Techno-economic Evaluation of CO2 Capture from Flue Gases Using

Article pubs.acs.org/IECR

Techno-economic Evaluation of CO2 Capture from Flue Gases Using Encapsulated Solvent Anggit Raksajati,†,§ Minh T. Ho,†,‡,§ and Dianne E. Wiley*,†,‡,§ †

School of Chemical Engineering, UNSW Australia, UNSW Sydney, Sydney, New South Wales 2052, Australia School of Chemical and Biomolecular Engineering, The University of Sydney, Sydney, New South Wales 2006, Australia § CO2CRC Ltd., Canberra, Australian Capital Territory 2601, Australia ‡

S Supporting Information *

ABSTRACT: Encapsulated solvents are a solvent system for CO2 capture, where the operating solvent fluid is enclosed in a thin membrane capsule with a diameter of 100−600 μm. The encapsulation provides a significantly higher surface area compared to conventional packings, which potentially reduces the absorber dimensions. In this paper, a high-level assessment of costs for postcombustion CO2 capture using encapsulated solvent systems is carried out to identify key areas for future development. Two process configurations for an encapsulated solvent system are assessed. In the first process configuration, multiple fixed-bed columns are used as the absorber and regenerator. In the second process configuration, a circulating fluidized-bed absorber and a bubbling fluidized-bed regenerator are used. For each system, possible cost reductions through improvements in the capsule properties are investigated. Key design and operational challenges for these systems are also evaluated. The capture costs for using an encapsulated MEA 30% wt. solvent system are found to be 60% to 2 times higher than a conventional MEA solvent system. Higher capture cost is due to the extra membrane resistance in the encapsulated system which increases the regeneration energy required, coupled with higher equipment and capital cost. To reduce cost, future developments for an encapsulated solvent system should consider implementing a suitable heat recovery scheme within the process, using novel absorber and/or regenerator column designs and using solvents encased in very thin capsules. The performance of the encapsulated system could also be improved by using solvents other than MEA with more favorable properties.

1. INTRODUCTION Encapsulated solvents are a solvent system for CO2 capture, where the operating solvent fluid is enclosed in a membrane capsule that isolates the working solvent.1−4 The CO2 is absorbed through the capsule wall before reacting with the solvent while the other components in the flue gas are emitted at the top of the absorber. In a recent study by Vericella et al.,1 the capsules for an encapsulated solvent system have been produced in the laboratory with dimensions of 100−600 μm in diameter and 10−30 μm capsule wall thickness. According to the literature,1−4 one advantage of an encapsulated solvent is that it offers significantly higher surface area compared to conventional packing, which potentially leads to smaller absorber dimensions. Another advantage of an encapsulated solvent is that it provides the possibility to use solvents that are impractical in conventional CO2 absorption systems, such as solvents with slow kinetics, solvents with high viscosity and solvents that involve solid precipitation.3 According to Vericella et al.,1 the encapsulated solvent is environmentally benign because the capsule could be used to isolate degradation products and prevent the release of toxic volatile organic compounds. © 2017 American Chemical Society

To date, development of encapsulated solvent systems has focused on the manufacturing process for the capsule, the material selection for the capsule shell and preliminary assessment of the working solvent.1−4 Optimal capsule properties such as the capsule diameter, capsule thickness, and membrane capsule permeability have not been formulated. Different process configurations including fixed-bed-, circulated fluidized-bed-, and bubbling fluidized-bed columns have been proposed as possible equipment for use as the absorber and the regenerator;4 however, the process design has not been examined in detail or optimized. For example, opportunities for heat recovery between the rich and the lean sorbent streams have not been investigated. Although characteristics of the encapsulated solvent have been studied in the laboratory, to our knowledge the cost of capture using this solvent system has not been evaluated. Therefore, a high-level assessment of costs for postcombustion Received: Revised: Accepted: Published: 1604

October 23, 2016 January 3, 2017 January 19, 2017 January 19, 2017 DOI: 10.1021/acs.iecr.6b04095 Ind. Eng. Chem. Res. 2017, 56, 1604−1620

Article

Industrial & Engineering Chemistry Research

Figure 1. (a) Simplified block diagram for CO2 capture process using encapsulated solvent for case FB cases; (b) Simplified block diagram for CO2 capture process using encapsulated solvent for case CFB cases; (c) Mass transfer mechanism in absorber and regenerator. Notes: Inset in Panels a and b shows the rich capsule equilibrium with the flue gas. aDiffusion in gas phase; bDiffusion in membrane phase; cDiffusion in liquid phase; d Absorption/desorption reaction. 1605

DOI: 10.1021/acs.iecr.6b04095 Ind. Eng. Chem. Res. 2017, 56, 1604−1620

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beds are often used in parallel with cycles of multiple steps that are optimized for each separation application. In the absorption mode, flue gas is fed to the bottom section of the absorber. The flue gas flows upward through the encapsulated solvent system. The mass transfer mechanism in the absorber is shown in Figure 1c. There are four steps during the absorption process: (1) diffusion of CO2 through the gas phase, (2) diffusion of CO2 through the capsule shell, (3) diffusion of CO2 through the liquid phase, and (4) absorption inside the capsule. The CO2 absorption reaction results in a CO2-rich solvent that is accumulated within the capsules. Meanwhile, the lean flue gas is emitted from the top of the absorber. Once the CO2 breaks through at the top of the bed, the column is shifted into regeneration mode, where the column is heated in order to release the captured-CO2 using the reverse of the mechanism in Figure 1c. The heat required for the regeneration process could be supplied indirectly by LP steam flowing in embedded tubes. Another method could be to use hot inert gas (such as nitrogen or steam) through a direct heating process but this option is counter-productive as it dilutes the stripped-CO2 stream, which probably requires additional separation and energy (hence additional cost) as described for the bubbling-bed regenerator below. Between the regeneration and the absorption mode, the column goes through a cooling mode in order to decrease the temperature ready for absorption. Note that the heat recovery scheme presented in Figure 1 is one simplified example. There may be other designs and optimization of the heat recovery system; however, they have not been considered in this paper. In a fixed-bed operation, it is difficult to recover heat from the lean sorbent stream because the encapsulated solvents are not transferred between the columns. In comparison, in a conventional solvent system, this heat can be easily exchanged through the cross heat exchanger. This difficulty in exchanging the heat between the rich and lean sorbent has been raised in the literature as one of the drawbacks of fixed-bed operation such as in a TSA system.6−9 However, this becomes an even bigger concern in the encapsulated solvent system because the overall specific heat of the encapsulated solvent (consisting of the specific heat of the membrane capsule and the liquid sorbent inside the capsules) can be significantly higher than that of the specific heat of a solid adsorbent in the TSA system. 2.1.2. Fluidized Bed. A second process configuration option for encapsulated solvents is to use a circulating fluidized bed (CFB) as the absorber and a bubbling fluidized bed as the regenerator. In contrast to a fixed-bed system, a fluidized-bed system can be operated at constant temperature and steadystate. In a fluidized-bed system, the mixing of gas and solids in a properly designed bed should be more efficient than in a fixedbed column, which translates to a better heat transfer efficiency.10 In a circulating fluidized-bed absorber, the flue gas is fed to the bottom and it flows upward cocurrently with the encapsulated solvent (Figure 1b). The four-step mass transfer mechanism in the fluidized-bed absorber is the same as that of the fixed-bed absorber (Figure 1c). The CO2 absorption reaction results in CO2-rich sorbent capsules that exit the column from the top outlet. The lean flue gas is emitted from the top as well after passing through a cyclone. The rich-solvent capsules are then sent to the regenerator (a bubbling fluidized-bed column), where the captured-CO2 is stripped by temperature swing at a high temperature. The optimal heating medium for the regeneration

CO2 capture using encapsulated solvent systems is carried out in this paper to determine its competitiveness in comparison with conventional absorption processes. Two process configurations for the encapsulated solvent absorption system are compared. In the first process configuration, multiple fixed-bed columns (FB) are used as the absorber and regenerator. Each fixed-bed column in this configuration operates as both the absorber and regenerator in a cyclic step. In the second process configuration, a circulating fluidized-bed (CFB) absorber and a bubbling fluidized-bed regenerator are used. Figure 1a,b illustrates the simplified block diagram for each configuration whereas Figure 1c illustrates the mass transfer mechanism for absorption and desorption mechanism across the capsule shell. Further, the analysis identifies areas of development that could result in possible cost reduction. For each configuration, possible cost reductions through improvements in the capsule properties (for example capsule diameter, thickness, and permeability) are investigated. Key design and operational challenges including the heat recovery scheme, and the capsule integrity and stability are also investigated.

2. METHODOLOGY 2.1. CO2 CAPTURE PROCESS DESCRIPTION. The pretreatment for both process configurations includes a direct contact cooling tower (DCC) if the flue gas temperature is higher than the absorber temperature of 40 °C, and a blower to reach the absorber pressure and counter the pressure drop across the piping and the absorber column. All cases evaluated in this paper do not include FGD and SCR in the pretreatment section of the capture plant. We have assumed that FGD and SCR facilities are already installed as part of the power plant reducing the SOx and NOx levels to below 20 and 10 ppm. We have also assumed that the capsule shell material does not permit transfer of SOx and NOx and any traces of these gases exit the absorber.1 2.1.1. Fixed Bed. One option for an encapsulated solvent system involves multiple fixed-bed columns (Figure 1a), which is similar to the fixed-bed swing adsorption processes. This configuration has also been used in many industrial applications (for example, hydrogen and oxygen production), thus the process design and operation is relatively well understood. One of the challenges in employing this process for CO2 absorption is the limitation with regards to the applicable superficial gas velocity, which must be below the minimum fluidization velocity of the capsules.3 As alternatives, it may also be possible to use mesh to contain the capsules and/or a downflow configuration for the gas flow. Although these configurations would help alleviate fluidization issues, other issues with fixedbed columns including pressure drop through the bed of capsules and capsule integrity if the capsule exhibits significant compressibility under pressure would still need to be addressed. These configurations would require further research in order to identify and address potential challenges and operational problems. For processes with very large flue gas volumes, this means that very large column diameters and/or a large number of absorber units may be required. The system operates in a cyclic process and hence the column is not operated at a constant temperature, which leads to a more complex process control compared to a conventional absorption system.5 When multiple fixed-bed columns are used and follow a basic four-step cycle process, at least one column is operated in absorption mode while another is operated in regeneration mode (heating, desorption, and cooling). In practice, several 1606


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Industrial & Engineering Chemistry Research process in this system has not been formulated. One option is to apply indirect heating using LP steam through embedded tubes. Embedding a heat transfer surface in fluidized beds has been applied commercially, for example in fluidized-bed combustion of coal.11 The heat is transferred mainly by particle and gas convection, and will be discussed further in Section 2.3.1.4. In a bubbling-bed regenerator, the fluidization gas is made up mostly by the CO2 that is stripped from the capsules. However, gas flow from the bottom of the regenerator is also required to ensure proper bubbling fluidization. This could be supplied by using the recycled-CO2 stream because using an inert purge gas or steam will dilute the stripped-CO2 and would require further separation steps to produce CO2 ready for sequestration or reuse, thereby incurring additional costs. Further, the use of steam is likely to require more total heat and/or energy input to the system in order to produce the steam for fluidization and to condense the steam in the stripped-CO2. Even though there are possibilities for heat integration under consideration for different systems, at this stage the use of recycled-CO2 is likely the best option as the fluidization initiator because it requires the least amount of additional equipment and energy (i.e., only a blower to compensate for the pressure drop along the regenerator). 2.2. BASELINE CASE STUDY. This paper assumes that the emission source is flue gas from a new build 500 MW supercritical black coal power plant in Australia. The flue gas contains 13%-mole of CO2, 75%-mole of N2, 5%-mole of O2, 7%-mole of H2O, 400 ppm of SOx, and 450 ppm of NOx. The CO2 emission intensity of the power plant is 0.808 tonne per MWh electrical. The pressure and temperature of flue gas coming into the capture plant is assumed to be 1 atm and 130 °C and is consistent with our previous studies.12,13 The analysis in this paper assumes that 90% of the CO2 contained in the flue gas is recovered by the capture plant. In the post-treatment process, the separated enriched CO2 stream is dehydrated and then compressed to 100 bar ready for transport. In this paper, four baseline cases are assessed: • Case FB represents the case where multiple fixed-bed columns are used as the absorber and the regenerator. • Case CFB represents the case where a circulating fluidized-bed column and a bubbling fluidized-bed column is used as the absorber and the regenerator, respectively. • For both FB and CFB cases, the subcases are: No HI, no heat is exchanged between the rich and lean sorbent streams; HI, heat recovery is used between the rich and lean sorbent streams. The baseline capsule properties (e.g., capsule diameter, thickness, and permeability) used in this paper are shown in Table 1 and are based on the values reported by Vericella et al.1 The suitable capsule diameter for the CFB cases is much smaller than that of the FB cases. This will be discussed in Section 3.2. The material of the capsule shell is assumed to be silicone-based, similar to that reported by Vericella et al.1 This capsule shell material offers a high CO2 permeability. There was no noticeable decrease in capsule permeability during the testing, which was conducted up to 150 °C. It was also reported that mainly only CO2 and H2O permeated through the capsule with the entrainment of other gases being negligible.1 The dissolved O2 and N2 was smaller than that of CO2 by approximately 4 orders of magnitude. Accordingly, it was also reported that the encapsulated solvent system provides a higher

Table 1. Processing Conditions and Solvent Properties Used for Baseline Cases Parameter

Value

CO2 capture rate (%) Absorption temperature (K) Regenerator pressure (bar) Regenerator temperature (K) ΔT approach in HE (K) Lean loadinga

90 313 2 393 10 0.28 (FB cases) 0.23 (CFB cases) 0.23 (packed bed)

Rich loadinga

0.46 (FB cases) 0.39 (CFB cases) 0.46 (packed bed) 82 30 2000 (FB cases) 175 (CFB cases) 10 1250 3260 50 1.5 0.38 (fixed bed) 0.50 (bubbling bed) 0.91 (CFB) 0.05 (packed bed)

Heat of reaction (MJ/kmole) Solvent concentration (%-wt.) Capsule diameter (μm) Capsule thickness (μm) Capsule density (kg/m3) Membrane shell permeability (barrer)b Membrane shell price (US$/kg) Solvent price (US$/kg) Average bed void fraction

a Loading is defined as the mole fraction of carbonate conversion to bicarbonate (in liquid and solid phases). b1 barrer = 3.348 × 10−19 kmole·m/(m2·s·Pa).

overall selectivity than a gas separation membrane and would be comparable to a conventional absorption solvent system. Because the focus of this paper is to understand how the encapsulated solvent system performs compared to the conventional absorption system, MEA 30%-wt. is used as the baseline solvent, even though different types of solvent can be used within the encapsulated solvent system. The performance of encapsulated solvent systems using other solvents will be investigated in future work. The baseline solvent properties are shown in Table 1. 2.3. TECHNICAL CALCULATIONS. The calculations are performed using an in-house techno-economic model developed by UNSW Australia for the CO2CRC. This section describes the equations and correlations used in the process model. 2.3.1. PROCESS MODEL. The process model calculates the mass and energy balance for all main streams and process units described in Figure 1a,b. Figure 2 shows a flowchart of the process model used. The model calculates the key outputs of each process units (dimension and/or duty). The process is modeled assuming steady state operation. The absorber and the regenerator are modeled as equilibrium-kinetic units, whereas the other units are modeled as mixing units. 2.3.1.1. Pretreatment. Pretreatment facilities include cooling of the flue gas to the absorption temperature and a blower to provide sufficient driving force across the absorber. The maximum pressure differential provided by the blower is assumed to be 2 bar (30 psi). FGD and SCR facilities are not included in the pretreatment because it is assumed that SOx and NOx cannot penetrate the capsule shell and FGD and SCR 1607


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Figure 2. Flowsheet of process model used. Note: Fabs,out is the flow rate of absorber outlet streams, Xabs,out is the composition of absorber outlet streams, Tabs,out is the temperature of absorber outlet streams, and %-solidabs,out is the mass solid fraction of absorber outlet streams.

dimensions of the absorber include the superficial gas velocity and the overall mass transfer rate. 2.3.1.2.1. Superficial Gas Velocity. 2.3.1.2.1.1. Fixed-Bed Column (FB). In fixed-bed column operation, the superficial gas velocity must be kept below the minimum fluidization velocity to prevent fluidization of the bed.3 The minimum fluidization

are assumed to be included in the power plant from which the flue gas is obtained. 2.3.1.2. Absorber. The absorber is operated adiabatically to capture 90% of the CO2 in the flue gas. The dimensions of the absorber are determined to ensure sufficient residence time for mass transfer. Key variables necessary to determine the 1608


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slowest steps in this system and therefore control the mass transfer. If one mass transfer coefficient is much smaller than the others, it will tend to dominate the total and the larger ones can be neglected. Thus, the overall mass transfer rate for encapsulated solvents can be simplified as

velocity is calculated using eq 2, which is derived from the Ergun equation:

Ufb < Umf

(1)

in which 150μUmf (1 − εmf ) 2

2

ϕs d p εmf

3

+

1.75Umf 2ρf (1 − εmf ) ϕsd pεmf

3

1

= g (ρp − ρf )

k total

(2)

(3)

(4)

in which

Utf =

− 3.35 × 10−2p − 3.08)]] (6)

18Rept + 2.7Rept

Upt =

Ar =

(11)

Kl = 0.6792x[F(L /μ)2/3 [1 + 5.7(αe − α)M exp(0.12T

(5)

Recμ d pρp 1.687

− 20.23Ar = 0

(7)

(8)

gL3ρf (ρs − ρf ) μ2

(12)

where Kl is the liquid-side mass transfer coefficient in kmole· h−1·m−3·kPa−1, F is the packing correction factor, L is the liquid flow rate (lb·h−1·ft−2), αe is the equilibrium CO2 loading in solution (moles of CO2/mol of solvent), α is the CO2 loading (moles CO2/mol solvent), M is the molar concentration of solvent (mol/L), and p is the partial pressure of CO2 flowing into the absorber (kPa). To use eq 12 for an encapsulated system, the packing factor needs to be estimated to describe the effective surface area. The packing factor for the encapsulated solvent system is estimated using the proportion of surface area (SA) provided by the capsule (m2/m3), as follows

Reptμ d pρf

(10)

where Mu is the membrane permeability (kmole·m/m2·s·Pa), ε is the column void fraction, R is the ideal gas constant, T is the column operating temperature (K), and Wth is the wall thickness of the capsule (m). The mass transfer rate of liquid-side reaction and diffusion are estimated using an empirical equation for MEA reported in Kohl et al.5

2.3.1.2.1.2. Circulating Fluidized-Bed Column (CFB). According to Li,15 the corresponding range of gas velocity for a fast fluidization regime can be estimated as follows:

Rec = 0.7584Ar 0.73

6Mu(1 − ε)RT d pWthε

km =

2

Utf < Ucfb < Upt

1 1 1 1 1 + + ≈ + kg km kl km kl

where ktotal is the overall mass transfer rate (s−1), kg is the mass transfer rate of the gas-side diffusion (s−1), km is the mass transfer rate of the diffusion in the membrane shell (s−1), and kl is the mass transfer rate of the liquid-side reaction and diffusion (s−1). As described by Aines et al.,3 the mass transfer rate through the membrane can be calculated using the following simplified equation:

where Ufb is the superficial gas velocity in the fixed-bed column (m/s), Umf is the minimum fluidization velocity (m/s), μ is the dynamic viscosity of the fluid (cP), εmf is the void fraction of the bed at minimum fluidization, ϕs is the particle sphericity factor, dp is the diameter of the capsule (m), ρf is the density of the fluid (kg/m3), ρp is the density of the particle (kg/m3), and g is the standard gravitational constant (m/s2). The pressure drop across the fixed-bed column is calculated using the Ergun equation: 150μUfb,s(1 − εmf ) 1.75Ufb,s ρf (1 − εmf ) ΔP = + 2 2 3 H ϕs d p εmf ϕsd pεmf 3

=

(9)

where Utf is the incipient fast fluidization velocity (m/s), Ucfb is the superficial gas velocity in the circulating fluidized-bed column (m/s), Upt is the pneumatic transport velocity (m/s), Rec is the Reynolds number at the incipient fast fluidization point, Ar is Archimedes number, Rept is Reynolds number at the pneumatic transport point, and L is the characteristic body length (m). The pressure drop across the fluidized-bed column is calculated using the method described in Fan and Zhou,16 which is dependent on the sum of the pressure drop across the riser, the pressure drop through the cyclone, the pressure drop across the down-comer, and the pressure drop through the solids flow control device. 2.3.1.2.2. Overall Mass Transfer Rate. As shown in Figure 1c, the mass transfer resistance of CO2 absorption in an encapsulated solvent system occurs in four zones. The diffusion through the liquid side and the reaction resistances are often reported as the overall liquid-side mass transfer resistance.5 As described in Vericella et al.,1 membrane resistance and liquidside reaction and diffusion resistance are overwhelmingly the

SA =

πd p 2 1 πd p3/(1 6

− ε)

(13)

2.3.1.2.3. Absorber Height and Diameter. For both fixedbed and fluidized-bed cases, the absorber height is proportional to the residence time of the gas inside of the absorber and the gas velocity (calculated using eq 14). This is based on the plug flow model, which would give a conservative value for the height. The residence time is calculated using eq 15 in which the overall mass transfer rate is assumed to be first-order with respect the CO2 concentration in the flue gas.3 H=

Uτgas (1 − ε)

(14)

in which 1609


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τgas =

⎛ CCO ,out ⎞ log⎜ C 2 ⎟ ⎝ CO2,in ⎠ −K total

of units and the dimensions of the moving-bed heat exchangers required for the baseline CFB HI case are shown in Table S1 (in the Supporting Information). 2.3.1.4. Regenerator. 2.3.1.4.1. Regenerator Height and Diameter. 2.3.1.4.1.1. FB. In the FB configuration, the fixedbed columns operate as both the absorber and regenerator. The process operates in a cyclic mode and the capsule remains in the column throughout the cycle. Therefore, the dimensions of the regenerator column are the same as those of the absorber. 2.3.1.4.1.2. CFB. The fluidized-bed regenerator is modeled as a multistage counter-current bubbling bed as described in Kunii and Levenspiel.11 In conventional packed-tower regenerator design, it is typically assumed that the sections are well mixed in the liquid and vapor phases.21,22 Thus, the dimensions are designed based on a mass transfer constraint. However, this may not always be the case for encapsulated solvent systems because the liquid and the gas are not in direct contact. Therefore, the dimensions of the regenerator may be determined by a mass transfer or a heat transfer constraint. To be conservative, the final dimensions in this paper are determined based on the constraint that results in the largest column. Based on a mass transfer constraint, the dimensions of the regenerator can be determined using the same approach as for the dimensions of the absorber (eqs 14 and 16). For the calculation of the dimensions based on a heat transfer constraint, the overall heat transfer coefficient in the bubbling-bed regenerator first needs to be estimated. The heat is transferred between a heating surface and a fluidized bed in three ways: particle convection, gas convection, and radiation. The radiation step can be neglected if the bed temperature is lower than 400 °C.16 Hence, the overall heat transfer coefficient is calculated as follows

(15)

where H is the absorber height (m), U is the superficial gas velocity in the column (equal to Ufb for the fixed-bed column and Ucfb for the circulating fluidized-bed column) (m/s), τgas is the gas residence time (s), CCO2,in is the concentration of CO2 in the flue gas at the absorber inlet, and CCO2,out is the concentration of CO2 in the flue gas at the absorber outlet. The absorber diameter is proportional to the superficial gas velocity and the gas feed flow rate, as shown in eq 16. The number of baseline absorber units used is selected so that the absorber diameter can be maintained at or below 25 m, which is considered as the maximum size of the absorber that can be constructed and delivered to a capture site.17 ⎛ V ⎞0.5 D = ⎜ feed ⎟ ⎝ 4πUNabs ⎠

(16)

where D is the absorber diameter (m), Vfeed is the feed gas flow rate (m3/s), and Nabs is the amount of absorber units. 2.3.1.3. Heat Recovery. 2.3.1.3.1. FB. For case FB HI, it is assumed that a fraction of heat from the rich sorbent stream can be recovered. For the baseline case, it is assumed that 60% of the total sensible heat duty in the regenerator can be provided. This heat recovery may be achieved by utilizing the gas stream from the cooling cycle for heating in the regeneration cycle and is similar to the process scheme that has been used in the dehydration unit of a natural gas processing plant using temperature swing adsorption (TSA).14 2.3.1.3.2. CFB. For case CFB HI, where heat recovery between the lean and rich sorbent streams is used, it is assumed that multiple moving-bed heat exchangers18 are used in parallel. The heat is exchanged indirectly between these two streams using a Dowtherm A heating medium system. The heat from the lean sorbent stream of the regenerator outlet is recovered using the heating medium stream in a moving-bed cooler. This heating medium stream is then reheated using LP steam before being sent to a moving-bed heater to increase the temperature of the rich sorbent stream from the absorber outlet. The moving-bed heater is used only to increase the temperature of the rich sorbent stream, whereas the regeneration process is operated in the bubbling fluidized-bed column. Assuming the moving-bed heat exchanger is used only to increase the temperature of the rich sorbent with no desorption, the energy balance is given by the following: ⎛ ∂ 2T ⎛ ∂T ⎞ ∂ 2Ts ⎞ ⎟ ρs (1 − ε)c ps⎜vz s ⎟ = ke0⎜ 2s + ⎝ ∂z ⎠ ∂x 2 ⎠ ⎝ ∂x

h = hpc + hgc + hr ≈ hpc + hgc

(18)

where h is the overall heat transfer coefficient (W/m ·K), hpc is the particle convective coefficient (W/m2·K), hgc is the gas convective coefficient (W/m2·K), and hr is the radiative coefficient (W/m2·K). Particle convective heat transfer occurs because of the convective flow of solid particles from the bulk of the bed to the region adjacent to the heating surface. There are two resistances in the particle convection; the resistance of the average packet (group of capsules) and the film resistance. Gas convective heat transfer occurs because of the gas percolating through the bed and also by gas voids adjacent to the heating surface. The detailed methods to determine the coefficients of particle and gas convective heat transfer are described in Fan and Zhou.16 To initiate fluidization, a recycled-CO2 stream is required. Table S1 (in the Supporting Information) shows the flow rate and the velocity of this stream, which are selected to ensure proper fluidization occurs especially at the bottom of the column and to prevent entrainment of capsules at the top of the column. Table S1 also shows the assumptions used to estimate the required heat transfer surface area in the regenerator, including the values for the tube diameter, tube pitch, and tube arrangement. The bubbling fluidized bed is designed as a rectangular box to simplify the construction of the embedded tubes. This configuration is the standard configuration for lowpressure fluidized-bed systems.19 The dimensions of the regenerator for the baseline cases are based on heat transfer constraints, which give the most conservative estimate for the baseline conditions used. 2

(17)

where cps is the solid specific heat capacity (MJ/kg·K), vz is the moving-bed velocity (m/s), Ts is the temperature of the capsule, k0e is the effective thermal conductivity of a bed of solid particles, x is the horizontal coordinate with x = 0 at the centerline of the moving bed, and z is the vertical coordinate along the length of the vertical heating surface with z = 0 at the inlet for the capsules at the top of the unit. Equation 17 can be solved by the separation of variables method as explained in Yang and Hoffman19 to estimate the dimensions of the moving-bed heat exchanger. The design and costing of the moving-bed heat exchanger are based on the values provided by Solex Thermal Science Inc.20 The number 1610


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Industrial & Engineering Chemistry Research 2.3.1.4.2. Regeneration Energy. The regeneration energy in the regenerator (Qreg) is calculated as the sum of the sensible heat to increase the temperature of the regenerator inlet to the temperature of the regenerator, the heat to perform the CO2 desorption reaction with the solvent, and the water vaporization duty. Hence, the regeneration energy in the regenerator (Qreg) can be calculated using

vaporized water during the regeneration mode can be estimated as the integral of the partial pressure ratio of H2O to CO2 for the interval between the rich and lean loadings, as described in Stolaroff and Bourcier.2 2.3.1.4.3. Total Heat Duty and Electrical Energy Requirement. Methods for calculating the total equivalent electrical energy and the conversion of thermal energy to electrical power for the reboiler have been described in Raksajati et al.12 If heat recovery is applied such as a moving-bed heat exchanger in the fluidized-bed configuration (Case CFB HI), the energy requirement for the heater needs to be accounted for in the total heat duty requirement and is given by the following:

where Qreg is the regeneration energy (MJ/kg CO2 captured), Treg is the solvent regeneration temperature, which is equal to the temperature of the regenerator bottom (K), Treg,in is the regenerator inlet temperature (K), msolvent is the solvent mass flow rate at the regenerator (kg/s), mcapsule is the capsule shell mass flow rate (kg/s), cp is the overall specific heat capacity (MJ/kg.K), nCO2 is the captured CO2 molar flow rate (mol/s), ΔHabs is the heat of absorption of solvent with CO2 at the regenerator temperature (MJ/mol of CO2 captured), R is the universal gas constant (MJ/mol·K), ΔHvap,H2O is the heat of vaporization of water at the regenerator top temperature (MJ/ mol), and nH2O is the molar flow rate of the vaporized water at the regenerator (mol/s). The capsule mass flow rate (mcapsule in eq 19) is determined by the capsule shell fraction, which represents the ratio of the volume of capsule membrane shell to the total volume of capsule. The capsule shell fraction is calculated as follows VFshell = =

Q total = Q stripper + Q HR

where Qtotal is the total heat duty (MJ/kg CO2 captured) and QHR is the heat duty of the heater in the heat recovery unit (MJ/kg CO2 captured), for CFB HI case. 2.4. Solvent Properties. The process model requires thermodynamic data for solvent properties to determine the mass and energy balance for each stream and process unit. As MEA 30%-wt. is used as the baseline solvent, the empirical equations from Oyenekan et al.21 are used to estimate the equivalent CO2 partial pressure, the absorption enthalpy, and the water vapor pressure (as shown in eqs 23−25). Table S2 (in the Supporting Information) provides the list of data sources for the solvent and physical properties used in this paper.

capsule shell volume total capsule volume

1 πd p 3 6

eq

ln(P CO2 ) = A + B ·α +

1

− 6 π (d p − 2Wth)3 1 πd p 3 6

=

(23) eq

d p3

where PCO2 is the equilibrium CO2 partial pressure (kPa), and A to F are empirical parameters given by Oyenekan et al.21

where VFshell is the capsule shell volume fraction. The vaporized water molar flow rate (nH2O in eq 19) is determined based on the gas composition at the top of the regenerator, as follows

∫ pH2Otop dα ∫ pCO2,top dα

C α2 α α + D· 2 + E· 2 + F · T T T T

d p3 − (d p − 2Wth)3 (20)

n H 2O =

(22)

eq

∂ln(P CO2 )

( T1 )

∂

=

ΔHabs α2 α = C + 2D· + 2E · + F ·α −R T T (24)

where ΔHabs is the heat of absorption (kJ/mol) and R is the gas constant (kJ/kmole·K)

nCO2 (21)

ln(P H2O) = G +

where pH2O,top and pCO2,top are the partial pressure of water vapor and CO2 in the gas phase at the top of the regenerator (Pa) respectively, and α is the solvent loading (moles CO2/mol solvent). For the CFB configuration, the regenerator is modeled as a multistage counter-current bubbling bed, as described in Kunii and Levenspiel.11 The regeneration energy duty is provided by the reboiler at the bottom of the column. A vapor stream flows upward (consisting of CO2 and H2O), however for this high level assessment and consistent with the findings of Stolaroff and Boucier,2 we assume that the amount of water lost from the capsule during regeneration is minimal and does not affect the composition of the solvent inside the capsule Thus, the gas composition at the top of the regenerator is constant for a steady-state operation. For the FB configuration, the gas composition at the top of regenerator varies throughout the operation because it is a cyclic and dynamic process. As a result, the amount of

H + I ·ln(T ) + J ·T K T

(25)

where PH2O is the vapor pressure of the solution (kPa), and G to K are empirical parameters given by Oyenekan et al.21 2.5. ECONOMIC ASSUMPTIONS. Methods for calculating the capture plant capital cost and operating cost, and the specific cost of CO2 avoided have been described in Raksajati et al.12 The cost year of this analysis is 2011, with cost reported in US dollars (US$). The costs are evaluated on a pretax basis using a real discount rate of 7% assuming the project life is 25 years. The load factor of the power plant and the capture plant is 85%. The cost for coal is US$1.5/GJ. The solvent and capsule shell prices used in this paper are shown in Table 1. In this paper, the encapsulated solvent is assumed to be replaced every 5 years to maintain the performance of the membrane shell and the solvent inside. In practice, replacement times may be longer or shorter than this. However, there is limited information on capsule lifetime and replacement to date. 1611


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calculated variables from the process model and the values of these variables from the literature. For the validation, the solvent and operating conditions (such as temperature and pressure of the absorber and the regenerator) are the same as those from Vericella et al.1 and Stolaroff and Bourcier.2 In these studies, the sensible duty and water vaporization duty of an encapsulated solvent system using sodium carbonate 30%-wt. were reported. Vericella et al.1 reported that for an encapsulated solvent using sodium carbonate 30%-wt., the ratio of water evaporation duty between H2O to CO2 ratio is 1:7.5, which translates to water vaporization duty of 0.13 MJ/kg CO2. In this paper, a similar number is also obtained. According to Stolaroff and Bourcier,2 the sensible heat duty in the encapsulated solvent system using sodium carbonate 30%-wt. is expected to be higher than a MEA solvent used in a conventional absorption system, and accounts for about 27% of total regeneration energy. This translates to a sensible duty requirement of 0.85 MJ/kg CO2. Using the same assumptions as Stolaroff and Bourcier,2 the duty is estimated to be 0.84 MJ/kg CO2, which is within 1% of their number. The heat of reaction duty for sodium carbonate in an encapsulated solvent system was not quantitatively measured and reported in either studies by Vericella et al.1 or Stolaroff and Bourcier.2 However, the heat of reaction of a solvent in an encapsulated solvent system is expected to be calculated in the same manner as the heat of reaction in a conventional absorption system because the

Figure 3. Capture cost breakdown and total heat duty for the baseline cases investigated.

3. RESULTS AND DISCUSSION 3.1. PROCESS MODEL EVALUATION. As there is no published study that has evaluated the performance of encapsulated solvent systems as a complete CO2 absorption unit to date, values for the dimensions of the absorber and total regeneration duty cannot be used to compare with those obtained in this study. However, to validate our process model the calculated values of variables that are used to determine the total regeneration energy and absorber dimensions have been benchmarked with available studies. Table S3 (in the Supporting Information) shows the key inputs required, the Table 2. Key Outputs for Baseline Cases Parameter

FB No HI

FB HI

CFB No HI

CFB HI

Benchmark MEA

Total heat duty (MJ/kg CO2) Regeneration energy in the regenerator (MJ/kg CO2) Sensible heat (MJ/kg CO2) Solvent sensible heat Capsule sensible heat Heat of reaction (MJ/kg CO2) Water vaporization duty (MJ/kg CO2) Heat recovery duty (MJ/kg CO2) Energy penalty (MWe) Number of absorbers Absorber bed height (m) Absorber bed diameter (m) Absorber pressure drop (bar) Superficial gas velocity in the absorber (m/s) Specific surface area of capsule/packing in absorber (m2/m3) Number of regenerators Regenerator bed height (m) Regenerator bed diameter/width (m) Total equipment cost (US$/kW) Capture cost (US$/ tonne CO2 avoided) Capture plant Capex (US$/kW) Capture plant Opex (US$/kW·year) Equipment cost breakdown (%) Vessels (absorber, regenerator, reboiler) Encapsulated solvent (working solvent + capsule) Solvent system (filter, humidifier, etc.) Internal heat exchangers Auxiliaries (pumps, blower, thermal oil system, pneumatic transport) Dehydration Unit Compressor General facilities

11.01 11.01 5.00 4.94 0.06 1.86 4.14 n/a 509.8 6

7.74 7.74 2.58 2.55 0.03 1.86 3.29 n/a 310.2 5

11.75 11.75 8.21 6.91 1.30 1.86 1.68 n/a 614.0

6.21 5.59 2.05 2.05 1.73 1.86 1.68 0.62 297.9

24.3

19.7

0.33 0.70 1860 6

23 16 5 12 4 7 20 13

29 1 7 3 1 9 34 17

16.7 0.29 4.75 2950

5

24.8 1507 197.4 3092 230.9

24.3 1198 134.4 2452 171.4

1 36.4 23.9 1097 196.4 2267 195.6

57 5 3 3 2 5 11 14

55 5 3 3 2 5 14 14

34 13 6 8 2 7 17 13

4.4

1612

1 17.3 20.2 802 104.2 1639 114.8

1 15.1

4.4 24.8

4.35 4.35 1.23 1.23 0 1.86 1.26 n/a 178.5 1 15.5 21.3 0.1 2.48 250 1 12.4 14.2 506 66.6 1037 92.1


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Figure 4. Total heat duty breakdown for the baseline cases.

Figure 5. Effect of lean loading on regeneration energy (a) for case FB HI (rich loading fixed at 0.46 mol/mol), (b) for case CFB HI (rich loading fixed at 0.39 mol/mol).

Figure 6. Effect of capsule diameter at a fixed number of absorbers on capture cost (a) for case FB HI, (b) for case CFB HI.

reaction occurs at the same operating temperature and pressure. The absorber height is calculated as a function of the superficial gas velocity, the overall bed void fraction, and the overall mass transfer rate (eqs 14 and 15). The superficial gas velocity is calculated using equations reported in Li15 (eqs 1 and 2 for FB configuration and eqs 2−9 for CFB configuration). The overall bed void fraction for each FB and CFB configuration is within the range of typical bed void fraction of packed-bed and fast-fluidized-bed operations as reported by Kunii and Levenspiel.11 The overall mass transfer rate consists of the mass transfer rate of the diffusion in the membrane shell and the mass

transfer rate of the liquid-side reaction and diffusion (eq 10). The values calculated by our process model for these two parameters are identical with the values reported in Vericella et al.1 Our model estimates the same value of the specific surface area of the capsule in a fixed-bed column as that reported by Vericella et al.,1 who reported the specific surface area as a function of capsule diameter (between 100 and 700 μm); for example, the specific area for a 500 mm capsule is 7440 m2/m3. 3.2. BASELINE ECONOMIC RESULTS. Figure 3 summarizes the total capture cost for the baseline cases investigated and for the benchmark MEA in a conventional packed-column, whereas the key outputs of the baseline cases are shown in Table 2. As shown in Figure 3, the capture costs for the baseline 1613


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absorber of a fluidized bed is more than 6-fold higher than for a fixed bed, thus a smaller area is needed (Table 2). There is an opportunity to reduce the capital cost of an encapsulated solvent system (both FB and CFB cases) by using carbon steel (CS) as the material of construction for some equipment (e.g., the absorber, stripper, and piping). Unlike the conventional MEA CO2 absorption where alloy is required (typically SS 304) due to the corrosive nature of MEA, CS may be applied in the encapsulated solvent system because the solvent does not directly contact the vessel material. The capture cost reduces by approximately 8% for FB cases and 4% for CFB cases if the absorber and regenerator (including the reboiler) are constructed from CS. Figure 4 shows the total heat duty breakdown for the baseline cases. The regeneration energy for the baseline FB and CFB cases are higher than that of a conventional MEA system by 84% and 23% respectively with HI, and about 153% and 170% respectively for no HI. The higher regeneration energy for the encapsulated systems is due to the higher sensible heat duties and water vaporization duties, while the heat of reaction duty remains the same. The higher sensible heat duty for an encapsulated solvent system compared to that of a conventional solvent system arises because of two causes: the additional capsule shell sensible heat and the lower working capacity that can be applied. As shown in eq 19, the capsule shell sensible heat is an extra component in the total regeneration energy. The results in Table 2 also show that the heat required for capsule shell sensible duty is higher for CFB cases than FB cases. This is because the suitable capsule diameter for the CFB cases is much smaller than that of the FB cases resulting in a much larger fraction of capsule shell that needs to be heated. In the baseline cases, the capsule diameter used in the FB configuration is about 10 times larger than that of the CFB configuration (Table 1) because of the different superficial velocities that are needed for each configuration (Table 2). If both configurations (FB and CFB) use the same capsule diameter size (for example 200 μm) and with the absorber diameter limited to 25 m, in the case of the CFB configuration one absorber is needed because of the high velocity that is generated in the absorber (to reach the fast fluidization regime). In contrast, for the FB configuration, because the capsule remains unmoved within the fixed-bed column (a low velocity is required within the absorber to keep it below the minimum fluidization velocity) and the number of absorbers needed is 10. Therefore, for the baseline FB configuration in order to reduce the number of absorbers needed (6 units), comparatively larger capsule diameters of 2000 μm are used. For both the FB and CFB cases, the working capacity is lower than that in a conventional absorption system. However, the reasons for the lower working capacity in each process configuration are different. For the FB cases, the working capacity is lower because of a higher lean loading, whereas for the CFB cases, the working capacity is lower because of a lower rich loading. In the case of the FB configuration, the working capacity is smaller than a conventional system because of the higher lean loading needed to prevent excessive water vaporization in the regenerator. In this study, the lean loading for the FB configuration was selected to be 0.28 mol CO2 per mole solvent compared to 0.23 mol CO2 per mole solvent for the conventional system. This value was selected as it resulted in the lowest regeneration duty for a range of lean loadings when

Figure 7. Effect of capsule diameter at a fixed number of absorbers on absorber diameter and height (a) for case FB HI a, (b) for case CFB HI.

encapsulated solvent systems are between 65% and 223% higher than that of commercial MEA solvent absorption in a conventional packed-column. This cost difference arises due to the higher capital costs and regeneration energy of these systems; for the encapsulated solvent system, the capital cost is about US$ 1640−3090/kW which is 58−198% higher than the conventional packed-column system of US$ 1037/kW (Table 2). For the FB cases, the higher capital cost is due to the large number of absorber and regenerator units required (five units for the baseline FB cases). A large number of columns is needed for the FB cases because the superficial gas velocity applicable in the columns is very low to prevent fluidization. For CFB cases, the high capital cost (compared to the benchmark conventional MEA) arises because of high capital costs of the moving-bed heat exchangers and the regenerator; with the capital cost of moving-bed heat exchangers being up 5 to 7 times higher than the cost of a typical cross-heat exchanger (Table 2). The moving-bed heat exchanger also has a lower overall heat transfer rate because of the lower contact and mixing efficiency, thus multiple units of MBHE with large dimensions are required (Table S1). Further, the CFB regenerator is larger than that of a conventional MEA system because extra space is required to install embedded tubes that are used to supply the heat required in this unit adding to the cost (Table 2). Between the two encapsulated processes, the capital cost of the CFB cases is lower than FB cases because only one unit of absorber and regenerator is required for the CFB configuration compared to five units for the FB configuration. The higher gas velocity that can be used in the 1614


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Figure 8. Effect of capsule diameter at a fixed number of absorbers on total heat duty and shell volume percentage (a) for case FB HI, (b) for case CFB HI.

the rich loading was fixed at 0.46 mol CO2 per mole (Figure 5a). As shown in Figure 5a, when the lean loading increases from 0.23 to 0.32 mol/mol the sensible heat of the system increases because the solvent working capacity decreases resulting in higher solvent circulation rates (msolvent in eq 19). At the same time, the vaporization duty decreases because when the lean loading increases the partial pressure ratio of H2O to CO2 at the top of the regenerator also decreases (eq 21). Between lean loadings of 0.23 and 0.28 mol CO2 per mole solvent, the overall regeneration duty decreases because the reduction in vaporization duty is greater than the increases in sensible heat. Beyond this point, the total regeneration energy starts to increase because the higher sensible heat outweighs any benefit of having a lower vaporization duty. For the CFB cases, the lower working capacity compared to the conventional system arises because a lower rich loading that is used: 0.39 mol CO2 per mole solvent compared to a rich loading of 0.46 mol CO2 per mole of solvent in the conventional system. In the CFB configuration, the feed gas stream and the capsule containing the lean solvent flow cocurrently from the bottom of the absorber (Figure 1b). Thus, the rich-solvent leaving the top of absorber is close to equilibrium with the lean flue gas (inset of Figure 1b). In contrast, in a conventional packed-column or FB absorber the rich solvent is in equilibrium with the feed gas at the bottom of

the absorber (inset of Figure 1a) which has a much lower partial pressure of CO2 (that is, the concentration of CO2 in the lean gas at the absorber outlet). As a consequence, much higher rich loadings can be achieved in a conventional system compared to the CFB configuration. Although the rich loading for the CFB cases is lower than for the conventional system, with the lean loading of the configuration being similar at 0.23 mol CO2 per mole of solvent. Similar to the FB configuration, the lean loading value was selected based on the value which resulted in the lowest regeneration duty for this configuration (Figure 5b). To overcome this limitation of the cocurrent flow of the CFB encapsulated solvent system, Yang and Hoffman19 have proposed a process whereby multistages CFB are used. This alternative process option could enable the gas and sorbent streams to flow counter-currently and allow the solvent to be operated at higher rich loadings similar to those in a conventional packed-column absorber. If this can be achieved without adversely impacting the absorber dimensions, the capture cost for case CFB HI would be reduced to US$ 87 per tonne CO2 avoided and the regeneration energy would be 5.0 MJ per kg CO2. Figure 4 also shows that the water vaporization duties for both encapsulated solvent processes are much higher than that for a conventional absorption process. For the FB process, the vaporization duty of the FB cases is more than triple than that 1615


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Figure 9. Effect of capsule permeability on (a) capture cost and absorber height, (b) mass transfer resistance.

regenerator decreases, and hence for the CFB cases, the partial pressure ratio of H2O to CO2 and the amount of water vaporized is also higher than for the conventional process (eq 19 and eq 21). As a result of the very high regeneration energy for the FB and CFB No HI cases presented in Table 2, the capture costs are estimated to be above US$ 200 per tonne CO2 avoided (Figure 3 and Table 2), making it unlikely that an encapsulation system would not have a heat recovery system in practical applications. Thus, the following sections in this paper will exclude the results of the No HI cases. 3.3. IMPACT OF CAPSULE PROPERTIES. Based on the above results, encapsulated solvents will only be economically competitive with other solvent-based capture technologies if the capital cost and the total heat duty can be reduced. One way to achieve this is through improvements in capsule properties. This includes optimizing the capsule diameter, increasing the CO2 permeance through the capsule shell, and improving the capsule shell stability and integrity. Increasing the CO2 permeance through the capsule shell can be achieved through increasing the capsule shell permeability and decreasing the capsule thickness. In the following analysis, one capsule property is varied from the baseline value at a time, with the other parameters held constant, in order to identify which parameters have the most effect on total cost. 3.3.1. Capsule Diameter. In an encapsulated solvent system, the capsule diameter affects the hydrodynamics of the columns

of conventional MEA, whereas for the CFB cases, the water vaporization duties are about 33% higher. As shown in eq 21, the amount of vaporized water is calculated as the integral of the partial pressure ratio of H2O to CO2 for the interval between the rich and lean loadings over p the integral for CO2 ( H2O ). In a conventional MEA process, pCO

2

because the composition of gas at the top of the regenerator is constant over time and is in equilibrium with the liquid phase, p the integral can be simplified to ( H2O ) at the stripper top pCO

2

conditions. In contrast, in the case of the FB configuration, the solvent loadings (and composition of CO2 gas) at the top of the regenerator are not constant. Because of the cyclic nature and dynamic process of the FB configuration, the solvent loading and the CO2 partial pressure at the top of the regenerator is similar to that of the conventional process only at the start of the regeneration cycle and then gradually decreases as the regeneration step continues. Therefore, the overall integral of the ratio of water partial pressure to CO2 partial pressure over time is much higher for the FB cases compared to a conventional MEA system resulting in a higher vaporization duty (eq 19 and eq 21). For the CFB cases, the water vaporization duty is higher than that of a conventional MEA system because of the lower rich loading that is achieved. Following eq 23, when the rich loading decreases the partial pressure of CO2 at the top of the 1616


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Figure 10. Effect of capsule thickness on (a) capture cost and absorber height, (b) total heat duty.

corresponding superficial gas velocity in the absorber, and the corresponding column diameter increases (Figure 7a,b). For the FB and CFB configurations, the absorber diameter increases by 26% and 50% respectively if the capsule diameter halves (from 2000 to 1000 μm for FB HI and from 200 to 100 μm for CFB HI). This is because the superficial gas velocity is proportional to the exponential function of capsule diameter. However, the capsule diameter is inversely proportional to the contactor surface area (m2/m3), as shown in eq 12, thus smaller capsule diameters provide a larger surface area and result in a shorter column height (Figure 7a,b). A shorter column also translates to a lower pressure drop in the column, which leads to a decrease in the blower duty and capital cost. As shown in Figure 7a,b, the absorber height for FB and CFB decreases by 65% and 76% when the capsule diameter halves When the capsule diameter decreases, the absorber height decreases more steeply for CFB cases (Figure 7b) than for FB cases (Figure 7a). This is because a fluidized-bed column has a higher void fraction compared to that of a fixed-bed column (Table 1), which results in the need for a taller absorber column (eq 13). The capsule diameter is also inversely proportional to the capsule shell volume percentage (eq 20), thus a smaller capsule diameter (at fixed capsule shell thickness) results in a higher shell volume percentage. As shown in Figure 6a,b, the range of optimum capsule diameter for FB is substantially bigger than that for CFB (as explained in Section 3.2) and thus the shell volume percentage

Figure 11. Effect of capsule replacement frequency on capture cost for constant capsule properties.

(such as the applicable superficial gas velocity), the surface area of the capsule, and the regeneration energy requirement. For a fixed number of absorber units (as shown in Table 2 for each case), Figure 6a,b shows the impact of capsule diameter on the capture cost for case FB HI and CFB HI respectively whereas Figure 7a,b shows the impact on column dimensions, and Figure 8a,b shows the impact on heat duty and shell volume. As shown by eqs 1 and 2 for the FB and eqs 4−9 for the CFB configurations, decreasing the capsule diameter decreases the 1617


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duty (Figure 10b) for both configurations through reductions in the sensible heat duty. This is especially pronounced for the CFB cases because of the higher shell volume percentage for this configuration. 3.3.4. Capsule Integrity and Stability. One of the components of operating costs for the absorption plant is the solvent replacement cost. This is due to the loss of solvent through vaporization and degradation of solvent over time. For an encapsulated solvent system, the replacement rate is also affected by the expected lifetime of the capsule shell in order to maintain its performance. Vericella et al.1 reported that the solvent loss for an encapsulated solvent system is expected to be lower than that of conventional system, especially because of significantly lower solvent vaporization in the absorber. This is due to the fact that water and CO2 can move across the membrane capsule, but not the solvent. Figure 11 shows the effect of changes in capsule replacement frequency on capture cost. The results show that capsule replacement frequency only slightly affects the capture cost, over a range of membrane shell prices of 10 to 100 US$ per kg. At the baseline shell price of US$ 50 per kg, if the replacement frequency decreases from 5 years to 2 years, the capture cost increases by approximately 3% for all cases. If the replacement frequency doubles from 5 years to 10 years, the capture cost decreases by approximately 1%. For a more expensive membrane shell material costing US$ 100 per kg, doubling the capsule replacement frequency to 10 years decreases the capture cost by 2%. As shown in Figure 11, the results suggest that although capsule replacement frequency does not have a critical influence on cost for CFB encapsulated solvent systems, the price of the membrane shell does. At a fixed replacement frequency, changing the membrane shell price from US$ 10 per kg to US$ 100 per kg increases the cost for the CFB cases by 20% and for the FB cases by 4%. The cost reduction is much more evident for CFB cases because the capsule accounts for about 13 to 16% of the total capital cost (Table 2), thus any reduction in price for the membrane shell price has a significant impact on the overall capture cost. For the FB cases, the costs are dominated by the absorber and regenerator vessels due to the number of multiple units required, and thus a reduction in membrane shell price has less impact.

for FB cases is significantly lower than for CFB cases. As a result, the capsule sensible heat duty increases significantly for CFB cases (Figure 8a) when the capsule diameter decreases, but these are negligible for FB cases (Figure 8a). For the FB HI case, the lowest capture costs occur at the trade-off point at which the capsule is small enough so that a high surface area is achieved, but the corresponding gas velocity is high enough to limit the absorber diameter (or the number of absorber units). For the CFB HI case, the same trade-off occurs coupled with an additional constraint of capsule sensible heat duty, such that the diameter must be big enough to limit the capsule sensible heat duty. For cases FB HI and CFB HI, a capsule diameter of 2000 and 175 μm, respectively, results in the lowest costs (as shown in Figure 6a,b). 3.3.2. Capsule Shell Permeability. Another possibility for reducing the capture cost is by increasing the permeability of the CO2 through the capsule shell. The mass transfer rate through the membrane increases when the capsule permeability increases (eq 11), and hence the overall mass transfer rate increases (eq 10). As shown in Figure 9a, the capture cost is largely insensitive to permeability for both FB and CFB systems unless at low values (less than 3000 barrer) when a decrease in mass transfer rate through the membrane (eq 11) leads to an increase in the column height (eqs 14 and 15). As shown in Figure 9b, the mass transfer resistance in the capsule shell (km) decreases when the capsule permeability increases, which leads to a decrease in the overall mass transfer resistance (kov). Above 3000 barrer, the cost is largely insensitive to capsule permeability because the absorber height decreases less steeply due to overall mass transfer resistance decreasing with increasing membrane permeability (Figure 9b). Even though membranes with a high permeability have been reported in the literature (up to 28 000 barrer),23−25 the results show that permeability above 3000 barrer is unlikely to reduce the costs further. An interesting feature of an encapsulated solvent is that the overall process permeability and selectivity can be independently modified, unlike a gas separation membrane process where these two parameters are typically competing parameters.23 This decoupling of the parameters can be achieved by improving the permeability of the capsule shell and the selectivity of the solvent that reacts with CO2 but not with N2.1 3.3.3. Capsule Shell Thickness. Another way to improve the permeance of the CO2 through the capsule shell is by reducing the wall thickness. This could be achieved by improving the capsule manufacturing process or by changing the capsule material. The membrane capsule needs to be thin enough in order to prevent excessive mass transfer resistance, but also strong enough to maintain its integrity and stability. This is especially pertinent during the regeneration process where the capsule is heated to a high temperature to release the capturedCO2. As shown in Figure 10a, the absorber height decreases with capsule thickness (at fixed other capsule properties) because the resistance of the membrane to mass transfer decreases (eq 10). Figure 10a shows that the impact of capsule shell thickness on the absorber height for the FB cases is much less sensitive compared to that of the CFB cases. This arises because the void fraction of the FB column is less than half that of a CFB column and there is less membrane resistance for the FB cases (eq 11). In addition to affecting the mass transfer, changes in the shell thickness also affect the heat transfer because decreasing the capsule shell thickness decreases the total heat

4. CONCLUSION The scoping level economic assessment shows that the capture costs for an encapsulated solvent system using MEA 30%-wt. without internal heat recovery is 195−223% higher than the capture cost of a MEA 30%-wt. solvent in a conventional absorption system. This is because of the higher capital cost and regeneration duty required. Among the different configurations assessed, the fluidizedbed configuration results in lower capture cost. In comparison, the fixed-bed configuration (FB) is more expensive because of the large number of absorber and regenerator columns needed. This is due to the lower superficial gas velocity required to prevent fluidization which is approximately an order of magnitude lower in the FB configuration than that of the CFB configuration. The other disadvantage for the FB configuration is the high water vaporization duty because of the equilibrium conditions at the top of regenerator column. For the encapsulated solvent, adding heat recovery between the rich and lean sorbent streams (FB and CFB HI) reduces the overall cost by 37−47%. However, using the properties of 1618


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Industrial & Engineering Chemistry Research Notes

current capsules and solvents, the capture costs for encapsulated solvent systems are uncompetitive with the conventional MEA process by 56−102%. For encapsulated solvent systems to be economically competitive, further improvement in the heat recovery scheme, reduction in the membrane wall thickness by up to an order of magnitude, and development of novel absorber and/or regenerator columns are necessary. Increasing the capsule permeability beyond 3000 barrer and changing the capsule diameter have less impact on cost reduction. Further cost reductions could also be achieved through the use of other solvents. For example, sodium carbonate has been proposed as an alternate solvent because of its lower heat of reaction and vaporization duty. There are other challenges that need to be addressed if encapsulated solvent systems are to become attractive and competitive, for example managing the water balance inside the capsule. Although encapsulated solvent systems should be more tolerant to changes in the water balance,2 an additional process would be required to rehydrate the capsule if some of water inside the capsule is vaporized and diffuses out of the capsule during the regeneration process along with the stripped-CO2. In a conventional liquid sorbent system, the vaporized water can be returned to the stripper column as a liquid reflux. To the best of the authors’ knowledge, water loss in encapsulated solvent systems and methods to compensate for such loss have not been evaluated to date. The other potential challenges are maintaining capsule integrity and evaluating the capsule resistance to impurities in the flue gas. It is important to develop and select capsule materials that are not permeable to or affected by the presence of SOx and NOx (i.e., the capsule membrane must not plasticize or age in the presence of impurities). If the capsule material can be penetrated by SOx and NOx, it will be very difficult to remove degradation or side-products (e.g., sulfate and nitrate salts) from inside the capsules and the capsules will need to be replaced. If the capsule material is significantly affected by SOx and NOx, it will be essential that FGD and SCR facilities within the power plant (as assumed in this paper) can provide low impurity flue gas to the capture system. If not, additional or enhanced pretreatment facilities would need to be included in the capture system to lower the SOx and NOx concentration entering the absorber, which would make this solvent system even less economically attractive than shown in this paper.

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The authors declare no competing financial interest.

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ACKNOWLEDGMENTS The authors acknowledge the funding provided by Australian Government through the CRC Program. We gratefully acknowledge Solex Thermal Science Inc. for providing the sizing and costing of the bulk solid heat exchanger. One of the authors (A.R.) also acknowledges UNSW Australia and the Faculty of Engineering for scholarship funding.

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ASSOCIATED CONTENT

S Supporting Information *

The Supporting Information is available free of charge on the ACS Publications website at DOI: 10.1021/acs.iecr.6b04095. Tables showing variables for the design of moving-bed heat exchanger and regenerator in the baseline CFB cases, the list of the data sources for the solvent properties, and data for the process model validation (PDF)

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REFERENCES

(1) Vericella, J. J.; Baker, S. E.; Stolaroff, J. K.; Duoss, E. B.; Hardin, J. O.; Lewicki, J.; Glogowski, E.; Floyd, W. C.; Valdez, C. A.; Smith, W. L.; Satcher, J. H.; Bourcier, W. L.; Spadaccini, C. M.; Lewis, J. A.; Aines, R. D. Encapsulated Liquid Sorbents for Carbon Dioxide Capture. Nat. Commun. 2015, 6, 6124. (2) Stolaroff, J. K.; Bourcier, W. L. Thermodynamic Assessment of Micronencapsulated Sodium Carbonate Slurry for Carbon Capture. In Proceedings of the 12th International Conference on Greenhouse Gas Control Technologies, Austin, USA, November 2014. (3) Aines, R. D.; Spaddaccini, C. M.; Duoss, E. B; Stolaroff, J. K.; Vericella, J.; Lewis, J. A.; Farthing, G. Encapsulated Solvents for Carbon Dioxide Capture. In Proceedings of the 11th International Conference on Greenhouse Gas Control Technologies, Kyoto, Japan, November 2012. (4) Aines, R. D., Bourcier, W. L., Spaddaccini, C. M., Stolaroff, J. K. Polymer-encapsulated Carbon Capture Liquids that Tolerate Precipitation of Solids for Increases Capacity. U.S. Patent WO 2013063501 A2, May 2013. (5) Kohl, A. L.; Riesenfeld, F. C. Gas Purification; Gulf Publishing Company: Houston, 1985. (6) Evaluation of Solid Sorbents as a Retrofit Technology for CO2 Capture from Coal-fired Power Plants, Technical Report; DENT0005649-05649FR01; NETL, ADA Environmental Solutions, 2011. (7) Pirngruber, G. D.; Guillou, F.; Gomez, A.; Clausse, M. A Theoretical Analysis of the Energy Consumption of Post-combustion CO2 Capture Process by Temperature Swing Adsorption using Solid Sorbents. Int. J. Greenhouse Gas Control 2013, 14, 74−83. (8) Liu, H.; Liu, B.; Lin, L.; Chen, G.; Wu, Y.; Wang, J.; Gao, X.; Lv, Y.; Pan, Y.; Zhang, X.; Zhang, X.; Yang, L.; Sun, C.; Smit, B.; Wang, W. A Hybrid Absorption-adsorption Method to Efficiently Capture Carbon. Nat. Commun. 2014, 5, 5147. (9) Lively, R. P.; Chance, R. R.; Koros, W. J. Enabling Low-cost CO2 Capture via Heat Integration. Ind. Eng. Chem. Res. 2010, 49, 7550− 7562. (10) Yang, W. Handbook of Fluidization and Fluid-particle Systems; Marcell Dekker: New York, NY, 2003. (11) Kunii, D.; Levenspiel, O. Fluidization Engineering; ButterworthHeinemann: Newton, MA, 1991. (12) Raksajati, A.; Ho, M. T.; Wiley, D. E. Reducing the Cost of CO2 Capture from Flue Gases using Aqueous Chemical Absorption. Ind. Eng. Chem. Res. 2013, 52, 16887−16901. (13) Raksajati, A.; Ho, M. T.; Wiley, D. E. Understanding the Impact of Process Design on the Cost of CO2 Capture for Precipitating Solvent Absorption. Ind. Eng. Chem. Res. 2016, 55, 1980−1994. (14) Engineering Data Book, FPS version, 12th ed.; Gas Processors Suppliers Association: Tulsa, OK, 2004. (15) Li, Y. Hydrodynamics. In Advances in Chemical Engineering Vol. 20 − Fast Fluidization; Kwauk, M., Ed.; Academics Press: San Diego, CA, 1994. (16) Fan, L. S., Zhou, C. Principle of Gas-Solid Flows; Cambridge University Press: New York, NY, 1998. (17) Reddy, S.; Johnson, D.; Gilmartin, J. Fluor’s Econamine FG PlusSM Technology for CO2 Capture at Coal-fired Power Plants. Presented at Power Plant Air Pollutant Control Mega Symposium, Baltimore, MD, August 2008.

AUTHOR INFORMATION

Corresponding Author

*D. E. Wiley. Tel.: +61 2 9351 2886. E-mail: dianne.wiley@ sydney.edu.au. ORCID

Anggit Raksajati: 0000-0001-5274-056X Dianne E. Wiley: 0000-0002-4655-0732 1619


Article

Industrial & Engineering Chemistry Research (18) www.solexthermal.com. (19) Yang, W.; Hoffman, J. Exploratory Design Study on Reactor Configurations for Carbon Dioxide Capture from Conventional Power Plants Employing Regenerable Solid Sorbents. Ind. Eng. Chem. Res. 2009, 48, 341−351. (20) Jordison, N. Solex Thermal Science. Personal communication, November 2014. (21) Oyenekan, B. A.; Rochelle, G. T. Energy Performance of Stripper Configurations for CO2 capture by Aqueous Amines. Ind. Eng. Chem. Res. 2006, 45, 2457−2464. (22) Oyenekan, B. A.; Rochelle, G. T. Rate modelling of CO2 stripping from potassium carbonate promoted by piperazine. Int. J. Greenhouse Gas Control 2009, 3, 121−132. (23) Ho, M.; Allinson, G. W.; Wiley, D. E. Reducing the Cost of CO2 Capture from Flue Gases using Membrane Technology. Ind. Eng. Chem. Res. 2008, 47, 1562−1568. (24) Du, K. H.; Park, H. B.; Dal-Cin, M. M.; Guiver, M. D. Advances in High Permebility Polymeric Membrane Materials for CO 2 Separations. Energy Environ. Sci. 2012, 5, 7306−7322. (25) Pera-Titus, M. Porous Inorganic Membranes for CO2 Capture: Present and Prospects. Chem. Rev. 2014, 114, 1413−1492.

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Techno-economic Evaluation of CO2 Capture from Flue Gases Using

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