Process for CO2 Capture Using Ionic Liquid That Exhibits Phase

Oct 27, 2014 - Pelayo García-Gutiérrez , Johan Jacquemin , Corina McCrellis , Ioanna Dimitriou , S. F. Rebecca Taylor , Christopher Hardacre , and R...
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Process for CO2 Capture Using Ionic Liquid That Exhibits Phase Change Ronald S. Eisinger* and George E. Keller, II Mid-Atlantic Technology, Research & Innovation Center (MATRIC), Post Office Box 8396, South Charleston, West Virginia 25303, United States ABSTRACT: A novel process for capturing carbon dioxide from the flue gas of a coal-fired power plant has been shown to reduce parasitic power consumption substantially. The process employs an ionic liquid created at the University of Notre Dame that has a high capacity for absorbing CO2 by chemical reaction. A distinguishing property of this ionic liquid is that it changes phase from solid to liquid upon reaction with CO2. The process uses heat generated by this phase transition to lower parasitic power consumption. The driving force for CO2 separation is a combination of temperature and pressure differences; the process could even work without the addition of heat. A realistic process was created to capture CO2 efficiently. Computer simulation of the process enabled calculation of viable process conditions and power usage. The main concepts of the process were shown to work using a lab-scale apparatus. Parasitic power consumes 23% of net power generation, 55% lower than that of the monoethanolamine (MEA) process. However, capital cost is higher. The cost of electricity (COE) is 28% lower than that of the MEA process.

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INTRODUCTION Concerns about global warming have spurred numerous initiatives to curb emissions of CO2. The largest target for emission reduction is electric power generation, especially coalfired power plants. Most proposals for carrying out such reductions involve retrofitting power plants to separate CO2 from the flue gas, compressing CO2, and then transporting the pressurized liquid to an appropriate site for sequestration underground. The technology closest to full-scale commercialization for separating and compressing CO2 uses a substantial fraction of the power output of the power plant. For every two coal-fired power plants retrofitted with CO2 capture technology, a third power plant would need to be built to maintain the same net power output to customers. Numerous researchers have examined the potential of ionic liquids (ILs) for sorption of CO2.1−8 ILs are salts in which either or both the cation and anion are organic components. Their melting points can be lower than 100 °C. An advantage of an IL is its very low vapor pressure. In addition, it is not necessary to use a solvent, thereby reducing heating and cooling requirements. Physical absorption of CO2 by ILs is generally too small to be viable in the separation process. A better approach is to incorporate functionality into the IL, usually an amine or related functionality. The sorption of CO2 in these cases is by chemical reaction. Prof. Joan Brennecke and fellow researchers at the University of Notre Dame9 have created functionalized ILs, in which the loading approaches 1 mol of CO2/mol of IL at the low partial pressure of CO2 in flue gas.10 Such ILs may have a role in reducing parasitic power usage during CO2 capture. Among the ILs created and tested at the University of Notre Dame, a few exhibited the property of phase change during reaction with CO2. Specifically, these phase change ionic liquids (PCILs) change from solid to liquid while capturing CO2. When the liquid IL·CO2 complex (called liquid complex in this paper) loses CO2 during regeneration, it becomes solid. The © XXXX American Chemical Society

heating requirement during regeneration is decreased by the amount of heat emitted during transition from the liquid complex to the solid phase. Further, the endothermic melting of the IL during CO2 capture will reduce the cooling usage. The concept of adding a phase-change material to an adsorbent to reduce the adverse effects of heating during adsorption of CO2 has been assessed.11 However, use of the enthalpy of phase transition of the sorbent itself to decrease parasitic power had not been considered. Sorbents that undergo phase transition during CO2 capture have been studied by others,12 but the direction of phase transition was disadvantageous and not considered for reduction of parasitic power. This work used an IL exhibiting phase change that was synthesized at the University of Notre Dame, tetraethylphosphonium benzimidazolide [P2222] [BnIm]. The structure of both the IL and the liquid complex are shown in Figure 1. The importance of the enthalpy of phase change at the regenerator can be gauged by examining two thermodynamic parameters of this IL, determined at the University of Notre Dame. The estimated endothermic heat of desorption of CO2 is 54.7 kJ/

Figure 1. Structures of (left) [P2222] [BnIm] and (right) product of reaction with CO2. Reprinted with permission from ref 9. Copyright 2014 American Chemical Society. Received: July 8, 2014 Revised: October 27, 2014

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mol of CO2. The measured heat generated by solidification of the IL when CO2 is desorbed is −19.9 kJ/mol. Thus, the reduction in heat supplied at the regenerator will be 19.9/54.7 or 36%. MATRIC, working as subcontractor to the University of Notre Dame, undertook the development of a viable CO2 capture process that takes advantage of the enthalpy of phase transition to decrease parasitic power. Each version of a process that was considered included an absorber and a regenerator. Regeneration employed not only temperature swing but also pressure swing. Promising low-cost, realistic processes were simulated using an Excel spreadsheet. The simulations provided operating conditions, including temperatures, pressures, and flow rates, and also calculated parasitic power. The basis for the simulations is a coal-fired power plant with net generation of 550 MWe (electric).13 The target for CO2 capture is 90%. Key unit operations of the most promising process were demonstrated in the lab in continuous mode. This paper focuses on those experimental results and their impact on a commercial process. Economics of the favored process were determined.

Figure 3. Phase transition diagram for IL [P2222] [BnIm]. The curved line is demarcation for solid (upper left) and liquid (lower right) phases; + denotes conditions in process simulation; and ∗ indicates experimental conditions in the spray dryer during two tests.

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THEORETICAL BASIS Modeling of CO2 sorption and phase transition is discussed here. Measured loading of CO2 on [P2222] [BnIm] as a function of CO2 partial pressure9 is plotted in Figure 2. Even at

Operating conditions at minimum parasitic power, obtained from process simulation, are denoted with + symbols for three locations in the process in the figure. These conditions were as follows: (1) Absorber top stage: in this stage, unreacted solid IL is present, along with liquid. (2) Absorber bottom stage: the sorbent complexed with CO2 leaving the absorber is liquid. (3) Spray dryer: this type of regenerator was chosen to atomize the liquid into fine droplets for rapid conversion into solids. Operating conditions are (barely) within the solid region. A substantial vacuum must be maintained in the spray dryer. The absorber comprises a multi-stage separation. A McCabe−Thiele diagram, modified for absorption, is shown in Figure 4. The equilibrium curve is based on measured sorption isotherm data, except at low CO2 loading, where data are not available. The operating line is defined by flow rates of the flue gas and IL and by inlet and outlet CO2 concentrations. The loading of CO2 in the CO2-rich slurry leaving the absorber will exceed 0.8 mol of CO2/mol of IL after just two theoretical stages. Process simulation indicates that an exit loading of at

Figure 2. CO2 solubility (loading) in [P2222] [BnIm] as a function of the CO2 partial pressure at temperatures of 60, 70, and 80 °C. Reprinted with permission from ref 9. Copyright 2014 American Chemical Society.

a CO2 partial pressure as low as 0.3 bar, loading at 70 °C reaches 0.85 mol of CO2/mol of IL. However, this plot does not provide an indication of the phase(s) present at each condition. Several increasingly complex thermodynamic models were developed at the University of Notre Dame to predict CO2 loading and liquid−solid phase transition. The most advanced model14 is a semi-empirical model using sorption isotherm data, which not only calculates conditions at which phase transition occurs but also estimates the CO2 loading in the vicinity of the transition. This advanced model is used in the current process simulation. The location of the phase transition of [P2222] [BnIm] is shown in Figure 3. It is a plot of the temperature versus CO2 partial pressure at the phase transition. Solids are favored at low CO2 partial pressure and high temperature. The liquid complex occurs to the right and below the transition curve, unless the temperature is below its melting point.

Figure 4. McCabe−Thiele diagram for CO2 sorption with IL [P2222] [BnIm]. The black line is the operating line; the blue line is the equilibrium relationship for CO2 between flue gas and sorbent; and the dashed red line shows the two theoretical stages required. B

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Figure 5. Process flow diagram for CO2 capture in a full-scale plant. slurry in the liquid complex, initially shunned, blossomed when experimental work demonstrated that a solid-rich slurry possessed flow behavior nearly as good as that of the liquid complex. The importance of this discovery for the process is immense. Before the slurry can be prepared, the liquid complex is made by reaction of IL with CO2. Synthesis of the liquid complex was carried out in the 2 L, stirred glass vessel shown in Figure 6. Before charging

least 0.6 mol/mol is sufficient for 90% CO2 capture. The actual number of sieve trays in the absorber will be considerably greater than two. Water and CO2 need to be included in the mass and heat balances of the process simulation. Not only does water affect flow rates throughout the process, but above a critical concentration, it prevents solidification of the liquid complex in the spray dryer. Measurements of water loading on both the solid IL and the liquid complex have been made.9 However, the effect of water vapor on the loading of CO2 on [P2222] [BnIm] has not been investigated. While water loading data are an integral part of the process simulation, water was not included in the flue gas composition used in the experimental demonstration.

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EXPERIMENTAL SECTION

Scope. The process envisioned for a full-scale plant is shown in Figure 5. In this plant, flue gas is compressed and cooled to dry it partially. It then enters the bottom of the absorber, where CO2 is sorbed by IL slurried in the liquid complex. The liquid complex exiting the bottom of the absorber is then heated, first by the compressed flue gas and then with steam, before entering a spray dryer for regeneration. The released CO2 gas is dried and compressed to 138 bar for transport to a sequestration site. The mixture of the liquid complex and solids leaves the spray dryer as a slurry, is cooled, and returns to the absorber. In the experimental program, each of the two main unit operations, absorption and regeneration, have been demonstrated separately in continuous mode. Flow rates were about 100 g/min for liquids and solids and 15 standard liters per minute (slpm, at 1 atm and 0 °C) for the simulated flue gas mixture. IL. The IL [P2222] [BnIm] was synthesized at the University of Notre Dame9 using a precursor material, tetraethylphosphonium bromide, that was prepared at MATRIC. It is solid at room temperature, having a melting point of 166 °C. The particles are deliquescent, leaving a liquid puddle within minutes of exposure to ambient air. Even when kept dry, solid particles agglomerate and do not flow well. Reaction with CO2 creates a viscous liquid with a melting point of 59 °C. The viscosity at 70 °C is 100 cP. The particulate IL is stored in a desiccator at room temperature. The liquid complex is stored in a sealed vessel at 70 °C in an atmosphere containing CO2. Its density at 70 °C, 1.12 g/cm3, is slightly greater than that of the IL. Additional properties of the IL can be found in ref 9. Slurry Preparation. Processes requiring handling of solids are often plagued by reliability issues. Early iterations of the process suffered from this issue, especially with the difficulty of cooling solid particles on a large scale. The concept of handling the solid IL as a

Figure 6. Prep vessel for the liquid complex and slurry.

IL particles to the vessel, the particles were screened in a glovebox to remove particles larger than 2 mm. Reaction took place at 70 °C and near-atmospheric pressure in the absence of humidity. Typically, a mixture of 34 vol % CO2 in N2 was fed to the vessel, yielding an estimated loading of 0.85 mol of CO2/mol of IL. Initial production of the liquid complex was tedious because the reaction rate decreases sharply after exposed surfaces of the particles have reacted. However, once an initial aliquot of the liquid complex was made, further production was facile. The production of a slurry is carried out by adding solid particles to the liquid complex without adding gas. The slurry flowed as a wellbehaved liquid with little change in viscosity. Slurries were made with liquid/solid ratios from 2:1 to 2.4:1 by weight. Absorber. Capture of CO2 was carried out using a distillation column in continuous, countercurrent mode. This unjacketed Oldershaw glass column had an inner diameter of 5 cm and 10 sieve trays, each separated by 5 cm. It was custom-made by Specialty C

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Glass, Inc. of Rosharon, TX. The absorber unit, including the Oldershaw column, is shown in Figure 7.

Figure 8. Spray dryer for regeneration of the liquid complex. pressure of 8 bar (100 psig), the liquid complex was atomized at a rate of 110 g/min.

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Figure 7. Absorber for CO2 capture.

RESULTS AND DISCUSSION Absorber. Nine experiments were carried out in the absorber. Their goals were to (1) identify flow issues with the slurry and flue gas, (2) determine the percentage of CO2 removed from the flue gas, and (3) explore effectiveness of CO2 capture at different CO2 partial pressures. In the first two experiments, flue gas and liquid complex were fed in countercurrent mode to observe flow issues. In subsequent experiments, a slurry of IL and liquid complex was employed to absorb CO2. Composition of the flue gas was varied from 21 to 34 vol % CO2. These concentrations are substantially higher than those from a coal-fired power plant, typically below 15 vol % CO2. The higher concentrations were chosen to mimic the CO2 partial pressure that would be experienced in a commercial absorber operating at elevated pressure. The absorber system encountered operational constraints. First, because the pressure drop is limited by a maximum safe pressure at which the glass column can be operated, the maximum gas flow was limited to 15 slpm, which corresponds to a gas velocity of 16.5 cm/s. This gas velocity was insufficient to force gas through more than half of the holes in each sieve tray. The velocity constraint decreases tray efficiency. Second, as comments in Table 1 indicate, slurry flow was often held up in a downcomer, resulting in flooding of the tray above it during filling of the column. With reference to Figure 7, downcomers are vertical, curved tubes, which extend from 4 mm above the tray down to a point a little above the next lower tray. They allow the slurry to flow by gravity through the column. Because holdup in a downcomer had also been observed when feeding the liquid complex without solids, flowability of the slurry was not the main source of holdup. Third, the gravity-driven flow rate of the slurry is constrained. When the slurry flow rate exceeded a value between 62 and 76 g/min, the slurry backed up on the top tray. The inability to feed the slurry at high rates precluded reaching the target of capturing 90% of CO2 entering the column. Absorption results are shown in Table 1. The percentage of CO2 captured from the flue gas is reported in column 4. The percentage of maximum possible CO2 recovery is shown in the

The simulated flue gas mixture consisted of CO2 and N2. The flow rate of each component was controlled by a Brooks mass flow meter to provide the desired mixture composition. The flue gas was heated before entering the bottom of the Oldershaw column. The slurry of IL and liquid complex was metered out of the prep vessel to the column using a Masterflex L/S Easy-Load peristaltic pump head with 5/16 in. inner diameter Norprene tubing. The CO2-rich slurry exiting the bottom of the column was collected in a second prep vessel. Thermocouples were located in three thermowells built into the column. The column was enclosed in a clear polycarbonate cabinet, which was maintained at about 73 °C with a Conair hair dryer. The flow rate of the CO2-lean gas exiting the top of the Oldershaw column was measured with a Brooks mass flow meter. This gas was also analyzed continuously for the CO2 concentration using a Quantek Instruments model 906 NDIR CO2 analyzer. A Moore data acquisition and control system was programmed to calculate the true exit gas flow rate from the raw flow rate and the gas composition. This calculation is needed because the response of the mass flow meter depends upon the composition of the mixture. The flow rates of the feed gases and exit gas and the composition of the exit gas were used to calculate the percentage of CO2 captured by the slurry. At a flue gas flow rate of 15 slpm, pressure drop across the column and mass flow meter at the exit was 0.075 bar (1 psi). Spray Dryer. To regenerate the liquid complex, a spray dryer was employed. It was selected because the fine droplets that it produces promote rapid CO2 desorption and phase change. Two different configurations of the spray-dryer vessel were tested. One was a 4 L glass vacuum flask. The other one, constructed of a 2 m length of a 5 cm diameter process glass pipe, is shown in Figure 8. Residence time in the tall vessel is less than 1 s. The figure shows the liquid complex being pumped through a heat exchanger and into the spray-dryer vessel. In a subsequent improvement, the gear pump was replaced with pressurized gas to push the liquid complex into the vessel. The vessel was maintained under a substantial vacuum to promote desorption of CO2. A key component of the spray-dryer system is the spray nozzle. Nozzles that provide the desired narrow-angle, full-cone pattern are not manufactured at the low flow rates used in this project. Following an experimental study of flow characteristics of purchased and homemade nozzles using viscous fluids, we selected a nozzle with a narrow, flat pattern, Unijet U150017-SS from Spraying Systems Co., with an equivalent diameter orifice of 0.33 mm (0.013 in). At a D

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Table 1. Absorber Conditions and Results in Oldershaw Sieve Tray Column fixed conditions column pressure (bottom) = 1.07 bar column temperature = 70−75 °C flue gas flow rate = 15 slpm slurry liquid/solid weight ratio = 2.4 (=2.0 for test 2b) residence time = 1.6−4.2 min slurry flow rate percent CO2 maximum possible CO2 efficiency (g/min) capture (%) recovery (%) (%)

test number

flue gas CO2 concentration (vol %)

1b 1c 2b 2c 2d 3a

34 20 34 34 34 27

50 60 117 89 76 85

3b 4

27 21

5

27

17 15−22

53 36 30 42

56 36−52

76 45

21 12−19

38 29

55 41−66

62

23

31

75

fifth column. Efficiency, shown in the sixth column, is the quotient of the percentage of CO2 captured divided by maximum possible CO2 recovery, expressed as a percent. Maximum possible CO2 recovery is the maximum percentage of CO2 in the flue gas that can be captured at the insufficient slurry flow rate used in the experiment. The absorber experiments demonstrated that a viscous slurry of IL particles in the liquid complex can flow continuously in countercurrent mode through a sieve tray column while capturing CO2. The effectiveness of CO2 capture is illustrated in Figure 9 using the last experiment. A fairly steady capture rate of 22−

comment liquid PCIL·CO2 only liquid PCIL·CO2 only slurry backed up on tray 1 could not prime peristaltic pump slurry flow blocked at downcomer 5 and then backed up on trays 1 and 2 slurry backed up on trays 1 and 2 slurry flow blocked at downcomer 1 and backed up on tray 1

was achieved at a CO2 partial pressure of 0.28 bar. Efficiencies measured in other successful experiments averaged 55%. Despite mediocre bubbling performance on the sieve trays and a low percentage of CO2 capture, the relatively high efficiencies are promising. The following three scale-up measures should enable the target of 90% CO2 recovery to be attained. First, doubling gas velocity to activate all holes can be readily achieved in a pressure-rated column. Second, the diameter of the downcomers must be widened to accommodate a high slurry flow rate. Third, it may be necessary to increase the number of trays beyond 10 to extend the residence time of the slurry in the absorber. Because literature on use of viscous fluids and slurries in absorption columns is limited, experimental data collected on a larger scale will be needed to estimate the number of trays. A key aspect of the process is compression of the flue gas before it is fed to the absorber. Process simulation indicated that flue gas compressed to 2.5 bar provides enough CO2 partial pressure, 0.39 bar, to achieve 90% CO2 capture in the absorber. The hot compressed flue gas provides most of the heating needed for the CO2-rich slurry leaving the absorber. After this heat exchange, the flue gas is cooled to 40 °C, at which temperature water is condensed. The dry flue gas proceeds to the absorber. Removal of water from the flue gas enables phase transition of the liquid complex to solid particles to occur in the spray dryer. An expander turbine powered by the CO2-lean flue gas leaving the absorber recovers 43% of the power expended by the flue gas compressor. The turbine and heat integration allow for compression of flue gas to be costeffective. Spray Dryer. The spray dryer is used to desorb complexed CO2. At the same time, solids must form to take advantage of the enthalpy of fusion. Specific goals of the spray-dryer tests are

Figure 9. Percent CO2 capture during the last absorber experiment. Flue gas = 15 slpm of 27 vol % CO2 in N2; PCIL slurry = 2.42:1 liquids/solids (wt/wt); and slurry flow rate = 62 g/min.

23% was maintained after gas and slurry flow rates lined out. The maximum possible CO2 recovery was nearly 31%; an efficiency of CO2 capture of 75% was attained. This efficiency Table 2. Operating Conditions for Spray-Dryer Experiments test number

spray-dryer vessel

pressurization mode

heated liquid complex temperature (°C)

heated liquid complex pressure (bar)

initial pressure in spray dryer (bar)

liquid complex flow (g/min)

1 2 3 4

flask flask glass pipe glass pipe

pump N2 press N2 press N2 press

140 122 139 124

7.1−8.2 8.2 8.1 7.8−8.6

0.062 0.025 0.30

109 108 108 109

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Table 3. Results of Spray-Drying Experiments spray-dryer temperature (°C) test number

a

upper

lower

pressure increase (bar)

desorbed CO2 from pressure increase (%)

desorbed CO2 from volatile analysis (%) 41a

1 2 3

100

53 51

0.019 0.015

29 23

15

4

97

67

0.008

13

14

observation droplets covered part of flask walls; liquid at the bottom bubbled after test liquid pooled at the bottom of flask; some droplets adhered to walls fine material on walls was grainy; rivulets of liquid flowed down walls; gas bubbled up through cloudy, opaque liquid at the bottom; this liquid “separated” over the next few days into a clear, yellow lower layer and a smaller, whitish, cloudy, upper layer walls remained largely clear below the top head; two rivulets flowed down walls; brown−yellow clear liquid collected at the bottom; no bubbles were seen

Value may be too high; some complexed CO2 could have been lost during the 2 weeks before collecting the solidified sample from the flask.

to (1) atomize the liquid complex, (2) desorb CO2, (3) change the liquid droplets to solid particles, and (4) determine effective temperature and pressure in the spray-dryer vessel. Four spray-drying experiments were carried out. In these experiments, liquid was pressurized to 8 bar and heated to 122−140 °C. Liquid flowed through the nozzle at about 108 g/ min. Test duration was less than 1 min for each of the first two experiments. In test numbers 3 and 4, test duration exceeded 3 min. Operating conditions are shown in Table 2. Experimental observations are summarized in Table 3. Trends of the temperature and pressure in the spray-dryer vessel are shown for test number 3 in Figures 10 and 11,

Figure 11. Pressure in the spray-dryer vessel during test number 3. The two vertical arrows indicate the beginning and end of the test. The increase in the pressure during the test is a measure of the rate of desorption of CO2 from the sprayed droplets.

number 3. Test number 4 was just outside of the desired solidphase region. Experimental evidence for CO2 desorption and phase transition to solids in these two tests mostly matched the predictions in Figure 3. A measured pressure increase in the spray-dryer vessel gives a quantitative estimate of the amount of CO2 desorbed. The vessel was calibrated for evolved CO2 flow rate as a function of the pressure increase under vacuum conditions. This relationship was used to calculate the percentage of CO2 desorbed from the liquid complex, listed in the fourth column of Table 3. Desorption of CO2 was highest in test numbers 2 and 3, and lowest in test number 4. The accuracy of the percentage of CO2 desorbed is ±10% absolute because of limited resolution of the pressure transducer. Another way to infer the percentage of CO2 desorbed is by analysis of volatiles in the material collected from the spraydryer vessel, shown in the sixth column of Table 3. Volatiles were measured by thermogravimetric analysis (TGA) at the University of Notre Dame. Unlike the pressure increase data, the TGA data suggest little difference in the percentage of CO2 desorbed during test numbers 3 and 4. Visual inspection of the inner walls and bottom of the spraydryer vessel, reported in Table 3, suggested that CO2 desorbed and/or solids formed in test numbers 1−3. In test number 4, there was no evidence of either. A corroboration of the presence of solid particles from the spray dryer is the observation of an apparent phase separation in the liquid

Figure 10. Spray-dryer vessel temperatures during test number 3. Blue points were measured 0.5 m below the spray nozzle, and red points were measured 1.5 m below the spray nozzle. The two vertical arrows indicate the beginning and end of the test.

respectively. With regard to Figure 10, because the upper temperature measurements were less-affected by heat loss, they more accurately represent the temperature of the sprayed droplets. The pressure measurements inside the spray dryer, shown in Figure 11, taken together with the temperature, are the thermodynamic parameters that determine whether phase change may occur. Test number 3 was carried out at a temperature and pressure in the spray-dryer vessel that would ensure CO2 desorption. Test number 4 attempted to demonstrate CO2 desorption at an elevated pressure, 0.30 bar. The thermodynamic likelihood of achieving CO2 desorption and phase change in these tests can be gauged from the phase transition graph in Figure 3. The ∗ symbol indicates operating conditions for these two experiments. The graph indicates a clear likelihood for success in test F

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process that takes advantage of the heat of fusion to reduce parasitic power is feasible.

collected in test number 3. Over a period of a few days, the lower 80% of the liquid lost its cloudiness. The particles causing the cloudiness apparently rose, forming a whitish, opaque suspension of particles of IL. The liquid was maintained at 70 °C during this time to avoid solidification of the liquid complex. The photograph in Figure 12 contrasts liquids collected from

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ECONOMICS Capital Cost. The cost of the process to capture CO2 from flue gas using IL exhibiting phase change is based on retrofit of a pulverized-coal-fired power plant with net power generation of 550 MWe. This economic analysis is specific to the IL [P2222] [BnIm]. Conditions of the flue gas entering the CO2 capture unit are those suggested by the National Energy Technology Laboratory (NETL)13 and assume that the flue gas is thoroughly desulfurized. Equipment sizing and parasitic power are derived from Excelbased process simulation at near-minimum parasitic power. Limits on available size of equipment dictate use of multiple pieces of identical equipment in parallel to handle the huge flow requirements. The flue gas compressor is single-stage, and the vacuum pump at the spray dryer employs two stages. Efficiency factors for each pump and compressor, which ranged from 75 to 90%, were based on values given by manufacturers. Most equipment costs are based on the process design text by Seider et al.15 Other costs were obtained from vendor quotes. The installed cost of equipment was estimated using a multiplier factor of 5.65.15 Costs have been adjusted to 2007 U.S. dollars for comparison to that of the aqueous monoethanolamine (MEA) process.16 The total installed cost of equipment is $837 million. Pumps account for 56% of these costs. Heat exchangers take up 31%, and vessels use 13%. Individual major equipment contributions to this capital cost are summarized in the pie chart in Figure 13. The most expensive item, accounting for one-quarter of the cost, is the flue gas compressor.

Figure 12. Recovered liquids collected from the spray-dryer vessel in test numbers 3 and 4.

test numbers 3 and 4. The whitish upper layer is seen in the flask on the left, from test number 3. The collected liquid on the right in the photo was free of cloudiness and remained that way. The absence of CO2 desorption and particles in test number 4 is predicted by the phase transition graph of Figure 3. However, the graph suggests that the conditions of test number 4 were very close to working. A seemingly small change in the pressure in the spray dryer is important; commercially, a singlestage centrifugal vacuum compressor cannot pull a vacuum of less than about 0.3 bar, and it is desirable to avoid using a second stage. Process simulation suggests that spray drying will work at a total pressure of 0.33 bar (CO2 partial pressure of 0.28 bar) and 96 °C but uses 24% more parasitic power than the minimum power calculated at 0.145 bar CO2 and 70 °C. The highest rate of CO2 desorption measured was equivalent to removal of 29% CO2 from the liquid complex. Process simulation indicates that 34% CO2 removal is needed when the slurry liquid/solid ratio is 2:1 by weight. This degree of removal ought to be achievable given the following benefits of commercial scale. A large spray dryer may have a residence time of 10 s, will have a much smaller bed aspect ratio (height/ diameter) to enhance back mixing and uniformity of the temperature, will be insulated, and can use large-capacity spray nozzles available with a solid cone pattern. Because regenerated IL particles form a suspension that remains stable for at least a few hours, a simple lock-hopper arrangement should be adequate to transfer the liquid suspension out of the spray dryer. Overall, the experimental results indicate not only that IL [P2222] [BnIm] is a viable sorbent for CO2 but also that a

Figure 13. Contribution of major equipment to capital cost.

Parasitic Power. The economics of the process are especially sensitive to the power consumed by the CO2 capture unit. This is so because parasitic power derates the saleable output of the power plant. The strength of the PCIL process is its relatively low parasitic power requirement. Power usage of each pump is listed in Table 4. Net pump power consumption is 123.5 MWe. The flue gas compressor consumes most of the G

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CONCLUSION A practical process to capture CO2 from flue gas of coal-fired power plants has been created, simulated, and tested using an IL sorbent that exhibits phase change. The process makes use of the enthalpy of change between solid and liquid to reduce parasitic power. It employs pressure swing in addition to temperature swing to regenerate the sorbent. A customized IL was employed to demonstrate process concepts experimentally. Operation of an absorption column with 10 sieve trays in continuous, countercurrent mode demonstrated an efficiency of 50−75% for removal of CO2 from synthetic flue gas. A spray dryer operating in continuous mode demonstrated removal of up to 29% of CO2 complexed with the IL, close to the target of 34%. Furthermore, there is evidence of particle formation in the spray dryer, inferring that the heat of fusion is available to heat the spray dryer. No desorption of CO2 occurred when the CO2 partial pressure in the spray dryer was relatively high, corroborating the advanced thermodynamic model used in the process simulation. A key finding that mixtures of solid IL and liquid complex form a wellbehaved slurry frees the process from solid-handling issues. Further development of this process should particularly address the need to raise operating pressure in the spray dryer, perhaps through modification of the IL molecule. The economics of the process appear substantially better than those of the MEA-based process. It is remarkable that the first IL exhibiting phase change to be tested on a lab/pilot scale to capture CO2 meets virtually all requirements for commercial application.

power. The product CO2 compressor accounts for 39% of net power consumption. Table 4. Power Consumption of Pumps pump

power (MWe)

flue gas compressor flue gas expander turbine CO2-rich PCIL slurry pump CO2-lean PCIL slurry pump product (CO2) vacuum pump CO2 multi-stage compressor CO2 final-stage pump grand total total excluding CO2 compressors CO2 compressors only

84.6 −36.1 1.2 0.4 24.7 45.3 3.4 123.5 74.8 48.7

Steam consumption is 1.28 × 1010 J/h. An efficiency factor of 0.3517 is used to convert steam usage into electrical power. Thus, parasitic steam usage is equivalent to 1.3 MWe. Total parasitic power is therefore 123.5 + 1.3 = 125 MWe. Steam accounts for just 1% of the total electric-equivalent parasitic power. Parasitic power derates the power plant by (125/550) × 100% = 23%. Cost of Electricity (COE). The COE is the methodology by which NETL considers both operating and fixed capital costs. The value obtained is essentially the additional price that would be paid by the consumer for electricity when the cost of the CO2 capture unit is included. The methodology used to compute COE is provided in a NETL publication.18 Table 5 summarizes the COE calculations.

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*E-mail: [email protected].

550 × TOC + OCfix + CF × OCvar)/(CF × MWh) capital charge factor, 5 years 0.124 total overnight cost $939000000 fixed operating cost, first year $2600000 variable operating cost, first year $12000000 plant capacity factor 0.85 MWe net at 100% CF 425 ¢/KWh 4.08

Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS Among those at MATRIC who assisted in this project, the skillful construction and operation of the experimental facility by Ron Brown is particularly recognized. The authors appreciate the close collaboration with faculty at the University of Notre Dame, including Professors Joan Brennecke, Mark McCready, Mark Stadtherr, Bill Schneider, and Ed Maginn. This work was carried out under subcontract to the University of Notre Dame as part of a contract with the Advanced Research Projects Agency-Energy (ARPA-E), U.S. Department of Energy, under Award DE-AR0000094.

The COE amounts to 4.1 ¢/kWh. The cost of avoided CO213 is $48/ton of CO2. Key costs are compared to analogous values for the MEA CO2 capture process16 in Table 6. Parasitic power

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Table 6. Comparison of Costs for CO2 Capture in PCIL and MEA Processes sorbent

[P2222] [BnIm]

MEA

investigator parasitic power (MWe) installed plant cost (2007 $ million) COE (¢/kWh)

MATRIC 125 837 4.1

Trimeric Corp. 276 450 5.7

AUTHOR INFORMATION

Corresponding Author

Table 5. COE: Additional Price Paid by Consumer MWe net COE = (CCF CCF TOC OCfix OCvar CF MW COE

Article

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is 55% lower for the PCIL process. However, the installed plant cost is higher. In comparison to the MEA process, the 1.6 ¢/kWh decrease in COE for the PCIL process represents a 28% reduction. H

dx.doi.org/10.1021/ef501546e | Energy Fuels XXXX, XXX, XXX−XXX

Energy & Fuels

Article

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dx.doi.org/10.1021/ef501546e | Energy Fuels XXXX, XXX, XXX−XXX