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Carbon dioxide recovery from carbonate solutions using bipolar membrane electrodialysis
Separation and Purification Technology 101 (2012) 49–59
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Separation and Purification Technology journal homepage: www.elsevier.com/locate/seppur
Carbon dioxide recovery from carbonate solutions using bipolar membrane electrodialysis Atsushi Iizuka a, Kana Hashimoto b, Hiroki Nagasawa c, Kazukiyo Kumagai c, Yukio Yanagisawa c, Akihiro Yamasaki d,⇑ a
Research Center for Sustainable Science and Engineering, Institute of Multidisciplinary Research for Advanced Materials, Tohoku University, 2-1-1 Katahira, Sendai, Miyagi 980-8577, Japan Department of Chemical System Engineering, Faculty of Engineering, The University of Tokyo, 7-3-1 Hongo Bunkyo-ku, Tokyo 113-8656, Japan c Department of Environment Systems, Graduate School of Frontier Sciences, The University of Tokyo, 5-1-5 Kashiwanoha Kashiwa, Chiba 277-8563, Japan d Department of Materials and Life Science, Faculty of Science and Technology, Seikei University, 3-3-1 Kichijoji-Kitamachi, Musashino, Tokyo 180-8633, Japan b
a r t i c l e
i n f o
Article history: Received 11 June 2012 Received in revised form 12 September 2012 Accepted 12 September 2012 Available online 23 September 2012 Keywords: CO2 recovery Bipolar membrane electrodialysis Power consumption Cost estimation Global warming Carbon capture and storage
a b s t r a c t Process design and cost estimation were conducted for CO2 recovery by liquid absorption with alkaline solution, coupled with bipolar membrane electrodialysis for CO2 gas separation and alkaline solution regeneration. The electrodialysis performances and power consumption for CO2 recovery and alkaline solution regeneration were examined using laboratory-scale electrodialysis equipment under various conditions; degree of CO2 absorption, CO2 recovery ratio, alkaline concentration, and type of cation exchange membrane. The total cost of CO2 recovery was estimated for the treatment of CO2 emitted from a 400 MW coal-fired thermal power plant. The minimum cost for the optimum condition was about US$180 per ton-CO2 removed, which is higher than that for the conventional process using amine absorption and thermal desorption. However, a sensitivity analysis indicates that the overall cost could be significantly reduced if the cost of membranes was reduced. Ó 2012 Elsevier B.V. All rights reserved.
1. Introduction Carbon capture and storage (CCS) has been widely recognized as an effective countermeasure to greenhouse gas emissions that mitigates global warming [1,2]. In the CCS process, carbon dioxide emitted from point sources would be captured and transported to storage sites, where liquid CO2 would be injected into either underground or the ocean waters. The main concern regarding CCS implementation is the need to reduce cost and power consumption for capturing CO2 from flue gas. More than 80% of the total power consumption and cost for CCS can be attributed to those for CO2 capture process. There are several methods for capturing and recovering CO2 from point sources, e.g., chemical absorption [3–8], physical adsorption including pressure swing adsorption (PSA) [9,10] and temperature swing adsorption (TSA) [11,12], membrane separation [13–17]. Among these methods, chemical absorption method using amines has been considered to be the most promising method from the viewpoints of both cost and energy penalty [1]. Aqueous solutions of mono ethanol amine (MEA) and its similar substances are used for capturing ⇑ Corresponding author. Tel.: +81 422373887. E-mail address: [email protected] (A. Yamasaki). 1383-5866/$ - see front matter Ó 2012 Elsevier B.V. All rights reserved. http://dx.doi.org/10.1016/j.seppur.2012.09.016
CO2 in the flue gas emitted from point sources. These amines form CO2-amine compounds at the lower temperature conditions. It is necessary to increase the temperature of the solution to recover CO2. The heating process would require a lot of heat; almost all the power consumption for the CO2 capture can be attributed to this regeneration process, which in turn increases the cost for CCS [1]. Electrochemical methods have been proposed for recovery of CO2 from chemically absorbed solutions [18–20]. Electrodialysis using bipolar membranes is utilized for these recovering methods of CO2 from solutions. In a previous paper [18], we proposed a method for recovering CO2 from carbonate solutions and demonstrated its recovery performances from sodium carbonate with a laboratory-scale experimental apparatus. The concept behind the method is as follows. First, CO2 in the flue gas is absorbed in an alkaline solution such as NaOH.
NaOHðaqÞ þ CO2ðgÞ ! HCO3ðaqÞ þ NaþðaqÞ
ð1Þ
The absorbed CO2 exists mainly in the form of bicarbonate ions, HCO 3 , in the solution, depending on the amount absorbed. The solution is then fed into electrodialysis cells, where the carbonate ions react with protons and are converted back to CO2 gas. Fig. 1 shows a simple cell configuration for the electro-
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Fig. 1. Configuration of electrodialysis cells for CO2 recovery from carbonate solution.
dialysis process, namely the two-cell configuration [18]. In this configuration, the two types of cells, the CO2 recovery cell and the alkali regeneration cell are alternately stacked to form the electrodialysis system. Both cells are sandwiched by a bipolar membrane and a cation exchange membrane. Cation exchange membranes (CEMs) are polymeric membranes containing chemical groups with fixed negative charge and mobile cations as a counter ion. Cations can permeate through the CEM, but anions are repelled by the fixed negative charges in the CEM, and cannot permeate through the CEM. The bipolar membrane (BPM) is a membrane composed of two types of ion-exchange membranes, a cation exchange membrane and an anion exchange membrane stacked together. Anion exchange membranes (AEMs) are the membranes with fixed positive charges in the membrane by the chemical groups connected to the main polymer chains of the membrane, and mobile anions are included in the membrane as a counter ion. When applied a voltage across the BPM that is higher than the voltage necessary for water splitting, the water contained in the BPM would dissociate into a proton and a hydroxyl ion. The proton can permeate through the CEM constituting the BPM, while the hydroxyl ion can permeate through the AEM constituting the BPM. As a result, protons can be obtained at the lower potential side of the BPM and hydroxyl ions can be obtained at the higher potential side of the BPM. In the two-cell type configuration of BPM and CEM, the BPM is located at the higher potential side of the CO2 recovery cell and the CEM is located at the lower potential side. The membranes in the alkali regeneration cell are configured conversely. Other kinds of the membrane configuration can be considered. It was demonstrated that the two cell type configuration with BPM and CEM is most efficient for the CO2 recovery from sodium carbonate solutions. Alkaline carbonate solution is supplied to the CO2 recovery cell, where protons are supplied from the bipolar membrane and react with bicarbonate ions to form CO2 gas.
HCO3ðaqÞ þ HþðaqÞ ! H2 O þ CO2ðgÞ "
ð2Þ
The alkaline metal ions in the CO2 recovery cell are transported to the alkali regeneration cell through the cation exchange membrane, driven by the potential slope. Hydroxide ions are transferred from the bipolar membrane to the alkali regeneration cells and the alkaline solution is thus regenerated by the alkaline ions transferred from the CO2 recovery cell.
Naþ þ OH ! NaOH
ð3Þ
The protons that reduce the pH in the absorption solution can be supplied by the water splitting reaction in the bipolar membrane. At the same time, hydroxyl ions formed by water splitting regenerate the alkaline solution for CO2 absorption. Because CO2 gas recovery and alkaline regeneration should be performed separately, it is necessary to design the cell configuration for efficient electrodialysis. In our previous study, it was demonstrated that gaseous CO2 can be recovered immediately after the CO2 recovery cell but this was examined under limited conditions, using laboratory scale experimental apparatus. To examine process feasibility, it is necessary to examine the CO2 recovery performance and power consumption in the process under a range of conditions. It is also necessary to consider the process of CO2 absorption from flue gas by alkaline solution. This study aimed to evaluate the total cost of the process, based on the experimental results, and to identify the key parameters involved in reducing CO2 recovery costs. 2. Outline of the CO2 absorption and recovery process A flow diagram for the overall process is shown in Fig. 2. The process comprises the following three steps: (1) CO2 absorption in an absorption tower: CO2 from the flue gas stream is passed through an alkaline solution such as sodium hydroxide in an absorption tower to form carbonate solution. (2) CO2 recovery by electrodialysis: the carbonate solution is fed to the CO2 recovery cells of the electrodialysis system, where gaseous CO2 is recovered. (3) Alkaline regeneration by electrodialysis: the carbonate solution, after releasing CO2, is fed into the alkali regeneration cells in the electrodialysis system, where alkaline solution is regenerated. The regenerated alkaline solution is then reused for CO2 capture in the absorption tower. Overall, the net inlet flow to the system is the flue gas, and the net outlet flow from the system is CO2 and the off-gas. The performance of the above process would depend on various operating parameters. Fig. 3 shows a diagram of the process operation for the present system, as a function of the concentrations of sodium ions and total carbonate. Point A in Fig. 3 represents the initial con-
A. Iizuka et al. / Separation and Purification Technology 101 (2012) 49–59
51
Fig. 2. Process flow diagram.
the concentration in equilibrium with the concentration of carbonate ions, CO2 gas will be generated, the condition of which is denoted by point C. As electrodialysis proceeds, the concentrations of sodium and carbonate ions decrease along the equilibrium line from point C, downward. CO2 recovery could be completed to the extent that the carbonate ion concentration is zero. However, considering process efficiency, CO2 recovery should be limited to some extent, e.g. to point D, where the carbonate concentration is equal to CCO2,0: the initial concentration in the alkaline solution before CO2 absorption. This parameter can be expressed by the CO2 recovery ratio, defined as
CO2 recovery ratio ½% ¼
Fig. 3. Diagram of the CO2 absorption and recovery cycle.
dition of the alkaline solution before CO2 absorption. The initial concentrations of sodium ions and carbonate ions in the solution are, CNa,0 [M] and CCO2,0 [M], respectively. The solution can absorb CO2 from flue gas up to the equilibrium limit, which is represented by the dashed-dotted line in the figure, under which the maximum carbonate concentration in the solution is CCO2,eq [M], which is in equilibrium with the sodium concentration of the initial solution, CNa,0. It is not necessary and is indeed impractical to absorb CO2 at the maximum for a given sodium concentration to the equilibrium limit, CCO2,eq. The extent of CO2 absorption is expressed by the CO2 absorption ratio. In the present case, the CO2 is absorbed up to the concentration CCO2,1 [M], denoted by point B in Fig. 3.
CO2 absorption ratio ½% ¼
C CO2;1 100 C CO2;eq
ð4Þ
The solution, after CO2 absorption, is then fed to the CO2 recovery cells in the electrodialysis system, where the concentrations of sodium ions are reduced by transport through cation exchange membranes to be replaced by protons generated from the bipolar membranes. When the sodium ion concentration is reduced to
C CO2;1 C CO2;0 100 C CO2;1
ð5Þ
The solution is then fed into the alkali regeneration cell, where sodium ions are transferred from the CO2 recovery cell and hydroxyl ions are supplied from the bipolar membranes. The electrodialysis is performed to the initial conditions of the absorption solution, where the composition of the solution in the alkali regeneration cell reaches the initial conditions denoted by point A, CNa,0 and CCO2,0. The cycle of the alkaline solution is thus completed. From the above description, the key parameters for CO2 absorption and recovery by electrodialysis are the CO2 absorption ratio and the CO2 recovery ratio. In this study, the effect of the above operating conditions on the performance of the electrodialysis process was examined, based on laboratory-scale electrodialysis equipment. 3. Experimental methods Fig. 4 shows the experimental set-up for the electrodialysis. Commercially available electrodialysis equipment was used [Model CH-0, Asahi Glass Company (AGC), Tokyo, Japan]. The membrane area per cell was 0.021 m2 and the distance between membranes was fixed at 0.75 mm. A maximum of 20 cell pairs (CO2 recovery and alkali regeneration cells) was available. The anode was SUS electrode, and the cathode was platinum coated titan (Pt/Ti). The electrode solution was 0.35 M Na2SO4. The power unit for electrodialysis was PK36-11 type made by Matsusada Precision Co., Japan. Three commercially available cation exchange membranes (CEM) were used, all of which were supplied by AGC, Japan; (1) Selemion CMV type of thickness 130 lm, standard CAM, (2) Selemion CMD type of thickness 400 lm, and (3) Selemion CSO type, with thickness 100 lm, a monovalent ion selective membrane. These cation
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Fig. 4. Flow diagram for the CO2 recovery experiment.
4. Results and discussion
feed solution to the recovery cell. The current density was fixed at 96 A m2, the number of the cell-stacks was 10, and the CMV was used as the cation exchange membrane for the stack. Pure sodium bicarbonate solutions with various concentrations (0.25–1.0 M) were used as the feed solution. The feed solution to the alkaline regeneration cells was 0.1 M NaHCO3 solution. The generation of CO2 gas was observed just after the electrodialysis was started, and the CO2 recovery rate increased with time up to 20 min and then remained a steady state at about 1.5 104 mol s1 for up to 60 min after the start for all the initial sodium concentrations in the feed liquid. The sodium ion concentration in the feed solution had no remarkable effect on the steady state rate of CO2 recovery within the experimental conditions studied. The pH of the CO2 recovery cells decreased gradually with time, and the final pH was higher when the initial sodium concentration was higher. From the thermodynamic equilibrium calculations, the CO2 gas should be released when the pH of the liquid phase reduced below 7.5. In the present case, the pH was reduced immediately after the electrodialysis was started, and CO2 gas generation was observed. However, it took about 20 min to reach the steady state. This time lag is due to the fact that it takes a certain time for carbonate ions to build up in the gas–liquid separator to reach the concentration in equilibrium with the gaseous CO2 at the atmospheric pressure. Since the ratio of the CO2 recovered by the electrodialysis in 60 min
10
0.20
0.15 8 pH [-]
CO2 recovery rate [mmol/s]
exchange membranes were selected because CMV is as a standard CEM, CMD is a heterogeneous membrane with greater thickness, and CSO is monovalent ion selective membrane containing surfactant molecules on the surface. The bipolar membrane used in this study was a BP-1E type supplied by Astom, Co. (Tokyo, Japan). The electrodialysis experiments were conducted under constant current density conditions varied over the range 48–192 A m2, and the potential difference between the electrodes was measured with a voltage logger (3635-06 type, HIOKI Co., Japan). The system was operated as a semi-batch system. Alkaline solution containing captured CO2 (feed solution) was supplied to the CO2 recovery cells by a liquid pump (MDR-type, Iwaki Ltd., Japan) at a given flow rate from the 2.0 L reservoir. The exit solution was then introduced to the gas–liquid separator, which was a short, wide cylindrical polycarbonate vessel. The volumetric CO2 generation rate was measured with a volumetric gas flow meter (VP-3U, Horiba S-Tec, Japan). Solution was supplied at the bottom of the vessel and gas released from solution was removed from the top of the vessel. Gas flow rate was measured by a dry gas flow meter (DCtype, Shinagawa Corporation, Japan). CO2-lean solution was fed into the alkali regeneration cell by a liquid pump (MDR-type, Iwaki Ltd., Japan) at a given rate from a 2.0 L reservoir. The pH of each solution was monitored with a pH meter (D-52 type, Horiba, Japan). Mixtures of sodium bicarbonate and sodium hydroxide solutions at different ratios were used as the feed solution to the CO2 recovery and alkali regeneration cells. The initial ratio fed to the CO2 recovery cells was varied corresponding to the CO2 absorption ratio (Eq. (4)), ranging from 80% to 100%, with a fixed total concentration. The initial ratio fed to the alkali regeneration cells was varied corresponding to the CO2 recovery ratio (Eq. (5)), ranging from 10% to 30%, with a fixed total concentration. In addition, the total sodium concentration was varied within the range 0.25–1.0 M, corresponding to the CO2 absorption capacity for one cycle of operation. All the solutions used in the experiments were prepared from reagent grade chemicals purchased from Wako Chemical, Japan, dissolved in deionized water.
0.10 6 rate pH
0.05
0.25 M 0.50 M 1.00 M
0.00
4 0
4.1. Effect of total sodium concentration of CO2 recovery cell feed solution Fig. 5 shows the time course for the CO2 recovery rate and pH in the CO2 recovery cells at various sodium ion concentrations in the
10
20
30 40 Time [min]
50
60
Fig. 5. Effect of total sodium ion concentration on the CO2 recovery rate and pH in the CO2 recover cells. Feed solutions: CO2 recovery cells: 1.0, 0.5, and 0.25 M NaHCO3; 1.2 L min1; alkali regeneration cells: 0.1 M NaHCO3; 1.2 L min1; electrode solution, 0.35 M Na2SO4; number of cell pairs, 10; current density: 96 A m2.
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Power
Efficiency 0.25 M 0.50 M 1.0 M
0.5
5
Current Efficicency [-]
Power Consumption [kWh/kg-CO2]
The CO2 recovery rate, rCO2 can be expressed by the following equation:
1.0
10
rCO2 ¼ gI
0
10
20
30 Time [min]
40
50
60
Fig. 6. Time course of power consumption and current efficiency for the electrodialysis process under the conditions of the runs shown in Fig. 5.
was lower than 50%, and the steady state CO2 recovery would continue after 60 min. However, because the process is a batch process, the electrodialysis can be operated until the feed solution is deionized, toward this point the electric resistance would increase. The power consumption for the CO2 recovery E [J/mol-CO2] can be evaluated by the following equation:
E¼
VI ViAm ¼ rCO2 rCO2
ð7Þ
where F [C mol1] is the Faraday constant, nc [–] is the number of electrodialysis cells. gI [–] is the current efficiency. Time change of the current efficiency is shown in Fig. 6. The current efficiency increased with time and almost kept constant after reaching the steady state. The current efficiency in the steady state is summarized in Table 1. The current efficiency in the steady state was in the range of 0.52–0.80 depending on the conditions.
0.0
0
Inc inc Am ¼ gI F F
ð6Þ
where I [A] is the current, i is the current density [A m2], A is the effective membrane area [m2], V [V] is the potential difference, and rCO2 [mol s1] is the CO2 recovery rate. Time change of the power consumption with time is shown in Fig. 6. The power consumption per unit mole of recovered CO2 decreased with time and remained almost constant after about 30 min from the start time. The average power consumption per kg of CO2 was in about 1.5 kW h under the steady state. The power consumption for CO2 recovery was lower at higher concentrations of sodium ions in the feed because the resistance of the solution is lower at higher concentrations. The power consumptions after reaching the steady state under other conditions are summarized in Table 1.
4.2. Effect of CO2 absorption ratio Fig. 7 shows the time course of the CO2 recovery rate and pH in the CO2 recovery cells for various compositions of the feed solution fed to the CO2 recovery cells. The ratio of sodium bicarbonate to sodium hydroxide was varied with the total sodium concentration fixed at 1.0 M. The composition was varied corresponding to the change in CO2 absorption ratios from 80% to 100%. All other conditions were fixed: current density, 96 A m2; number of the cellstacks, 10; and the CMV cation exchange membrane was used in the stack. When pure 1.0 M sodium bicarbonate solution was fed, corresponding to a 100% CO2 absorption ratio, no time lag for the CO2 gas generation was observed. With a decreasing ratio of sodium bicarbonate, corresponding to a lower CO2 absorption ratio, the time lag for CO2 gas recovery increased. This is because the initial pH of the feed solution was lower at higher CO2 absorption ratios and the pH should be lower than 7.5 for gaseous CO2 to be released from solution, thus it would take longer for the pH of the feed solution to reach pH 7.5 when the CO2 recovery ratio is higher, and consequently the initial pH is higher. The time lag corresponds to the pathway between points B and C in the operation diagram of Fig. 3. The power consumptions at the steady state are shown in Table 1. The increasing CO2 absorption ratio decreased the power consumption.
Table 1 Power consumption and current efficiency for CO2 recovery under steady state conditions. Conditions
Power consumption [kW h-kg-CO2]
Current efficiency
Effect of sodium concentration (Fig. 5)
[Na+] = 0.25 [Na+] = 0.50 [Na+] = 1.0
1.85 1.42 1.03
0.55 0.66 0.65
Effect of CO2 absorption ratio (Fig. 7)
80% 90% 100%
1.99 1.78 1.42
0.46 0.52 0.66
Effect of CO2 recovery ratio (Fig. 8)
10% 20% 30%
1.78 1.46 1.35
0.50 0.63 0.69
Effect of current density (Fig. 9)
48 A m2 96 A m2 144 A m2 192 A m2
1.18 1.05 2.54 3.18
0.65 0.64 0.79 0.80
Effect of cell numbers (Fig. 10)
10 cells 20 cells
1.42 1.13
0.66 0.67
Effect of CEM (Fig. 11)
CMV CSO CMD
1.42 1.32 1.48
0.66 0.73 0.73
Flow rate
1.2 L min1 2.4 L min1 3.6 L min1
1.42 1.31 1.28
0.66 0.73 0.74
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10
0.8
10
0.20
rate pH
0.10 rate pH
6 80% 90% 100%
0.05
4
0.00 0
10
20
30 40 Time [min]
50
0.6
8 pH [-]
8
CO2 recovery rate [mmol/s]
0.15
pH [-]
CO2 recovery rate [mmol/s]
2
48 A/cm 2 96 A/cm 2 144 A/cm 2 192 A/cm
0.4 6 0.2
4
0.0
60
0
Fig. 7. Effect of the CO2 absorption ratio on the CO2 recovery rate and pH in the CO2 recover cells. Feed solutions: CO2 recovery cells: 1.0 M NaHCO3 for 100% absorption, 0.9 M NaHCO3 + 0.1 M NaOH for 90% absorption, 0.8 M NaHCO3 + 0.2 M NaOH for 80% absorption, 1.2 L min1, alkali regeneration cells: 0.1 M NaHCO3; electrode solution: 0.35 M Na2SO4; number of cell pairs, 10; current density, 96 A m2.
10
20
30 40 Time [min]
50
60
Fig. 9. Effect of current density on the CO2 recovery rate and pH in the CO2 recover cells. Feed solutions: CO2 recovery cells: 0.5 M NaHCO3; 1.2 L min1; alkali regeneration cells: 0.45 M NaHCO3; 1.2 L min1; electrode solution: 0.35 M Na2SO4; number of cell pairs, 10; current density, 96 A m2.
4.4. Effect of the current density 4.3. Effect of CO2 recovery ratio Fig. 8 shows the effect of the sodium bicarbonate solution concentration fed into the alkali regeneration cells on CO2 recovery rate and pH change in the CO2 recovery cells. The concentration of sodium bicarbonate corresponds to the CO2 recovery ratios given by Eq. (5). The concentration of the sodium bicarbonate solution was changed over the range 0.35–0.45 M, corresponding to CO2 recovery ratios from 10% to 30% when 0.50 M sodium bicarbonate solution was fed to the CO2 recovery cells. Other operation conditions were kept the same: current density, 96 A m2; number of the cell-stacks, 10; and a CMV cation exchange membrane was used. The flow rate of the feeds was 1.2 L min1. The observed generation rate of gaseous CO2 increased with decreasing sodium bicarbonate concentration, which corresponds to an increasing CO2 recovery rate. This is because the transport of sodium ions through the cation exchange membranes would be enhanced by a larger difference between the two cells. The power consumption for the recovery of a unit amount of CO2 is lower for the conditions of higher CO2 recovery ratio during the steady state as shown in Table 1.
4.5. Effect of the number of stacks Fig. 10 shows the time course for the CO2 recovery rate with differing numbers of cell stacks and pH change in the CO2 recovery cells. The feed to the CO2 recovery cell was 0.5 M sodium carbonate solution and that to the alkali regeneration cell was 0.1 M. The current density was fixed at 96 A m2, and the CMV cation exchange membrane was used. The CO2 recovery rate was increased with increasing the number of stacks. The power consumption per unit amount of CO2 recovered was smaller with more stacks as shown in Table 1 because there was a smaller contribution to power consumption by electrolysis at electrodes.
10
pH [-]
8
rate pH
6
10% 20% 30%
0.05
0.00
4 0
10
20
30
40
Time [min] Fig. 8. Effect of CO2 recovery ratio on the CO2 recovery rate and pH in the CO2 recover cells. Feed solutions: CO2 recovery cells: 0.5 M NaHCO3; 1.2 L min1; alkali regeneration cells: 0.35–0.45 M NaHCO3; 1.2 L min1; electrode solution: 0.35 M Na2SO4; number of cell pairs, 10; current density, 96 A m2.
0.3 8 pH [-]
0.15
0.10
10
0.4 CO2 recovery rate [mmol/s]
0.20 CO2 recovery rate [mmol/s]
Fig. 9 shows the effect of current density on the CO2 recovery rate and pH change in the CO2 recovery cells. The feed to the CO2 recovery cell was 1.0 M sodium carbonate solution and that to the alkali regeneration cell was 0.1 M. The number of cell-stacks was fixed at 10 and the CMV cation exchange membrane was used. The CO2 recovery rate increased in almost in proportional with the current density. The power consumption for recovering a unit amount of CO2 shows a minimum at 96 A m2, and increased with increasing current density after that as shown in Table 1.
0.2 6 0.1
rate pH 20 cells 10 cells
0.0
4 0
10
20
30 40 Time [min]
50
60
Fig. 10. Effect of the number of cell pairs on the CO2 recovery rate and pH in the CO2 recover cells. Feed solutions: CO2 recovery cells: 0.50 M NaHCO3; 1.2 L min1; alkali regeneration cells: 0.10 M NaHCO3; 1.2 L min1; electrode solution: 0.35 M Na2SO4; current density, 96 A m2.
A. Iizuka et al. / Separation and Purification Technology 101 (2012) 49–59
5.1. Process design and cost estimation and methods
10
0.15 8 pH [-]
CO2 recovery rate [mmol/s]
0.20
0.10 6
rate pH
0.05
CMV CSO CMD
0.00
4 0
10
20
30 40 Time [min]
50
55
60
Fig. 11. Effect of cation exchange membranes on the CO2 recovery rate and pH in the CO2 recover cells. Conditions: feed flow rates, 1.2 L min1; feed compositions: CO2 recovery cells: 0.5 M NaHCO3; alkali regeneration cells: 0.1 M NaHCO3; electrode solution: 0.35 M Na2SO4; number of cell units, 10; current, 2.0 A (current density, 96 A m2).
4.6. Effect of the cation exchange membrane Fig. 10 shows the time course of the CO2 recovery rate for various cation exchange membranes used for the electrodialysis and pH change in the CO2 recovery cells. The CO2 recovery rate was slightly higher, especially during the initial stage, when using CSO membrane, which is monovalent-ion selective. The power consumption was lowest when using the CSO membrane as shown in Table 1. However, the power consumption and current efficiency were not significantly affected by the cation exchange membrane used.
4.7. Effect of feed flow rates The feed flow rates to both cells were changed from 1.2 to 3.6 L min1. The CO2 generation rate was found to increase slightly at the increased flow rates, while the power consumption for CO2 recovery decreased with increased feed flow rate as shown in Table 1. However, the effects of flow rate on the CO2 recovery rate and power consumption were not significant. The above experimental results for the power consumption and CO2 recovery performance were comparable with the results obtained for a bipolar membrane electrodialysis system comprising a two-cell type with a different configuration [18], where potassium hydroxide rather than sodium hydroxide was used as the alkali, and an anion exchange membrane was used instead of the cation exchange membrane to separate the CO2 recovery and alkali generation cells. In this study, a cost analysis was conducted to examine the process feasibility from an economic point of view.
5. Cost estimations The cost for the CO2 recovery process of the present system was estimated using the experimental results of electrodialysis described above. First, the emission amount of CO2 to be treated with the present system was assumed to be 300 metric tons per hour (1.89 103 mol s1), which corresponds to the CO2 emission rate from a 400 MW coal-fired power plant. It is assumed that the concentration of CO2 in the flue gas is at 10% and 90% of the CO2 in the flue gas is captured in the chemical absorption process. Thus, the rate of CO2 treatment by the process, RCO2 [mol s1], is 1.89 103 0.9 = 1.70 103 mol s1. Continuous operation was assumed, 24 h per day for 350 days per year.
5.1.1. Absorption process A packing-tower type absorption tower was assumed for capturing CO2 in the flue gas with a counter flow of sodium hydroxide aqueous solution of a given concentration. The tower was made of 8-mm thickness steel (SB410 carbon steel). The tower is packed with 1-in. ceramic Raschig rings, of which the diameter is 1 in., the thickness is 1/8 in., and the surface area of the rings per unit volume of the tower, at [m2 m3]. In this study, at is assumed at 190 m2 m3, the packing density of the Raschig rings at 47,700 m3, and the void ratio of the tower, ed [m3 m3] at 0.73. The overall material balance of CO2 in the absorption tower can be expressed by the following equation,
GM LM N ¼ Q ðPB Pr Þ ¼ m ðC A2 C A1 Þ
qM
ð8Þ
where N [mol m2 s1] is the molar absorption rate of CO2 per cross-sectional area of the absorption tower; GM [mol m2 s1] is the total molar flow rates of gas, LM [mol m2 s1] is the total molar flow rate of liquid, per cross-sectional area of the absorption tower; P [Pa] is the total pressure; PT and PB are the partial pressures of CO2 at the top and bottom of the tower, respectively [Pa], CA1 and CA2 are the concentrations of the absorbent (sodium hydroxide in the present case) at the top and bottom of the tower [mol m3], respectively; qM is the molar density of the absorption liquid [mol m3]; m [–] = 2, is the stoichiometric factor for the absorption reaction. The partial pressures of CO2 at the top and the bottom of the absorption tower is PB = 0.1 atm = 1.013 104 Pa, and PT = 0.01 atm = 1.013 103 Pa, respectively assuming the 10% of CO2 in the flue gas at the atmospheric pressure to be reduced at 1% of CO2 in the ventilation gas. The sodium hydroxide concentration, CA1 correspond to that of the feed solutions to the CO2 recovery cells of the electrodialysis, which represents the CO2 absorption ratio. The concentration, CA2 corresponds to that in the alkaline regeneration cells of the electrodialysis, representing the CO2 recovery ratio. The ratio of GM/LM can be determined by Eq. (8). The gas flow rate, GM is determined based on the flooding speed, GF [kg m2 s1]. The flooding speed is the critical flow rate of gas to avoid flooding of the absorbent solution in the absorption tower. The flooding speed can be estimated based on the following empirical equation by proposed by Sawistowski [21], considering the pressure drop in the absorption tower,
" ln
G2F at g qL qG e3d
lL lW
0:2 #
" #0:25 0:5 L qG ¼ 4 G qL
ð9Þ
where qG [kg m3] and qL [kg m3] are the densities of gas and liquid, respectively; lL and lW [kg m2 s1] are the viscosities of the absorption liquid (sodium hydroxide solution) and water, respectively; G [kg m2 s1] and L [kg m2 s1] are the mass flow rates of gas and liquid, respectively; g [m s2] is the acceleration of gravity. The absorption tower is assumed to be operated under the gas flow rate, G, at 70% of the flooding speed, GF. By using the determined mass flow rate, G, the necessary diameter, DT [m], can be calculated by the following equation:
G ¼ 0:70GF ¼ MW av GM
ð10Þ
1
where MWav [kg mol ] is the average molecular weight of the gas. Once gas flow rate is determined, the diameter of the absorption tower can be determined.
DT ¼
rffiffiffiffiffiffi 4S
p
¼
rffiffiffiffiffiffiffiffiffiffiffiffiffiffi 4 RCO2 p N
ð11Þ
56
A. Iizuka et al. / Separation and Purification Technology 101 (2012) 49–59
The height Z [m] of the absorption tower can be estimated by a standard method for designing absorption towers [21,22]. For an instantaneous, irreversible reaction, the tower height can be estimated by,
þ kH GM G PB þ rC A2 GM L ln 1 s M þs LM PT þ rC A2 LM ai P 1 s GLMM
ð12Þ
where kL [m s1] and kG [mol m2 s1 Pa1] are the mass transfer coefficient based on the liquid and gas sides, respectively; H [Pa m3 mol1] is the Henry constant for CO2 solution in water, ai [m2 m3] is the effective interfacial area for liquid–gas contact per unit volume of the absorption tower. The parameter r is given by,
Ps
r¼
ð13Þ
mqM
and the surface renewal parameter, s [s1] is given by,
qM H DL
s¼
P
DG
ð14Þ
where DL [m2 s1] and DG [m2 s1] are the diffusion coefficients in the liquid and gas phases, respectively. The parameters of mass transfer coefficients and effective interfacial area were estimated by the empirical correlation equations as follows.
0:7 kG RT G lG 1=3 ¼ 5:23 ðat Dp Þ2:0 at DG at lG qG DG
ð15Þ
and,
kL
qL lL g
PL ¼
1 kG
Z¼
The power consumption for a liquid pump for the absorption tower can be estimated by,
1=3
2=3 L lL 1=2 ¼ 0:0051 ðat Dp Þ0:4 aW lL qL DL
ð16Þ
where Dp [m] is the size of the ring (=1 in. = 2.54 102 m); aw [m2 m3] is the wetted area of the ring per volume of the tower. On the other hand, ai [m2 m3] can be estimated by the following correlation,
ai ¼ 0:0406L0:455 ð1000rÞn at
ð17Þ
where
n ¼ 0:83D0:48 p
ð18Þ
Based on the absorption tower design, the cost for operation and equipment were estimated. The equipment cost comprises the costs for the tower material and the packing material. The material cost for the tower construction is the weight of the steel necessary for construction of the absorption tower. The weight of the necessary steel is calculated from the tower volume, V ¼ ð1=4ÞZD2T . The tower volume is determined by the concentrations of the alkaline solutions at the inlet and outlet of the liquid flow. The cost for the packing material is in proportional to the total tower volume because the volume fraction of the packing material is fixed. Note that the volume of the absorption tower is determined by the concentrations of sodium hydroxide at the inlet and outlet of the power, in other words, the CO2 recovery ratio and CO2 absorption ratio for a given condition. Thus, the capital cost of the absorption tower can be determined by these ratios. Other costs comprising the capital cost were not taken into account in this study. The costs of the power consumption for pumping the absorbent liquid and that for blowing the flue gas into the absorption tower were considered as the operating cost of the absorption process. The cost for power consumption comprises the costs for the liquid pump (PL [W]) for transport of the absorbent solution and the blower (PG [kW]) for flue gas injection into the absorption tower.
LSZg
ð19Þ
gP
where gP [–] is the pump efficiency, and it is assumed 0.15. The power consumption for the blower is given by,
PG ¼
Dpg N RT gP P
ð20Þ
where gB [–] is the blower efficiency assumed at 0.40, and Dp is the pressure drop [Pa] in the absorption tower, which is given by Leva’s equation, 2 bL G Dp ¼ a10qL qG Z
ð21Þ
where a and b are empirical parameters determined experimentally [5] depending on the packing condition; for 1-in. Raschig ring; a = 438, and b = 51.1. The above power consumptions are then converted to the cost by using the electricity price at USD0.12 per kW h. 5.1.2. Electrodialysis process A commercialized large-scale electrodialysis system, type CG-5 by AGC, is composed of a cell stack composed of (nc=) 2400 cell pairs with an effective membrane area at Am = 1.785 m2. It is assumed that electrodialysis of the absorbent solution is performed with equipment of the same specification (nc = 2400, Am = 1.785 m2) as that of CG-5. The current efficiency, gI, is assumed to be the same as the one obtained experimentally with a laboratory-scale electrodialysis equipment used in this study for a given electrodialysis conditions. The necessary number of electrodialysis stacks, NS, to treat the CO2 form the flue gas is given by,
NS ¼
RCO2 F ¼ RCO2 r CO2 gI inc Am
ð22Þ
where rCO2 is the CO2 regeneration rate of an electrodialysis stack with nc = 2400, Am = 1.785 m2. The capital cost of the electrodialysis stacks can be estimated from the number of stacks with a cost for one stack. The total power consumption, PED [W] for the CO2 recovery can be expressed by,
PED ¼ NS VI ¼ RCO2
FV
gI nc
ð23Þ
Thus, the operation cost can be estimated from Eq. (23) with unit cost for electricity at USD 0.12. The average current efficiency over the experimental period for each experimental condition was used for the power consumption estimation. Further assumptions made for the cost estimation are shown in Table 2. 5.2. Results of the cost estimation Fig. 12a–e shows the effect of operating parameters on the total cost for CO2 capture with the present system based on the experimental results of the electrodialysis correspond to those in Figs. 5– 10. The total cost can be broken down into four sectors; equipment cost and power cost for operation of absorption and electrodialysis, respectively. However, the contributions of the absorption process, both for equipment and operation, are almost negligible compared with the costs for electrodialysis. Thus, it would be reasonable to concentrate the discussion on the cost for electrodialysis. The effect of total sodium ions in the absorption solution is shown in Fig. 12a and was not significant. With an increasing CO2 absorption ratio, the CO2 recovery cost decreased, as shown in Fig. 12b, mainly because a higher recovery ratio can increase
57
A. Iizuka et al. / Separation and Purification Technology 101 (2012) 49–59 Table 2 Assumptions used for cost estimation.
Absorption tower Packing Electrodialysis Bipolar membrane area Cation exchange membrane area Liquid pump Blower Electric power price
Specification
Cost in USD
Years in operation
Packed tower with 8 mm-thickness steel 1-in. Raschig ring 2400 cell pairs 1.785 m2 1.785 m2 Power efficiency 15% Efficiency 40%
1000 per tonne-steel 12 kg1 2.5 million unit1 750 m2 300 m2
15 2 15 5 5
0.12 kW h1
1200 Power for ED Equipment for ED
Power for ED Equipment for ED
1000
400 Cost [USD]
Cost [USD]
800
200
600 400 200 0
0 0.0
0.5
1.0
70
1.5
+
80
90
100
CO2 absorption ratio [%]
[Na ] [M]
(b) CO2 absorption ratio
(a) Total sodium ion concentration in the absorption liquid
Power for ED Equipment for ED
Power for ED Equipment for ED
400
Cost [USD]
Cost [USD]
400
200
0
200
0 0
10
20 30 CO2 recovery ratio [%]
40
0
100 -2 Current density [A m ]
200
(d) Current density
(c) CO2 recovery ratio 400
Cost [USD]
Power for ED Equipmeny for ED
200
0 CSO
CMV
CMD
Membrane CEM
(e) Cation exchange membrane Fig. 12. Effect of operating parameters on the cost for CO2 capture and recovery.
the CO2 recovery rate and consequently reduce the power consumption for CO2 recovery by electrodialysis. The number of operating cycles for a given amount of CO2 can be reduced with a higher CO2 absorption ratio and, consequently, the cost can also be reduced. The cost for CO2 recovery decreased with increasing CO2
recovery ratio but the effect was not significant, as shown in Fig. 12c. However, the experimentally tested values of the CO2 recovery ratio were 30% at the highest due to the experimental limit. However, the cost can be reduced when the CO2 recovery ratio is higher. The CO2 recovery cost was more sensitive to the CO2
58
A. Iizuka et al. / Separation and Purification Technology 101 (2012) 49–59
Table 3 Optimal operating conditions.
Table 4 Breakdown of the CO2 recovery costs at the optimal condition.
Electrodialysis Membrane area Cell volume Feed flow into the CO2 recovery cell Feed flow into the alkali regeneration cell Cation exchange membrane Current density Absorption tower volume
9.2 105 m2 6.9 102 m3 1.0 M NaHCO3; 1.2 L/min 0.42 M NaHCO3; 1.2 L/min CSO 240 A m2 50 m3
absorption ratio, as shown in Fig. 12b than to the CO2 recovery ratio. The effect of current density on the CO2 recovery cost is shown in Fig. 12d. The minimum cost was achieved at a current density of 96 A m2. The equipment cost was a decreasing function of the current density and vice versa for the cost of power consumption. Thus, the minimum cost was realized at an intermediate value of current density at 96 A m2. Fig. 12e shows the CO2 recovery cost for various cation exchange membranes, tested in this study. The total cost was lowest for the CSO membrane but the difference was not significant. The CO2 recovery process was then optimized by changing all the parameters related to the experimental results with extrapolations. The optimum operating condition is shown in Table 3, which achieves the minimum cost of US$180 per ton CO2. The cost breakdown is shown in Table 4. The cost of electric power accounts about 60% of the total cost and that of membranes accounts 30% of the total cost. The minimum cost is higher than that for conventional CO2 recovery processes based on amine absorption/thermal desorption, which is in the range lower than US$100 [23], based on the usage of the best performance amine such as KS-1. About 60%
Absorption Operation Equipment
0.068 0.005
Electrodialysis Electric power Equipment Membrane
110 11 56
Total
180 USD/t-CO2
of the total cost of the amine-based absorption process can steam supply for recovery of CO2 from the solvent. Possibilities of the cost reduction were examined by sensitivity analysis of the total cost. The electric power cost can be reduced by improving the current efficiency and the electric power cost itself. The equipment cost should be significantly reduced by reducing the membrane prices. The costs for membranes should be significantly reduced under a larger scale production of ion-exchange membranes and the bipolar membrane because the material costs for the membrane fabrication may not be high; note the above costs for the membranes are based on the prices for a very small-scale supply for laboratory experimental uses. The sensitivity analysis was conducted by choosing the current density, current efficiency, and prices of the membrane and electric power as the independent variables. Fig. 13a shows the total cost of the CO2 recovery process and its breakdown for equipment, power as a function of the current density, with a fixed current density at 75%, and prices of membranes and electric power are the same as the base case shown in Fig. 12d.
600 150 Total Equipment for ED Power for ED
100 Cost [USD]
Cost [USD]
400
200
Total Equipment for ED Power for ED
50
0
0 0
200
400
0
600
200
400
600
2
-2
Current density [A m ]
Current density [A/m ]
(a) Base case
(b) Membrane prices: 1/10, electricity: 0.12 USD/kWh, current efficiency: 90%
Total Equpment for ED POwer for ED
Cost [USD]
100
50
0 0
200
400
600
2
Current density [A/m ]
(c) Membrane prices: 1/10, electricity price: 0.06 USD/kWh, current efficiency: 90%. Fig. 13. Sensitivity analysis of the cost for electrodialysis: (a) Base case: current efficiency, 75%; membrane costs: as they are; electric power cost USD 0.12 per 1 kW h, (b) current efficiency, 90%; membrane cost: 1/10; electric power cost: USD 0.06 per 1 kW h, and (c) current efficiency, 90%; membrane cost: 1/10, electric power cost: USD 0.06 per 1 kW h.
A. Iizuka et al. / Separation and Purification Technology 101 (2012) 49–59
Fig. 13b shows the total cost and its breakdown of the CO2 recovery process as a function of the current density with the higher current efficiency at 90%, and lower membrane costs at 10% of the base case. The total cost could be significantly reduced, and the minimum total cost is about US$100 per metric ton-CO2, which is almost half that of the present case and comparable with the conventional processes. In Fig. 13c the estimated cost of the process was shown when the electric power price is halved to US$ 0.06 per kW h, and the membrane costs are 1/10 and the current density is 90%. The minimum cost is about US$70, which is competitive with the conventional process of amine chemical absorption. The total cost can be further reduced by enlarging the lifetime of the membranes (5 years for the base case). It should be noted that all the estimated costs are based on the experimental results carried out with a semi-batch system in this study. The power consumption for the CO2 recovery process was significantly high at the initial stage before reaching the steady state. The CO2 recovery performance can be avoided when the system is operated in a continuous mode. A significant cost reduction can be anticipated when the present system is operated in a continuous mode. In the future, more cost reduction might be possible by reducing membrane costs through mass production so that the present CO2 recovery process may be competitive with the conventional method. Another advantage of the present process is that it can be operated near room temperature and at atmospheric pressure conditions. This enables rapid startup and shutdown of the operation and high process safety. 6. Conclusions From the above experimental results and cost estimation, the optimal conditions for the recovery of CO2 based on chemical absorption with an alkaline solution, combined with electrodialysis using bipolar membranes, was determined. The sensitivity analysis showed a significant cost reduction can be achieved when the membrane cost is reduced and the current efficiency is improved. References [1] B. Metz, O. Davidson, H. de Coninck, M. Loos, L. Meyer (Eds.), IPCC, Carbon Dioxide Capture and Storage, Cambridge University Press, UK, 2005. [2] A. Yamasaki, Journal of Chemical Engineering of Japan 36 (2003) 361–375. [3] R. Barchas, R. Davis, The Kerr-McGee/ABB Lummus Crest technology for the recovery of CO2 from stack gases, Energy Conversion and Management 33 (5– 8) (1992) 333–340.
59
[4] D.S. Arnold, D.A. Barrett, R.H. Isom, CO2 can be produced from flue gas, Oil and Gas Journal (November) (1982) 130–136. [5] M.T. Sander, C.L. Mariz, The Fluor DanielÒ Econamine™ FG process: past experience and present day focus, Energy Conversion Management 33 (5–8) (1992) 341–348. [6] D.G. Chapel, C.L. Mariz, J. Ernest, Recovery of CO2 from flue gases: commercial trends, in: Paper No. 340 at the Annual Meeting of the Canadian Society of Chemical Engineering, Saskatoon, Canada, October, 1999. [7] T. Mimura, S. Satsumi, M. Iijima, S. Mitsuoka, Development on energy saving technology for flue gas carbon dioxide recovery by the chemical absorption method and steam system in power plant, in: P. Riemer, B. Eliasson, A. Wokaun (Eds.), Greenhouse Gas Control Technologies, Elsevier Science Ltd., United Kingdom, 1999, pp. 71–76. [8] T. Mimura, T. Nojo, M. Iijima, T. Yoshiyama, H. Tanaka, Recent developments in flue gas CO2 recovery technology, in: J. Gale, Y. Kaya (Eds.), Greenhouse Gas Control Technologies, Proceedings of the 6th International Conference on Greenhouse Gas Control Technologies (GHGT-6), 1–4 October 2002, Kyoto, Japan, Elsevier Science Ltd., Oxford, UK, 2003. [9] M. Ishibashi, K. Otake, S. Kanamori, A. Yasutake, Study on CO2 removal technology from flue gas of thermal power plant by physical adsorption method, in: P. Riemer, B. Eliasson, A. Wokaun (Eds.), Greenhouse Gas Control Technologies, Elsevier Science, Ltd., United Kingdom, 1999, pp. 95–100. [10] T. Yokoyama, Japanese R&D on CO2 capture, in: J. Gale, Y. Kaya (Eds.), Greenhouse Gas Control Technologies, Proc. of the 6th International Conference on Greenhouse Gas Control Technologies (GHGT-6), 1–4 October 2002, Kyoto, Japan, Elsevier Science Ltd., Oxford, UK, 2003, pp. 13–18. [11] J.R. Hufton, S. Mayorga, S. Sircar, Sorption enhanced reaction process for hydrogen production, AIChE Journal 45 (1999) 248–254. [12] K. Nakagawa, T. Ohashi, A novel method of CO2 capture from high temperature gases, Journal of the Electrochemical Society 145 (4) (1998) 1344–1346. [13] H. Herzog, D. Golomb, S. Zemba, Feasibility, modeling and economics of sequestering power plant CO2 emissions in the deep ocean, Environmental Progress 10 (1) (1991) 64–74. [14] J.P. Van der Sluijs, C.A. Hendriks, K. Blok, Feasibility of polymer membranes for carbon dioxide recovery from flue gases, Energy Conversion Management 33 (5–8) (1992) 429–436. [15] P.H.M. Feron, Membranes for carbon dioxide recovery from power plants, in: J. Paul, C.M. Pradier (Eds.), Carbon Dioxide Chemistry: Environmental Issues, The Royal Society of Chemistry, Cambridge, United Kingdom, 1994, pp. 236–249. [16] P. Luis, T. Van Gerven, B. Van der Bruggen, Recent developments in membranebased technologies for CO2 capture, Progress in Energy and Combustion Science 38 (2012) 419–448. [17] P. Luis, B. Van der Bruggen, T. Van Gerven, Non-dispersive absorption for CO2 capture: from the lab to the industry, Journal of Chemical Technology and Biotechnology 86 (2011) 769–775. [18] H. Nagasawa, A. Iizuka, K. Kumagai, A. Yamasaki, Y. Yanagisawa, AIChE Journal 55 (2009) 3286–3293. [19] M.D. Eisaman, L. Alvarado, D. Larner, P. Wang, B. Garg, K.A. Littau, Energy and Environmental Science 4 (2011) 1319–1328. [20] M.D. Eisaman, L. Alvarado, D. Larner, P. Wang, K.A. Littau, Energy and Environmental Science 4 (2011) 4031–4037. [21] H. Sawistowski, Flooding velocities in packed columns operating at reduced pressures, Chemical Engineering Science 6 (1957) 138–140. [22] Society of Chemical Engineers, Japan, Kagaku Kogaku Binran (Chemical Engineering Handbook), 7th ed., Maruzen, Tokyo, Japan (in Japanese). [23] M. Inui, Ph.D. Thesis, Feasibility Study on CO2 Subseabed Storage in the Form of Gas Hydrate, The University of Tokyo, 2010.
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Separation and Purification Technology journal homepage: www.elsevier.com/locate/seppur
Carbon dioxide recovery from carbonate solutions using bipolar membrane electrodialysis Atsushi Iizuka a, Kana Hashimoto b, Hiroki Nagasawa c, Kazukiyo Kumagai c, Yukio Yanagisawa c, Akihiro Yamasaki d,⇑ a
Research Center for Sustainable Science and Engineering, Institute of Multidisciplinary Research for Advanced Materials, Tohoku University, 2-1-1 Katahira, Sendai, Miyagi 980-8577, Japan Department of Chemical System Engineering, Faculty of Engineering, The University of Tokyo, 7-3-1 Hongo Bunkyo-ku, Tokyo 113-8656, Japan c Department of Environment Systems, Graduate School of Frontier Sciences, The University of Tokyo, 5-1-5 Kashiwanoha Kashiwa, Chiba 277-8563, Japan d Department of Materials and Life Science, Faculty of Science and Technology, Seikei University, 3-3-1 Kichijoji-Kitamachi, Musashino, Tokyo 180-8633, Japan b
a r t i c l e
i n f o
Article history: Received 11 June 2012 Received in revised form 12 September 2012 Accepted 12 September 2012 Available online 23 September 2012 Keywords: CO2 recovery Bipolar membrane electrodialysis Power consumption Cost estimation Global warming Carbon capture and storage
a b s t r a c t Process design and cost estimation were conducted for CO2 recovery by liquid absorption with alkaline solution, coupled with bipolar membrane electrodialysis for CO2 gas separation and alkaline solution regeneration. The electrodialysis performances and power consumption for CO2 recovery and alkaline solution regeneration were examined using laboratory-scale electrodialysis equipment under various conditions; degree of CO2 absorption, CO2 recovery ratio, alkaline concentration, and type of cation exchange membrane. The total cost of CO2 recovery was estimated for the treatment of CO2 emitted from a 400 MW coal-fired thermal power plant. The minimum cost for the optimum condition was about US$180 per ton-CO2 removed, which is higher than that for the conventional process using amine absorption and thermal desorption. However, a sensitivity analysis indicates that the overall cost could be significantly reduced if the cost of membranes was reduced. Ó 2012 Elsevier B.V. All rights reserved.
1. Introduction Carbon capture and storage (CCS) has been widely recognized as an effective countermeasure to greenhouse gas emissions that mitigates global warming [1,2]. In the CCS process, carbon dioxide emitted from point sources would be captured and transported to storage sites, where liquid CO2 would be injected into either underground or the ocean waters. The main concern regarding CCS implementation is the need to reduce cost and power consumption for capturing CO2 from flue gas. More than 80% of the total power consumption and cost for CCS can be attributed to those for CO2 capture process. There are several methods for capturing and recovering CO2 from point sources, e.g., chemical absorption [3–8], physical adsorption including pressure swing adsorption (PSA) [9,10] and temperature swing adsorption (TSA) [11,12], membrane separation [13–17]. Among these methods, chemical absorption method using amines has been considered to be the most promising method from the viewpoints of both cost and energy penalty [1]. Aqueous solutions of mono ethanol amine (MEA) and its similar substances are used for capturing ⇑ Corresponding author. Tel.: +81 422373887. E-mail address: [email protected] (A. Yamasaki). 1383-5866/$ - see front matter Ó 2012 Elsevier B.V. All rights reserved. http://dx.doi.org/10.1016/j.seppur.2012.09.016
CO2 in the flue gas emitted from point sources. These amines form CO2-amine compounds at the lower temperature conditions. It is necessary to increase the temperature of the solution to recover CO2. The heating process would require a lot of heat; almost all the power consumption for the CO2 capture can be attributed to this regeneration process, which in turn increases the cost for CCS [1]. Electrochemical methods have been proposed for recovery of CO2 from chemically absorbed solutions [18–20]. Electrodialysis using bipolar membranes is utilized for these recovering methods of CO2 from solutions. In a previous paper [18], we proposed a method for recovering CO2 from carbonate solutions and demonstrated its recovery performances from sodium carbonate with a laboratory-scale experimental apparatus. The concept behind the method is as follows. First, CO2 in the flue gas is absorbed in an alkaline solution such as NaOH.
NaOHðaqÞ þ CO2ðgÞ ! HCO3ðaqÞ þ NaþðaqÞ
ð1Þ
The absorbed CO2 exists mainly in the form of bicarbonate ions, HCO 3 , in the solution, depending on the amount absorbed. The solution is then fed into electrodialysis cells, where the carbonate ions react with protons and are converted back to CO2 gas. Fig. 1 shows a simple cell configuration for the electro-
50
A. Iizuka et al. / Separation and Purification Technology 101 (2012) 49–59
Fig. 1. Configuration of electrodialysis cells for CO2 recovery from carbonate solution.
dialysis process, namely the two-cell configuration [18]. In this configuration, the two types of cells, the CO2 recovery cell and the alkali regeneration cell are alternately stacked to form the electrodialysis system. Both cells are sandwiched by a bipolar membrane and a cation exchange membrane. Cation exchange membranes (CEMs) are polymeric membranes containing chemical groups with fixed negative charge and mobile cations as a counter ion. Cations can permeate through the CEM, but anions are repelled by the fixed negative charges in the CEM, and cannot permeate through the CEM. The bipolar membrane (BPM) is a membrane composed of two types of ion-exchange membranes, a cation exchange membrane and an anion exchange membrane stacked together. Anion exchange membranes (AEMs) are the membranes with fixed positive charges in the membrane by the chemical groups connected to the main polymer chains of the membrane, and mobile anions are included in the membrane as a counter ion. When applied a voltage across the BPM that is higher than the voltage necessary for water splitting, the water contained in the BPM would dissociate into a proton and a hydroxyl ion. The proton can permeate through the CEM constituting the BPM, while the hydroxyl ion can permeate through the AEM constituting the BPM. As a result, protons can be obtained at the lower potential side of the BPM and hydroxyl ions can be obtained at the higher potential side of the BPM. In the two-cell type configuration of BPM and CEM, the BPM is located at the higher potential side of the CO2 recovery cell and the CEM is located at the lower potential side. The membranes in the alkali regeneration cell are configured conversely. Other kinds of the membrane configuration can be considered. It was demonstrated that the two cell type configuration with BPM and CEM is most efficient for the CO2 recovery from sodium carbonate solutions. Alkaline carbonate solution is supplied to the CO2 recovery cell, where protons are supplied from the bipolar membrane and react with bicarbonate ions to form CO2 gas.
HCO3ðaqÞ þ HþðaqÞ ! H2 O þ CO2ðgÞ "
ð2Þ
The alkaline metal ions in the CO2 recovery cell are transported to the alkali regeneration cell through the cation exchange membrane, driven by the potential slope. Hydroxide ions are transferred from the bipolar membrane to the alkali regeneration cells and the alkaline solution is thus regenerated by the alkaline ions transferred from the CO2 recovery cell.
Naþ þ OH ! NaOH
ð3Þ
The protons that reduce the pH in the absorption solution can be supplied by the water splitting reaction in the bipolar membrane. At the same time, hydroxyl ions formed by water splitting regenerate the alkaline solution for CO2 absorption. Because CO2 gas recovery and alkaline regeneration should be performed separately, it is necessary to design the cell configuration for efficient electrodialysis. In our previous study, it was demonstrated that gaseous CO2 can be recovered immediately after the CO2 recovery cell but this was examined under limited conditions, using laboratory scale experimental apparatus. To examine process feasibility, it is necessary to examine the CO2 recovery performance and power consumption in the process under a range of conditions. It is also necessary to consider the process of CO2 absorption from flue gas by alkaline solution. This study aimed to evaluate the total cost of the process, based on the experimental results, and to identify the key parameters involved in reducing CO2 recovery costs. 2. Outline of the CO2 absorption and recovery process A flow diagram for the overall process is shown in Fig. 2. The process comprises the following three steps: (1) CO2 absorption in an absorption tower: CO2 from the flue gas stream is passed through an alkaline solution such as sodium hydroxide in an absorption tower to form carbonate solution. (2) CO2 recovery by electrodialysis: the carbonate solution is fed to the CO2 recovery cells of the electrodialysis system, where gaseous CO2 is recovered. (3) Alkaline regeneration by electrodialysis: the carbonate solution, after releasing CO2, is fed into the alkali regeneration cells in the electrodialysis system, where alkaline solution is regenerated. The regenerated alkaline solution is then reused for CO2 capture in the absorption tower. Overall, the net inlet flow to the system is the flue gas, and the net outlet flow from the system is CO2 and the off-gas. The performance of the above process would depend on various operating parameters. Fig. 3 shows a diagram of the process operation for the present system, as a function of the concentrations of sodium ions and total carbonate. Point A in Fig. 3 represents the initial con-
A. Iizuka et al. / Separation and Purification Technology 101 (2012) 49–59
51
Fig. 2. Process flow diagram.
the concentration in equilibrium with the concentration of carbonate ions, CO2 gas will be generated, the condition of which is denoted by point C. As electrodialysis proceeds, the concentrations of sodium and carbonate ions decrease along the equilibrium line from point C, downward. CO2 recovery could be completed to the extent that the carbonate ion concentration is zero. However, considering process efficiency, CO2 recovery should be limited to some extent, e.g. to point D, where the carbonate concentration is equal to CCO2,0: the initial concentration in the alkaline solution before CO2 absorption. This parameter can be expressed by the CO2 recovery ratio, defined as
CO2 recovery ratio ½% ¼
Fig. 3. Diagram of the CO2 absorption and recovery cycle.
dition of the alkaline solution before CO2 absorption. The initial concentrations of sodium ions and carbonate ions in the solution are, CNa,0 [M] and CCO2,0 [M], respectively. The solution can absorb CO2 from flue gas up to the equilibrium limit, which is represented by the dashed-dotted line in the figure, under which the maximum carbonate concentration in the solution is CCO2,eq [M], which is in equilibrium with the sodium concentration of the initial solution, CNa,0. It is not necessary and is indeed impractical to absorb CO2 at the maximum for a given sodium concentration to the equilibrium limit, CCO2,eq. The extent of CO2 absorption is expressed by the CO2 absorption ratio. In the present case, the CO2 is absorbed up to the concentration CCO2,1 [M], denoted by point B in Fig. 3.
CO2 absorption ratio ½% ¼
C CO2;1 100 C CO2;eq
ð4Þ
The solution, after CO2 absorption, is then fed to the CO2 recovery cells in the electrodialysis system, where the concentrations of sodium ions are reduced by transport through cation exchange membranes to be replaced by protons generated from the bipolar membranes. When the sodium ion concentration is reduced to
C CO2;1 C CO2;0 100 C CO2;1
ð5Þ
The solution is then fed into the alkali regeneration cell, where sodium ions are transferred from the CO2 recovery cell and hydroxyl ions are supplied from the bipolar membranes. The electrodialysis is performed to the initial conditions of the absorption solution, where the composition of the solution in the alkali regeneration cell reaches the initial conditions denoted by point A, CNa,0 and CCO2,0. The cycle of the alkaline solution is thus completed. From the above description, the key parameters for CO2 absorption and recovery by electrodialysis are the CO2 absorption ratio and the CO2 recovery ratio. In this study, the effect of the above operating conditions on the performance of the electrodialysis process was examined, based on laboratory-scale electrodialysis equipment. 3. Experimental methods Fig. 4 shows the experimental set-up for the electrodialysis. Commercially available electrodialysis equipment was used [Model CH-0, Asahi Glass Company (AGC), Tokyo, Japan]. The membrane area per cell was 0.021 m2 and the distance between membranes was fixed at 0.75 mm. A maximum of 20 cell pairs (CO2 recovery and alkali regeneration cells) was available. The anode was SUS electrode, and the cathode was platinum coated titan (Pt/Ti). The electrode solution was 0.35 M Na2SO4. The power unit for electrodialysis was PK36-11 type made by Matsusada Precision Co., Japan. Three commercially available cation exchange membranes (CEM) were used, all of which were supplied by AGC, Japan; (1) Selemion CMV type of thickness 130 lm, standard CAM, (2) Selemion CMD type of thickness 400 lm, and (3) Selemion CSO type, with thickness 100 lm, a monovalent ion selective membrane. These cation
52
A. Iizuka et al. / Separation and Purification Technology 101 (2012) 49–59
Fig. 4. Flow diagram for the CO2 recovery experiment.
4. Results and discussion
feed solution to the recovery cell. The current density was fixed at 96 A m2, the number of the cell-stacks was 10, and the CMV was used as the cation exchange membrane for the stack. Pure sodium bicarbonate solutions with various concentrations (0.25–1.0 M) were used as the feed solution. The feed solution to the alkaline regeneration cells was 0.1 M NaHCO3 solution. The generation of CO2 gas was observed just after the electrodialysis was started, and the CO2 recovery rate increased with time up to 20 min and then remained a steady state at about 1.5 104 mol s1 for up to 60 min after the start for all the initial sodium concentrations in the feed liquid. The sodium ion concentration in the feed solution had no remarkable effect on the steady state rate of CO2 recovery within the experimental conditions studied. The pH of the CO2 recovery cells decreased gradually with time, and the final pH was higher when the initial sodium concentration was higher. From the thermodynamic equilibrium calculations, the CO2 gas should be released when the pH of the liquid phase reduced below 7.5. In the present case, the pH was reduced immediately after the electrodialysis was started, and CO2 gas generation was observed. However, it took about 20 min to reach the steady state. This time lag is due to the fact that it takes a certain time for carbonate ions to build up in the gas–liquid separator to reach the concentration in equilibrium with the gaseous CO2 at the atmospheric pressure. Since the ratio of the CO2 recovered by the electrodialysis in 60 min
10
0.20
0.15 8 pH [-]
CO2 recovery rate [mmol/s]
exchange membranes were selected because CMV is as a standard CEM, CMD is a heterogeneous membrane with greater thickness, and CSO is monovalent ion selective membrane containing surfactant molecules on the surface. The bipolar membrane used in this study was a BP-1E type supplied by Astom, Co. (Tokyo, Japan). The electrodialysis experiments were conducted under constant current density conditions varied over the range 48–192 A m2, and the potential difference between the electrodes was measured with a voltage logger (3635-06 type, HIOKI Co., Japan). The system was operated as a semi-batch system. Alkaline solution containing captured CO2 (feed solution) was supplied to the CO2 recovery cells by a liquid pump (MDR-type, Iwaki Ltd., Japan) at a given flow rate from the 2.0 L reservoir. The exit solution was then introduced to the gas–liquid separator, which was a short, wide cylindrical polycarbonate vessel. The volumetric CO2 generation rate was measured with a volumetric gas flow meter (VP-3U, Horiba S-Tec, Japan). Solution was supplied at the bottom of the vessel and gas released from solution was removed from the top of the vessel. Gas flow rate was measured by a dry gas flow meter (DCtype, Shinagawa Corporation, Japan). CO2-lean solution was fed into the alkali regeneration cell by a liquid pump (MDR-type, Iwaki Ltd., Japan) at a given rate from a 2.0 L reservoir. The pH of each solution was monitored with a pH meter (D-52 type, Horiba, Japan). Mixtures of sodium bicarbonate and sodium hydroxide solutions at different ratios were used as the feed solution to the CO2 recovery and alkali regeneration cells. The initial ratio fed to the CO2 recovery cells was varied corresponding to the CO2 absorption ratio (Eq. (4)), ranging from 80% to 100%, with a fixed total concentration. The initial ratio fed to the alkali regeneration cells was varied corresponding to the CO2 recovery ratio (Eq. (5)), ranging from 10% to 30%, with a fixed total concentration. In addition, the total sodium concentration was varied within the range 0.25–1.0 M, corresponding to the CO2 absorption capacity for one cycle of operation. All the solutions used in the experiments were prepared from reagent grade chemicals purchased from Wako Chemical, Japan, dissolved in deionized water.
0.10 6 rate pH
0.05
0.25 M 0.50 M 1.00 M
0.00
4 0
4.1. Effect of total sodium concentration of CO2 recovery cell feed solution Fig. 5 shows the time course for the CO2 recovery rate and pH in the CO2 recovery cells at various sodium ion concentrations in the
10
20
30 40 Time [min]
50
60
Fig. 5. Effect of total sodium ion concentration on the CO2 recovery rate and pH in the CO2 recover cells. Feed solutions: CO2 recovery cells: 1.0, 0.5, and 0.25 M NaHCO3; 1.2 L min1; alkali regeneration cells: 0.1 M NaHCO3; 1.2 L min1; electrode solution, 0.35 M Na2SO4; number of cell pairs, 10; current density: 96 A m2.
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A. Iizuka et al. / Separation and Purification Technology 101 (2012) 49–59
Power
Efficiency 0.25 M 0.50 M 1.0 M
0.5
5
Current Efficicency [-]
Power Consumption [kWh/kg-CO2]
The CO2 recovery rate, rCO2 can be expressed by the following equation:
1.0
10
rCO2 ¼ gI
0
10
20
30 Time [min]
40
50
60
Fig. 6. Time course of power consumption and current efficiency for the electrodialysis process under the conditions of the runs shown in Fig. 5.
was lower than 50%, and the steady state CO2 recovery would continue after 60 min. However, because the process is a batch process, the electrodialysis can be operated until the feed solution is deionized, toward this point the electric resistance would increase. The power consumption for the CO2 recovery E [J/mol-CO2] can be evaluated by the following equation:
E¼
VI ViAm ¼ rCO2 rCO2
ð7Þ
where F [C mol1] is the Faraday constant, nc [–] is the number of electrodialysis cells. gI [–] is the current efficiency. Time change of the current efficiency is shown in Fig. 6. The current efficiency increased with time and almost kept constant after reaching the steady state. The current efficiency in the steady state is summarized in Table 1. The current efficiency in the steady state was in the range of 0.52–0.80 depending on the conditions.
0.0
0
Inc inc Am ¼ gI F F
ð6Þ
where I [A] is the current, i is the current density [A m2], A is the effective membrane area [m2], V [V] is the potential difference, and rCO2 [mol s1] is the CO2 recovery rate. Time change of the power consumption with time is shown in Fig. 6. The power consumption per unit mole of recovered CO2 decreased with time and remained almost constant after about 30 min from the start time. The average power consumption per kg of CO2 was in about 1.5 kW h under the steady state. The power consumption for CO2 recovery was lower at higher concentrations of sodium ions in the feed because the resistance of the solution is lower at higher concentrations. The power consumptions after reaching the steady state under other conditions are summarized in Table 1.
4.2. Effect of CO2 absorption ratio Fig. 7 shows the time course of the CO2 recovery rate and pH in the CO2 recovery cells for various compositions of the feed solution fed to the CO2 recovery cells. The ratio of sodium bicarbonate to sodium hydroxide was varied with the total sodium concentration fixed at 1.0 M. The composition was varied corresponding to the change in CO2 absorption ratios from 80% to 100%. All other conditions were fixed: current density, 96 A m2; number of the cellstacks, 10; and the CMV cation exchange membrane was used in the stack. When pure 1.0 M sodium bicarbonate solution was fed, corresponding to a 100% CO2 absorption ratio, no time lag for the CO2 gas generation was observed. With a decreasing ratio of sodium bicarbonate, corresponding to a lower CO2 absorption ratio, the time lag for CO2 gas recovery increased. This is because the initial pH of the feed solution was lower at higher CO2 absorption ratios and the pH should be lower than 7.5 for gaseous CO2 to be released from solution, thus it would take longer for the pH of the feed solution to reach pH 7.5 when the CO2 recovery ratio is higher, and consequently the initial pH is higher. The time lag corresponds to the pathway between points B and C in the operation diagram of Fig. 3. The power consumptions at the steady state are shown in Table 1. The increasing CO2 absorption ratio decreased the power consumption.
Table 1 Power consumption and current efficiency for CO2 recovery under steady state conditions. Conditions
Power consumption [kW h-kg-CO2]
Current efficiency
Effect of sodium concentration (Fig. 5)
[Na+] = 0.25 [Na+] = 0.50 [Na+] = 1.0
1.85 1.42 1.03
0.55 0.66 0.65
Effect of CO2 absorption ratio (Fig. 7)
80% 90% 100%
1.99 1.78 1.42
0.46 0.52 0.66
Effect of CO2 recovery ratio (Fig. 8)
10% 20% 30%
1.78 1.46 1.35
0.50 0.63 0.69
Effect of current density (Fig. 9)
48 A m2 96 A m2 144 A m2 192 A m2
1.18 1.05 2.54 3.18
0.65 0.64 0.79 0.80
Effect of cell numbers (Fig. 10)
10 cells 20 cells
1.42 1.13
0.66 0.67
Effect of CEM (Fig. 11)
CMV CSO CMD
1.42 1.32 1.48
0.66 0.73 0.73
Flow rate
1.2 L min1 2.4 L min1 3.6 L min1
1.42 1.31 1.28
0.66 0.73 0.74
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A. Iizuka et al. / Separation and Purification Technology 101 (2012) 49–59
10
0.8
10
0.20
rate pH
0.10 rate pH
6 80% 90% 100%
0.05
4
0.00 0
10
20
30 40 Time [min]
50
0.6
8 pH [-]
8
CO2 recovery rate [mmol/s]
0.15
pH [-]
CO2 recovery rate [mmol/s]
2
48 A/cm 2 96 A/cm 2 144 A/cm 2 192 A/cm
0.4 6 0.2
4
0.0
60
0
Fig. 7. Effect of the CO2 absorption ratio on the CO2 recovery rate and pH in the CO2 recover cells. Feed solutions: CO2 recovery cells: 1.0 M NaHCO3 for 100% absorption, 0.9 M NaHCO3 + 0.1 M NaOH for 90% absorption, 0.8 M NaHCO3 + 0.2 M NaOH for 80% absorption, 1.2 L min1, alkali regeneration cells: 0.1 M NaHCO3; electrode solution: 0.35 M Na2SO4; number of cell pairs, 10; current density, 96 A m2.
10
20
30 40 Time [min]
50
60
Fig. 9. Effect of current density on the CO2 recovery rate and pH in the CO2 recover cells. Feed solutions: CO2 recovery cells: 0.5 M NaHCO3; 1.2 L min1; alkali regeneration cells: 0.45 M NaHCO3; 1.2 L min1; electrode solution: 0.35 M Na2SO4; number of cell pairs, 10; current density, 96 A m2.
4.4. Effect of the current density 4.3. Effect of CO2 recovery ratio Fig. 8 shows the effect of the sodium bicarbonate solution concentration fed into the alkali regeneration cells on CO2 recovery rate and pH change in the CO2 recovery cells. The concentration of sodium bicarbonate corresponds to the CO2 recovery ratios given by Eq. (5). The concentration of the sodium bicarbonate solution was changed over the range 0.35–0.45 M, corresponding to CO2 recovery ratios from 10% to 30% when 0.50 M sodium bicarbonate solution was fed to the CO2 recovery cells. Other operation conditions were kept the same: current density, 96 A m2; number of the cell-stacks, 10; and a CMV cation exchange membrane was used. The flow rate of the feeds was 1.2 L min1. The observed generation rate of gaseous CO2 increased with decreasing sodium bicarbonate concentration, which corresponds to an increasing CO2 recovery rate. This is because the transport of sodium ions through the cation exchange membranes would be enhanced by a larger difference between the two cells. The power consumption for the recovery of a unit amount of CO2 is lower for the conditions of higher CO2 recovery ratio during the steady state as shown in Table 1.
4.5. Effect of the number of stacks Fig. 10 shows the time course for the CO2 recovery rate with differing numbers of cell stacks and pH change in the CO2 recovery cells. The feed to the CO2 recovery cell was 0.5 M sodium carbonate solution and that to the alkali regeneration cell was 0.1 M. The current density was fixed at 96 A m2, and the CMV cation exchange membrane was used. The CO2 recovery rate was increased with increasing the number of stacks. The power consumption per unit amount of CO2 recovered was smaller with more stacks as shown in Table 1 because there was a smaller contribution to power consumption by electrolysis at electrodes.
10
pH [-]
8
rate pH
6
10% 20% 30%
0.05
0.00
4 0
10
20
30
40
Time [min] Fig. 8. Effect of CO2 recovery ratio on the CO2 recovery rate and pH in the CO2 recover cells. Feed solutions: CO2 recovery cells: 0.5 M NaHCO3; 1.2 L min1; alkali regeneration cells: 0.35–0.45 M NaHCO3; 1.2 L min1; electrode solution: 0.35 M Na2SO4; number of cell pairs, 10; current density, 96 A m2.
0.3 8 pH [-]
0.15
0.10
10
0.4 CO2 recovery rate [mmol/s]
0.20 CO2 recovery rate [mmol/s]
Fig. 9 shows the effect of current density on the CO2 recovery rate and pH change in the CO2 recovery cells. The feed to the CO2 recovery cell was 1.0 M sodium carbonate solution and that to the alkali regeneration cell was 0.1 M. The number of cell-stacks was fixed at 10 and the CMV cation exchange membrane was used. The CO2 recovery rate increased in almost in proportional with the current density. The power consumption for recovering a unit amount of CO2 shows a minimum at 96 A m2, and increased with increasing current density after that as shown in Table 1.
0.2 6 0.1
rate pH 20 cells 10 cells
0.0
4 0
10
20
30 40 Time [min]
50
60
Fig. 10. Effect of the number of cell pairs on the CO2 recovery rate and pH in the CO2 recover cells. Feed solutions: CO2 recovery cells: 0.50 M NaHCO3; 1.2 L min1; alkali regeneration cells: 0.10 M NaHCO3; 1.2 L min1; electrode solution: 0.35 M Na2SO4; current density, 96 A m2.
A. Iizuka et al. / Separation and Purification Technology 101 (2012) 49–59
5.1. Process design and cost estimation and methods
10
0.15 8 pH [-]
CO2 recovery rate [mmol/s]
0.20
0.10 6
rate pH
0.05
CMV CSO CMD
0.00
4 0
10
20
30 40 Time [min]
50
55
60
Fig. 11. Effect of cation exchange membranes on the CO2 recovery rate and pH in the CO2 recover cells. Conditions: feed flow rates, 1.2 L min1; feed compositions: CO2 recovery cells: 0.5 M NaHCO3; alkali regeneration cells: 0.1 M NaHCO3; electrode solution: 0.35 M Na2SO4; number of cell units, 10; current, 2.0 A (current density, 96 A m2).
4.6. Effect of the cation exchange membrane Fig. 10 shows the time course of the CO2 recovery rate for various cation exchange membranes used for the electrodialysis and pH change in the CO2 recovery cells. The CO2 recovery rate was slightly higher, especially during the initial stage, when using CSO membrane, which is monovalent-ion selective. The power consumption was lowest when using the CSO membrane as shown in Table 1. However, the power consumption and current efficiency were not significantly affected by the cation exchange membrane used.
4.7. Effect of feed flow rates The feed flow rates to both cells were changed from 1.2 to 3.6 L min1. The CO2 generation rate was found to increase slightly at the increased flow rates, while the power consumption for CO2 recovery decreased with increased feed flow rate as shown in Table 1. However, the effects of flow rate on the CO2 recovery rate and power consumption were not significant. The above experimental results for the power consumption and CO2 recovery performance were comparable with the results obtained for a bipolar membrane electrodialysis system comprising a two-cell type with a different configuration [18], where potassium hydroxide rather than sodium hydroxide was used as the alkali, and an anion exchange membrane was used instead of the cation exchange membrane to separate the CO2 recovery and alkali generation cells. In this study, a cost analysis was conducted to examine the process feasibility from an economic point of view.
5. Cost estimations The cost for the CO2 recovery process of the present system was estimated using the experimental results of electrodialysis described above. First, the emission amount of CO2 to be treated with the present system was assumed to be 300 metric tons per hour (1.89 103 mol s1), which corresponds to the CO2 emission rate from a 400 MW coal-fired power plant. It is assumed that the concentration of CO2 in the flue gas is at 10% and 90% of the CO2 in the flue gas is captured in the chemical absorption process. Thus, the rate of CO2 treatment by the process, RCO2 [mol s1], is 1.89 103 0.9 = 1.70 103 mol s1. Continuous operation was assumed, 24 h per day for 350 days per year.
5.1.1. Absorption process A packing-tower type absorption tower was assumed for capturing CO2 in the flue gas with a counter flow of sodium hydroxide aqueous solution of a given concentration. The tower was made of 8-mm thickness steel (SB410 carbon steel). The tower is packed with 1-in. ceramic Raschig rings, of which the diameter is 1 in., the thickness is 1/8 in., and the surface area of the rings per unit volume of the tower, at [m2 m3]. In this study, at is assumed at 190 m2 m3, the packing density of the Raschig rings at 47,700 m3, and the void ratio of the tower, ed [m3 m3] at 0.73. The overall material balance of CO2 in the absorption tower can be expressed by the following equation,
GM LM N ¼ Q ðPB Pr Þ ¼ m ðC A2 C A1 Þ
qM
ð8Þ
where N [mol m2 s1] is the molar absorption rate of CO2 per cross-sectional area of the absorption tower; GM [mol m2 s1] is the total molar flow rates of gas, LM [mol m2 s1] is the total molar flow rate of liquid, per cross-sectional area of the absorption tower; P [Pa] is the total pressure; PT and PB are the partial pressures of CO2 at the top and bottom of the tower, respectively [Pa], CA1 and CA2 are the concentrations of the absorbent (sodium hydroxide in the present case) at the top and bottom of the tower [mol m3], respectively; qM is the molar density of the absorption liquid [mol m3]; m [–] = 2, is the stoichiometric factor for the absorption reaction. The partial pressures of CO2 at the top and the bottom of the absorption tower is PB = 0.1 atm = 1.013 104 Pa, and PT = 0.01 atm = 1.013 103 Pa, respectively assuming the 10% of CO2 in the flue gas at the atmospheric pressure to be reduced at 1% of CO2 in the ventilation gas. The sodium hydroxide concentration, CA1 correspond to that of the feed solutions to the CO2 recovery cells of the electrodialysis, which represents the CO2 absorption ratio. The concentration, CA2 corresponds to that in the alkaline regeneration cells of the electrodialysis, representing the CO2 recovery ratio. The ratio of GM/LM can be determined by Eq. (8). The gas flow rate, GM is determined based on the flooding speed, GF [kg m2 s1]. The flooding speed is the critical flow rate of gas to avoid flooding of the absorbent solution in the absorption tower. The flooding speed can be estimated based on the following empirical equation by proposed by Sawistowski [21], considering the pressure drop in the absorption tower,
" ln
G2F at g qL qG e3d
lL lW
0:2 #
" #0:25 0:5 L qG ¼ 4 G qL
ð9Þ
where qG [kg m3] and qL [kg m3] are the densities of gas and liquid, respectively; lL and lW [kg m2 s1] are the viscosities of the absorption liquid (sodium hydroxide solution) and water, respectively; G [kg m2 s1] and L [kg m2 s1] are the mass flow rates of gas and liquid, respectively; g [m s2] is the acceleration of gravity. The absorption tower is assumed to be operated under the gas flow rate, G, at 70% of the flooding speed, GF. By using the determined mass flow rate, G, the necessary diameter, DT [m], can be calculated by the following equation:
G ¼ 0:70GF ¼ MW av GM
ð10Þ
1
where MWav [kg mol ] is the average molecular weight of the gas. Once gas flow rate is determined, the diameter of the absorption tower can be determined.
DT ¼
rffiffiffiffiffiffi 4S
p
¼
rffiffiffiffiffiffiffiffiffiffiffiffiffiffi 4 RCO2 p N
ð11Þ
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A. Iizuka et al. / Separation and Purification Technology 101 (2012) 49–59
The height Z [m] of the absorption tower can be estimated by a standard method for designing absorption towers [21,22]. For an instantaneous, irreversible reaction, the tower height can be estimated by,
þ kH GM G PB þ rC A2 GM L ln 1 s M þs LM PT þ rC A2 LM ai P 1 s GLMM
ð12Þ
where kL [m s1] and kG [mol m2 s1 Pa1] are the mass transfer coefficient based on the liquid and gas sides, respectively; H [Pa m3 mol1] is the Henry constant for CO2 solution in water, ai [m2 m3] is the effective interfacial area for liquid–gas contact per unit volume of the absorption tower. The parameter r is given by,
Ps
r¼
ð13Þ
mqM
and the surface renewal parameter, s [s1] is given by,
qM H DL
s¼
P
DG
ð14Þ
where DL [m2 s1] and DG [m2 s1] are the diffusion coefficients in the liquid and gas phases, respectively. The parameters of mass transfer coefficients and effective interfacial area were estimated by the empirical correlation equations as follows.
0:7 kG RT G lG 1=3 ¼ 5:23 ðat Dp Þ2:0 at DG at lG qG DG
ð15Þ
and,
kL
qL lL g
PL ¼
1 kG
Z¼
The power consumption for a liquid pump for the absorption tower can be estimated by,
1=3
2=3 L lL 1=2 ¼ 0:0051 ðat Dp Þ0:4 aW lL qL DL
ð16Þ
where Dp [m] is the size of the ring (=1 in. = 2.54 102 m); aw [m2 m3] is the wetted area of the ring per volume of the tower. On the other hand, ai [m2 m3] can be estimated by the following correlation,
ai ¼ 0:0406L0:455 ð1000rÞn at
ð17Þ
where
n ¼ 0:83D0:48 p
ð18Þ
Based on the absorption tower design, the cost for operation and equipment were estimated. The equipment cost comprises the costs for the tower material and the packing material. The material cost for the tower construction is the weight of the steel necessary for construction of the absorption tower. The weight of the necessary steel is calculated from the tower volume, V ¼ ð1=4ÞZD2T . The tower volume is determined by the concentrations of the alkaline solutions at the inlet and outlet of the liquid flow. The cost for the packing material is in proportional to the total tower volume because the volume fraction of the packing material is fixed. Note that the volume of the absorption tower is determined by the concentrations of sodium hydroxide at the inlet and outlet of the power, in other words, the CO2 recovery ratio and CO2 absorption ratio for a given condition. Thus, the capital cost of the absorption tower can be determined by these ratios. Other costs comprising the capital cost were not taken into account in this study. The costs of the power consumption for pumping the absorbent liquid and that for blowing the flue gas into the absorption tower were considered as the operating cost of the absorption process. The cost for power consumption comprises the costs for the liquid pump (PL [W]) for transport of the absorbent solution and the blower (PG [kW]) for flue gas injection into the absorption tower.
LSZg
ð19Þ
gP
where gP [–] is the pump efficiency, and it is assumed 0.15. The power consumption for the blower is given by,
PG ¼
Dpg N RT gP P
ð20Þ
where gB [–] is the blower efficiency assumed at 0.40, and Dp is the pressure drop [Pa] in the absorption tower, which is given by Leva’s equation, 2 bL G Dp ¼ a10qL qG Z
ð21Þ
where a and b are empirical parameters determined experimentally [5] depending on the packing condition; for 1-in. Raschig ring; a = 438, and b = 51.1. The above power consumptions are then converted to the cost by using the electricity price at USD0.12 per kW h. 5.1.2. Electrodialysis process A commercialized large-scale electrodialysis system, type CG-5 by AGC, is composed of a cell stack composed of (nc=) 2400 cell pairs with an effective membrane area at Am = 1.785 m2. It is assumed that electrodialysis of the absorbent solution is performed with equipment of the same specification (nc = 2400, Am = 1.785 m2) as that of CG-5. The current efficiency, gI, is assumed to be the same as the one obtained experimentally with a laboratory-scale electrodialysis equipment used in this study for a given electrodialysis conditions. The necessary number of electrodialysis stacks, NS, to treat the CO2 form the flue gas is given by,
NS ¼
RCO2 F ¼ RCO2 r CO2 gI inc Am
ð22Þ
where rCO2 is the CO2 regeneration rate of an electrodialysis stack with nc = 2400, Am = 1.785 m2. The capital cost of the electrodialysis stacks can be estimated from the number of stacks with a cost for one stack. The total power consumption, PED [W] for the CO2 recovery can be expressed by,
PED ¼ NS VI ¼ RCO2
FV
gI nc
ð23Þ
Thus, the operation cost can be estimated from Eq. (23) with unit cost for electricity at USD 0.12. The average current efficiency over the experimental period for each experimental condition was used for the power consumption estimation. Further assumptions made for the cost estimation are shown in Table 2. 5.2. Results of the cost estimation Fig. 12a–e shows the effect of operating parameters on the total cost for CO2 capture with the present system based on the experimental results of the electrodialysis correspond to those in Figs. 5– 10. The total cost can be broken down into four sectors; equipment cost and power cost for operation of absorption and electrodialysis, respectively. However, the contributions of the absorption process, both for equipment and operation, are almost negligible compared with the costs for electrodialysis. Thus, it would be reasonable to concentrate the discussion on the cost for electrodialysis. The effect of total sodium ions in the absorption solution is shown in Fig. 12a and was not significant. With an increasing CO2 absorption ratio, the CO2 recovery cost decreased, as shown in Fig. 12b, mainly because a higher recovery ratio can increase
57
A. Iizuka et al. / Separation and Purification Technology 101 (2012) 49–59 Table 2 Assumptions used for cost estimation.
Absorption tower Packing Electrodialysis Bipolar membrane area Cation exchange membrane area Liquid pump Blower Electric power price
Specification
Cost in USD
Years in operation
Packed tower with 8 mm-thickness steel 1-in. Raschig ring 2400 cell pairs 1.785 m2 1.785 m2 Power efficiency 15% Efficiency 40%
1000 per tonne-steel 12 kg1 2.5 million unit1 750 m2 300 m2
15 2 15 5 5
0.12 kW h1
1200 Power for ED Equipment for ED
Power for ED Equipment for ED
1000
400 Cost [USD]
Cost [USD]
800
200
600 400 200 0
0 0.0
0.5
1.0
70
1.5
+
80
90
100
CO2 absorption ratio [%]
[Na ] [M]
(b) CO2 absorption ratio
(a) Total sodium ion concentration in the absorption liquid
Power for ED Equipment for ED
Power for ED Equipment for ED
400
Cost [USD]
Cost [USD]
400
200
0
200
0 0
10
20 30 CO2 recovery ratio [%]
40
0
100 -2 Current density [A m ]
200
(d) Current density
(c) CO2 recovery ratio 400
Cost [USD]
Power for ED Equipmeny for ED
200
0 CSO
CMV
CMD
Membrane CEM
(e) Cation exchange membrane Fig. 12. Effect of operating parameters on the cost for CO2 capture and recovery.
the CO2 recovery rate and consequently reduce the power consumption for CO2 recovery by electrodialysis. The number of operating cycles for a given amount of CO2 can be reduced with a higher CO2 absorption ratio and, consequently, the cost can also be reduced. The cost for CO2 recovery decreased with increasing CO2
recovery ratio but the effect was not significant, as shown in Fig. 12c. However, the experimentally tested values of the CO2 recovery ratio were 30% at the highest due to the experimental limit. However, the cost can be reduced when the CO2 recovery ratio is higher. The CO2 recovery cost was more sensitive to the CO2
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A. Iizuka et al. / Separation and Purification Technology 101 (2012) 49–59
Table 3 Optimal operating conditions.
Table 4 Breakdown of the CO2 recovery costs at the optimal condition.
Electrodialysis Membrane area Cell volume Feed flow into the CO2 recovery cell Feed flow into the alkali regeneration cell Cation exchange membrane Current density Absorption tower volume
9.2 105 m2 6.9 102 m3 1.0 M NaHCO3; 1.2 L/min 0.42 M NaHCO3; 1.2 L/min CSO 240 A m2 50 m3
absorption ratio, as shown in Fig. 12b than to the CO2 recovery ratio. The effect of current density on the CO2 recovery cost is shown in Fig. 12d. The minimum cost was achieved at a current density of 96 A m2. The equipment cost was a decreasing function of the current density and vice versa for the cost of power consumption. Thus, the minimum cost was realized at an intermediate value of current density at 96 A m2. Fig. 12e shows the CO2 recovery cost for various cation exchange membranes, tested in this study. The total cost was lowest for the CSO membrane but the difference was not significant. The CO2 recovery process was then optimized by changing all the parameters related to the experimental results with extrapolations. The optimum operating condition is shown in Table 3, which achieves the minimum cost of US$180 per ton CO2. The cost breakdown is shown in Table 4. The cost of electric power accounts about 60% of the total cost and that of membranes accounts 30% of the total cost. The minimum cost is higher than that for conventional CO2 recovery processes based on amine absorption/thermal desorption, which is in the range lower than US$100 [23], based on the usage of the best performance amine such as KS-1. About 60%
Absorption Operation Equipment
0.068 0.005
Electrodialysis Electric power Equipment Membrane
110 11 56
Total
180 USD/t-CO2
of the total cost of the amine-based absorption process can steam supply for recovery of CO2 from the solvent. Possibilities of the cost reduction were examined by sensitivity analysis of the total cost. The electric power cost can be reduced by improving the current efficiency and the electric power cost itself. The equipment cost should be significantly reduced by reducing the membrane prices. The costs for membranes should be significantly reduced under a larger scale production of ion-exchange membranes and the bipolar membrane because the material costs for the membrane fabrication may not be high; note the above costs for the membranes are based on the prices for a very small-scale supply for laboratory experimental uses. The sensitivity analysis was conducted by choosing the current density, current efficiency, and prices of the membrane and electric power as the independent variables. Fig. 13a shows the total cost of the CO2 recovery process and its breakdown for equipment, power as a function of the current density, with a fixed current density at 75%, and prices of membranes and electric power are the same as the base case shown in Fig. 12d.
600 150 Total Equipment for ED Power for ED
100 Cost [USD]
Cost [USD]
400
200
Total Equipment for ED Power for ED
50
0
0 0
200
400
0
600
200
400
600
2
-2
Current density [A m ]
Current density [A/m ]
(a) Base case
(b) Membrane prices: 1/10, electricity: 0.12 USD/kWh, current efficiency: 90%
Total Equpment for ED POwer for ED
Cost [USD]
100
50
0 0
200
400
600
2
Current density [A/m ]
(c) Membrane prices: 1/10, electricity price: 0.06 USD/kWh, current efficiency: 90%. Fig. 13. Sensitivity analysis of the cost for electrodialysis: (a) Base case: current efficiency, 75%; membrane costs: as they are; electric power cost USD 0.12 per 1 kW h, (b) current efficiency, 90%; membrane cost: 1/10; electric power cost: USD 0.06 per 1 kW h, and (c) current efficiency, 90%; membrane cost: 1/10, electric power cost: USD 0.06 per 1 kW h.
A. Iizuka et al. / Separation and Purification Technology 101 (2012) 49–59
Fig. 13b shows the total cost and its breakdown of the CO2 recovery process as a function of the current density with the higher current efficiency at 90%, and lower membrane costs at 10% of the base case. The total cost could be significantly reduced, and the minimum total cost is about US$100 per metric ton-CO2, which is almost half that of the present case and comparable with the conventional processes. In Fig. 13c the estimated cost of the process was shown when the electric power price is halved to US$ 0.06 per kW h, and the membrane costs are 1/10 and the current density is 90%. The minimum cost is about US$70, which is competitive with the conventional process of amine chemical absorption. The total cost can be further reduced by enlarging the lifetime of the membranes (5 years for the base case). It should be noted that all the estimated costs are based on the experimental results carried out with a semi-batch system in this study. The power consumption for the CO2 recovery process was significantly high at the initial stage before reaching the steady state. The CO2 recovery performance can be avoided when the system is operated in a continuous mode. A significant cost reduction can be anticipated when the present system is operated in a continuous mode. In the future, more cost reduction might be possible by reducing membrane costs through mass production so that the present CO2 recovery process may be competitive with the conventional method. Another advantage of the present process is that it can be operated near room temperature and at atmospheric pressure conditions. This enables rapid startup and shutdown of the operation and high process safety. 6. Conclusions From the above experimental results and cost estimation, the optimal conditions for the recovery of CO2 based on chemical absorption with an alkaline solution, combined with electrodialysis using bipolar membranes, was determined. The sensitivity analysis showed a significant cost reduction can be achieved when the membrane cost is reduced and the current efficiency is improved. References [1] B. Metz, O. Davidson, H. de Coninck, M. Loos, L. Meyer (Eds.), IPCC, Carbon Dioxide Capture and Storage, Cambridge University Press, UK, 2005. [2] A. Yamasaki, Journal of Chemical Engineering of Japan 36 (2003) 361–375. [3] R. Barchas, R. Davis, The Kerr-McGee/ABB Lummus Crest technology for the recovery of CO2 from stack gases, Energy Conversion and Management 33 (5– 8) (1992) 333–340.
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