 80% capture ratio. It means that the energy consumption caused by the increase of the flue gas pressure (proportional to the energy demand by the flue gas compressor) makes more contribution to the total capture cost than the membrane area. It is noted that pursuing an excessively high capture ratio probably causes a much higher capture cost while keeping CO2 purity as a constant (95% in this work). For example, the capture cost will reduce to about 9% (from 45.1 to U.S. $41.3/ ton of CO2) when the capture ratio decreases from 90 to 80%. Therefore, the optimal capture ratio was found to be in the range of 65−70% by considering the total capture capacity and the cost in a real process design. A similar tendency was proposed with an optimal range of 70−85% by Merkel et al.23 However, the optimal capture ratio also depends upon the operating parameters and process configurations. For example, if the membrane price is assumed as U.S. $100/m2 (as in the capture cost 2 curve in Figure 9), the contribution by the energy consumption will be less and the optimal range of the capture ratio moves to 80−85% with the lowest cost of U.S. $48/ton of CO2. 4.6. Profiles of Capture Performance Related to the Membrane Parameters. Three-dimensional profiles reflecting the relationship between permeance, selectivity, and specific energy consumption and capture cost for a two-stage process are created, as shown in Figure 10. The ranges of the parameters were chosen on the basis of the recommendations from the literature4,23 and the above analysis results, and CO2 permeance is designated from 2 to 8 Nm3 m−2 bar−1 h−1 with CO2/N2 selectivity from 50 to 120. Such ranges represent the parameters that have industrialization potentials and also have good feasibilities considering the progress on the membrane material development in recent years.11,13,28,47,48 The profiles in Figure 10 provide the visual change trends of the capture energy consumption and cost in different membrane parameters. The curves could also be used for screening or developing membrane materials or judging whether membranes with specific parameters have any economic feasibility

value (i.e., 1.8−1.9 MJe/kg of CO2) compared to a MEA-based system with 1.423 MJe/kg of CO2 (including power for CO2 compression) based on the benchmark of MEA technology39 and 1.57 MJe/kg of CO2 reported by Abu-Zahra et al.45 Of course, the capture energy consumption can be further reduced by improving the process configuration and optimizing the operational parameters or integrating the capture process with a power plant, for example, using income boiler air as a sweep stream.18,21,23 However, this will increase the process complexity in real industrial applications. Similar to MEA technology, one of the main issues is the low CO2 feed concentration, which means that membrane technology for CO2 capture from flue gas still faces many challenges. The price of the membrane is another key parameter for the gas separation membrane. Merkel et al.23 discussed the cost of the membrane and indicated that today’s commercial gas separation membrane skids range from 500 to U.S. $700/m2; however, considering the lower pressure operational condition and large scale of CO2 capture, around U.S. $50/m2 was possible.23 However, it is hard to figure out the price of the membrane before such technology has been run at least in a pilot-scale plant continually, especially for some new materials studied in the laboratory. To understand how the membrane price influences the capture cost, several cases are simulated for a two-stage membrane process based on different membrane prices, and the results are shown in Figure 8. If the membrane price decreases from 200 to U.S. $50/m2, the total capture cost reduces to U.S. $12.7/ton of CO2 (from 55.6 to U.S. $42.9/ton of CO2) to the greatest extent, equivalent to 25% reduction. For a 800 MW power plant, the total cost of about U.S. $50 million/year will be saved. Thus, reducing the membrane price as low as possible is an important goal in this field. 4.5. Influence of Capture Ratios for a Two-Stage Process. Capturing CO2 from flue gases in a power plant is not like traditional gas processing or purification, there are no strict requirements on the decarbonized sweet gas, which means no strict requirement on the CO2 concentration in the sweet gas or CO2 capture ratio. In general, the recommendation and calculation benchmark of the capture ratio is around 90%.39,46 Figure 9 shows the influence of the capture ratio on the capture energy consumption, membrane area, and capture cost considering two membrane prices. With the increase of the capture ratio, the energy consumption increases evidently, which is mainly derived from the increase of the flue gas 4145


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compared to a MEA-based system or other capture technology. On the basis of the assumptions and the models used in the current work, it was found that the optimal CO2/N2 selectivity is 70−90 and the higher the CO2 permeance, the lower the capture cost.

Table 4. Thermodynamic Variables of the Membrane Capture Process

5. EXERGY ANALYSIS OF A TWO-STAGE GAS SEPARATION MEMBRANE PROCESS Different from a MEA-based capture system, which is a typical chemisorption process, the driving force of separating CO2 from the gas mixture is well-known as the partial pressure difference of the gases between the feed and permeate sides. Merkel et al.23 compared the different modules of membrane processes and concluded that power consumption is still the key issue in a membrane separation system for post-combustion carbon capture. On the basis of our above analysis, the specific energy consumption is still as high as 1.7−1.8 MJe/kg of CO2 (1.422 MJe/kg of CO2 for the MEA-based case), even using a two-stage process with higher CO2 permeance (6 Nm3 m−2 bar−1 h−1) and high CO2/N2 selectivity (100), which is almost the upper limit of the reported polymer membranes.11 Thus, understanding the energy consumption status is another key point for the membrane process based on the thermodynamic principles. On the basis of the flow sheet in Figure 3c, a schematic diagram of a two-stage membrane system is represented in Figure 11 to illustrate the mass and energy flows involved in the system. The main material streams, power flows, and heat flows have been marked in the diagram. The input and output of the flow descriptions and the thermodynamic variables of the membrane process are listed in Table 4. With the method described in section 2.3, the exergy values of streams, power flows, and heat flows are calculated. According to the exergy balance shown in Figure 2, the exergy loss of each unit is obtained. The total net power demand is estimated to be 261 MW, as shown in Table 4. Most of the power consumption is still required by the flue gas compressors, and the total power consumption used for creating the pressure difference between the feed and permeate streams in the first stage is about 120 MW, including the power consumption for the vacuum pumps and the recovered power from the expanders, which account for 46% of the total net power consumption. In the second stage, the gas flow rate (mass) is only 27% of the feed flue gas; therefore, the total net power consumption reduces to 72 MW (about 60% of the net power demand in the first stage), where the main part is ascribed to the recompression of the vacuum permeate gas coming out from the first stage. Thus, improving the pressure of the permeate side in the first stage and reducing the feed pressure of the second stage is supposed to reduce the net power consumption. Moreover, compressing and pumping the captured CO2 to 110 bar requires 69.7 MW of power, accounting for 27% of the total energy consumption, which can be reduced to 66.6 MW if the CO2 purity can increase to 99% (mol). However, this is obviously not advisable because purification of CO2 from 95 to 99% needs more energy or even more stages in a membrane separation process. For instance, the specific energy consumption for CO2 capture increases from 1.37 to 1.53 MJe/kg of CO2, excluding the CO2 compression power, when the CO2 purity increases from 90 to 95% (mol), approximately 10% growth in this case. Thus, just like the capture ratio, a reasonable CO2 purity for capture and storage should be traded off too.

S1: flue gas to compressor 1 S2: flue gas from compressor 1 to membrane 1 S3: permeate from membrane 1 to compressor 2 S4: permeate from compressor 2 to membrane 2 S5: retentate from membrane 2 to expander 2 S6: retentate from expander 2 to mixer S7: retentate from membrane 1 to expander 1 S8: retentate from expander 1 to mixer S9: permeate from membrane 2 to compressor 3 S10: permeate from compressor 3 to transport

flow rate (kg/s)

flow description

Streams 748.88 737.77

pressure/ temperature (bar/°C)

exergy (MW)

1.02/25 3.82/25

56.24 135.20

199.03

0.2/25

39.78

199.03

2.5/25

71.44

49.53

2.5/25

5.62

49.53

1.02/25

1.10

538.74

3.82/25

71.16

538.74

1.02/25

9.15

149.50

0.2/25

50.73

149.50

110/25

98.66

Power Consumption W1: power for compressor 1 W2: power for vacuum 1 W3: power for compressor 2 W4: power for vacuum 2 W5: power for compressor 3 W6: power from expander 1 W7: power from expander 2 net power consumption Heat Flows Q1: heat from compressor 1 134.59 Q2: heat from vacuum 1 52.89 Q3: heat from compressor 2 45.74 Q4: heat from vacuum 2 29.26 Q5: heat from compressor 3 87.88 Q6: cooling from expander 1 30.55 Q7: cooling from expander 2 3.92 total exergy of heat flows Exergy Losses (Irreversibilities) membrane 1 + vacuum 1 membrane 2 + vacuum 2 compressor 1 compressor 2 compressor 3 expander 1 expander 2 total exergy loss (irreversibilities)

112.06 52.50 45.43 28.95 69.70 −45.10 −2.43 261.10 24.15 10.34 8.13 4.77 15.13 6.41 0.22 69.16 76.75 44.04 33.10 13.76 21.78 16.91 2.09 208.43

(66.42) (39.27) (8.95) (5.63) (6.64) (10.50) (1.87) (139.28)

Figure 12 shows the exergy efficiency of each unit in the whole process. Figure 12a indicates the exergy efficiency of each unit under the premise that the exergy of the heat flows produced by the compression processes can be recovered, and the efficiency is calculated by eq 4a. This means that the exergy losses are mainly caused by the irreversibilities of the processes. Figure 12b shows the exergy losses considering both the external losses and irreversibilities, and the efficiency are calculated with eq 4b. The exergy loss of each unit in Figure 4146


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Figure 12. Exergy losses and exergy efficiencies of units in the membrane process.

changes of the fluids. For example, the power demand for compressing CO2 to 110 bar is as high as 69.7 MW, but the irreversible losses are much lower because multi-stage compression is used. Another interesting result is the proportion of power consumption of each unit versus the total power consumption and the exergy losses in each unit versus the total exergy losses. The percentage of the irreversible losses of the CO 2 compressor only accounts for about 5% of the total irreversible losses, although its power demand accounts for 27% of the total power consumption. Contrarily, the irreversible losses in the first stage account for as high as 54% of the total irreversible losses, while its power demand only accounts for 46% of the total power consumption. For the second stage, these two values are 32 and 28%, respectively. Thus, it is worth noting that more efforts should be focused on the membrane separation process itself instead of the CO2 compression unit,

12a is smaller than that in Figure 12b, and correspondingly, the exergy efficiency in Figure 12a is higher than that in Figure 12b. In Figure 12a, the largest irreversibility occurs in the firststage membrane unit and then the second-stage membrane unit. This irreversibility is derived from the high pressure ratio between the feed and permeate sides based on the principle that the gas separation is an entropy-decreasing process. It means that the exergy losses are irreversible when pressurebased exergy is transformed to concentration-based exergy. Reducing the pressure ratio would decrease the irreversible losses, but it will be restricted by the required CO2 capture ratio and CO2 purity. Increasing the stage number of the membrane process will reduce the pressure ratio for each stage because it will be closer to a ideal reversible process, but more membrane area and more power equipment will be needed, which will increase the capital cost and make the process more complex. Moreover, the irreversible losses in the other power equipment are relatively lower, which are determined by the pressure 4147


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exergy efficiency has less potential for energy savings, although it costs around 27% of the total power consumption. In conclusion, improving the performance of the membrane materials and optimizing the separation configuration are the two main means to enhance the competitions of the membrane technology, and more efforts should be focused on reducing the membrane costs and increasing the long-term durability of the membrane materials to promote the advantages of membrane technology. This study discusses some key issues of membrane processes from different aspects, and the results will be important for both membrane material development and future industrial applications.

which has less potential for energy savings on the whole CCS chain. In general, the exergy efficiency of the whole capture process is 80% by only considering the irreversible losses and decreases to 53% when accounting for both internal and external exergy losses. This means that recovering and reusing the heat produced from this pressure swing process will increase the exergy efficiency of the whole system, for example, recovering part of the heat produced by compressing the gas to preheat the boiling water in a power plant.

6. SUMMARY AND CONCLUSION In this work, parametric study, capture cost, and exergy analysis were adopted to assess the technological, energetic, and economic performances of gas separation membranes for post-combustion carbon capture. As for membrane materials, higher CO2/N2 selectivity can reduce the energy consumption but increase the required membrane area. An optimal range for selectivity is found to be 70−90 for post-combustion carbon capture in a coal-fired power plant using a two-stage membrane process under certain assumptions. Increasing the CO2 permeance can reduce the required membrane area to decrease the capture cost; thus, higher permeance is preferred. Emerging materials will reduce the energy consumption and capture cost gradually and make it possible to realize industrialization successfully; for example, chemical modification of polymeric membranes is a promising approach for greatly enhancing the separation performance, but the thermally and chemically robust and long-term stability, resistant to plasticization and resistant to aging of the polymer membrane, should also be considered. With regard to membrane process configurations, multi-stage and stream circulation processes are supposed to satisfy the required CO2 capture ratio and CO2 purity with lower energy consumption. This study shows that, when CO2/N2 selectivity is low, achieving a higher CO2 purity (95%) with a one-stage process is unrealizable because an extremely high pressure ratio is required. With a two-stage process, the specific energy consumption and cost can be reduced, and it was found that the capture load distribution between the first and second stages affects the separation performance considerably. However, more stages will not be recommended because recompression and suction would increase the energy consumption and process complexity. Additionally, too high of a CO2 capture ratio and CO2 purity would increase the specific energy consumption, so that suitable ranges should be addressed. Besides using multi-stage processes and stream circulation, hybrid processes by combining membrane systems with the other separation technologies, such as solvent absorption and cryogenic separation, could be promising candidates for CO2 capture from post-combustion power plants. The capture cost using membrane technology has predominance compared to MEA-based systems in this work, but membrane technologies still suffer from high energy consumptions even using a multi-stage process or stream circulation. Thus, the energy consumption is still a challenge, at least for the currently developed membrane materials. The exergy analysis shows that the main energy consumption bottleneck of carbon capture is the membrane processes, thus improving the exergy efficiencies by optimizing the pressure ratio and process configuration, and integration of the heatexchanger network will be important for developing a membrane process. Relatively, CO2 compression with high

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AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS This publication has been produced with support from the BIGCCS Centre, performed under the Norwegian research program Centers for Environment-Friendly Energy Research (FME). The authors acknowledge the following partners for their contributions: Aker Solutions, ConocoPhillips, Det Norske Veritas, Gassco, Hydro, Shell, Statoil, TOTAL, GDF SUEZ, and the Research Council of Norway (193816/S60). This work was also supported by the National Basic Research Program of China (973 Program) (2013CB733506) and the Key Program of the National Natural Science Foundation of China (21036007). We also thank Danahe Marmolejo Correa for fruitful discussions on the exergy analysis method and Prof. Suojiang Zhang for his meaningful suggestions.

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Post-combustion Carbon Capture with a Gas Separation Membrane

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