Moving bed adsorption process based on a PEI-silica sorbent for CO2 capture

International Journal of Greenhouse Gas Control 67 (2017) 10–19

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Moving bed adsorption process based on a PEI-silica sorbent for CO2 capture ⁎

MARK

Wonho Jung, Junhyung Park, Kwang Soon Lee

Department of Chemical and Biomolecular Engineering, Sogang University, 35 Baekbeom-ro, Mapo-gu, Seoul 04107, Republic of Korea

A R T I C L E I N F O

A B S T R A C T

Keywords: CO2 capture Moving bed process Heat integration Amine-functionalized solid sorbent

A polyethylenimine (PEI)-silica adsorbent-based true moving bed adsorption process with a heat integration scheme was investigated as an energy conservative CO2 capture process. The process was configured using four separate beds with internal plate heat exchangers for adsorption, cooling, desorption and heating. An in-house steady-state simulator was constructed to analyze process behaviors, which focused on the energy demand and bed size under various operating conditions. A recently developed amine-functionalized adsorbent, 0.37EB-PEI, which uses silica fume as a support, was utilized in the process.

1. Introduction Carbon capture and sequestration (CCS) is regarded as one of the essential measures to mitigate atmospheric CO2 concentration below a catastrophic level by reducing CO2 emissions from both power and large industrial plants (D'Alessandro et al., 2010; Haszeldine, 2009; Karl and Trenberth, 2003; Rochelle, 2009; Samanta et al., 2011; SernaGuerrero et al., 2008). A post-combustion CC process based on amine scrubbing is technically mature, and CCS plants that treat 100–250 MW scale flue gas emissions from coal-fired power plants have been constructed and are being demonstrated at enhanced oil recovery sites (Rochelle, 2009). However, the overall CC cost needs to be reduced, and the environmental effects of amine degradation products must be further investigated before wide-scale adoption and deployment of the process (Brennecke and Gurkan, 2010; Gouedard et al., 2012; Lepaumier et al., 2009; Nguyen et al., 2010; Wang et al., 2011). Energy demand is one of the key factors that determine the economic feasibility of a CC process. Solid sorbent CC technology has been studied as a viable and cost-effective option that may replace absorption-based amine scrubbing technology (Chaikittisilp et al., 2011; D'Alessandro et al., 2010). The most important advantage of the solid sorbent process over the absorption process is that latent heat for water evaporation is not required (Chaikittisilp et al., 2011; Rochelle, 2009). However, the solid sorbent process still has several hurdles to overcome before it can be commercially viable. The sorbent material requires further improvement to retain a larger cyclic capacity, a faster adsorption rate, and higher thermal, chemical, and mechanical stabilities (Alptekin et al., 2012; D'Alessandro et al., 2010; Serna-Guerrero et al., 2008; Tarka et al., 2006; Wang et al., 2011; Yang et al., 2008). Additionally, there are several issues in the process design. The reactor

⁎

Corresponding author. E-mail address: [email protected] (K.S. Lee).

http://dx.doi.org/10.1016/j.ijggc.2017.10.004 Received 12 July 2017; Received in revised form 2 October 2017; Accepted 15 October 2017 1750-5836/ © 2017 Elsevier Ltd. All rights reserved.

type must be optimally determined among the fixed bed, moving bed (MB), fluidized bed, and their variants (Tarka Jr et al., 2006). Irrespective of the reactor type, the effective sensible heat recovery from the hot sorbent stream is critical for high energy efficiency, but the heat exchanger design for solid particles has inherent difficulties, including the low heat transfer capability and limited transportability of solid particles. A solid sorbent process design requires the contemplation and integration of various aspects. However, from the viewpoint of energy efficiency, the MB process is the definite choice since CO2 loading in exiting adsorbent particles from the adsorption bed can be maximized by counter-current contact between gas and solid streams. This minimizes the solid circulation rate, Fs, and the sensible heat demand. The MB process has been researched by several groups for this reason (Alptekin et al., 2012; Hornbostel, 2018; Kim et al., 2013; Yang and Hoffman, 2008). Yang and Hoffman (2008) numerically investigated a solid sorbent process using an MB regenerator. They assumed a hypothetical sorbent and proposed to use a Solex®-type plate heat exchanger bed for the thermal energy supply to the MB regenerator. Alptekin et al. (2012) developed an MB process based on a physical adsorbent and operated using a vacuum swing adsorption process. Heat exchangers were not installed in the beds. Instead, the heat of adsorption was removed by exiting solid and gas streams, and regeneration heat was conveyed by directly injected steam, which also enhances CO2 desorption by decreasing the CO2 partial pressure in the bed. The process concept was experimentally proved in a bench-scale plant. SRI (2018) also studied a physical adsorbent-based MB technology in a 1 t-CO2/day-scale pilot process. The adsorber, pre-heater, desorber, dehydrator, and cooler were vertically positioned from top to bottom. Steam was used as both a

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Nomenclature A C F ΔHCO2 L N Nu P Pr Q Qflue Re Sc T U Wi W a b be c dp k kb ke kf q qe qm sa

sd u Δx y

Cross-sectional area (m2) Molar concentration of the gas phase (mol/m3) Mass flow rate (ton/h) Heat of adsorption (kJ/mol) Bed height (m) The number of heat exchanger plates Nusselt number, hL/k Pressure (kPa) Prandtl number, cμ/k Consumed thermal energy (kJ/t-CO2) Mass flow rate of flue gas (ton/h) Reynolds number, ρudp/μ Schmidt number, μ/ρD Temperature (K) Overall heat transfer coefficient (W/m2·K) Equivalent work (kWh/t-CO2) Width of heat exchanger plates (m) Heat transfer area between the gas and solid phases per unit volume (m2/m3) Half distance between plates (m) Equilibrium constant (bar−1) Specific heat capacity (J/kg K) Particle diameter (m) Thermal conductivity (J/m s K) Backward reaction constant (sec−1) Effective thermal conductivity of the sorbent bed (J/m sec K) Forward reaction constant (sec−1bar−1) Amount of CO2 adsorption (kg-CO2/kg-sorbent) Amount of CO2 adsorption at equilibrium (kg-CO2/ kg-sorbent) Maximum possible adsorbed quantity of CO2 (kg-CO2/ kg-sorbent) Temperature-dependent parameter at adsorption

Temperature-dependent parameter at desorption Interstitial velocity (m/sec) Interval of stacked solid sorbent (m) Mole fraction of CO2

Greek letters β ρ ε κ μ η γ

Inter-phase drag coefficient (kg/m3 sec) Density (kg/m3) Bed porosity Conveyer efficiency Viscosity (kPa sec) Efficiency Heat capacity ratio, cp/cv

Subscripts ADB CB CW DES DSB HB fin g h hs in ini liq p s sen sg stm top

Adsorption bed Cooling bed Cooling water CO2 desorption Desorption bed Heating bed Final Gas Heating or cooling medium Between medium and solid Inlet Initial Liquefaction Particle Solid Sensible heat Between solid and gas Steam Top

2012; Sayari and Belmabkhout, 2010; Serna-Guerrero et al., 2008; Tarka Jr et al., 2006; Wang et al., 2011). The high heat of reaction, which can be regarded as a drawback, actually means a low desorption temperature according to the Van’t Hoff equation (Samanta et al., 2011). One of the disadvantages of the amine-functionalized adsorbent is the lack of long-term stability under repeated adsorption-desorption cycles. Recently, Choi et al. (2016) developed an amine-functionalized adsorbent, 0.37EB-PEI, to have long-term stability by partially epoxidizing PEI. In this research, we investigated a heat-integrated MB process for CO2 capture based on an amine-functionalized solid sorbent. The process is composed of four beds for CO2 adsorption, solid heating, CO2 adsorption, and solid cooling. The Solex®-type heat exchanger bed was assumed for all beds for heat exchange. The adsorbent 0.37EB-PEI with a 4 mm particle diameter was used in the process. Rigorous steady-state mass and energy balance equations were established for the MB process with adsorption and desorption kinetic models investigated by Jung et al. (2017). For energy demand assessment, energy models for the blower, compressor, and solid transportation were introduced together with thermal energy requirements. The total energy demand for 80% of the CO2 recovery from a flue gas with 15% CO2 was computed for various adsorption and desorption temperatures. For each set of adsorption/desorption temperatures, cooling water flow rates were optimized for maximum heat recovery. Finally, a sensitivity analysis was performed to investigate the effects of various process parameters on the total energy demand.

heating medium in the pre-heater and desorption bed and as a drying medium in the dehydrator. Air was used as a cooling medium in the cooler. Okumura et al. (2013) tested a 10 t-CO2/day-scale MB process. They used an amine-functionalized solid sorbent that can desorb CO2 at low temperatures. The process consists of adsorption, desorption, and drying beds. Pre-heated air was used to evaporate the water in the wet solid sorbent in the drying bed. Low-temperature desorption reduced the thermal burden of the process. The abovementioned solid sorbent processes did not consider sensible heat recovery from the hot solid stream from the desorber, although sensible heat consumes a large portion of the total regeneration energy. Son et al. (2014) proposed a heat-integrated MB process that assumed Solex®-type heat exchanger beds for both the adsorber and desorber, which improved the results from the study in Kim et al. (2013). Son and co-workers used zeolite 13X as the adsorbent and vacuum-temperature swing operation. Water was used as the medium that transferred the heat evolution from the adsorber to the desorber. The proposed heat integration scheme attempted substantial energy savings. There are many different solid sorbents that can be used for CC (Chaikittisilp et al., 2011; D'Alessandro et al., 2010; Fauth et al., 2012; Samanta et al., 2011). To date, physical adsorbents are still being investigated to benefit from the low heat of adsorption. However, the recent trend involves the amine-functionalized adsorbent that can potentially provide high CO2 selectivity, a fast adsorption rate, and high adsorption capacity (Bollini et al., 2012; Ebner et al., 2011; Fauth et al.,

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2. Process description 2.1. Moving bed process with heat integration Fig. 1(a) shows a schematic diagram of the proposed heat-integrated MB adsorption process. The process is composed of four separate beds that are operated under atmospheric pressure. Sorbent particles flow downward through the beds, while the flue gas flows upward in the adsorption bed (ADB). The desorbed CO2 gas is removed from both the top of the desorption bed (DSB) and heating bed (HB). Cooling water (CW) flows in the reverse direction of the solid stream. CW recovers the heat of adsorption in the ADB and the sensible heat of the sorbent particles in the cooling bed (CB) and delivers the heat to the sorbent in the heating bed (HB). A separate heating medium, steam, is used in the DSB to produce CO2 gas. Steam supplies heat to sorbent particles indirectly using an embedded plate heat exchanger. The CB can be integrated with the ADB in a single bed, as in Kim et al. (2013). In this case, the exhaust gas temperature becomes higher than in Fig. 1. These results not only demonstrate additional heat loss through the hot exhaust gas but also unnecessary desorption of CO2 from the hot sorbent stream from the DSB. The sorbent in the DSB is under a 100% CO2 environment and preserves a small quantity of CO2, even under high desorption temperatures. Desorption is induced by the low CO2-content exhaust gas, which decreases the CO2 recovery rate. The inlet gas to the CB is to maintain the bed pressure at 1 bar. Without this gas supply, vacuum can be developed during sorbent cooling. Determination of CO2 composition in the gas is discussed in Section 5.1. Because it is advantageous to maintain a low ADB temperature, a sufficient amount of CW is supplied to the ADB. The CW outlet stream from the ADB is split into two so that an optimum amount of the CW is transferred to the CB and HB for heat integration. The CW split ratio is an important variable to adjust for maximum heat recovery. Fig. 1(b) presents a sketch of temperature profiles along the beds. The sensible heat exchange between the CB and HB is the key part of the heat integration, which is the same as in the solvent absorption process. Solex®-type vessels with multiple heat exchanger panels were used for all beds for efficient heat transfer between the solid particles and heat transfer media. Fig. 2 shows the reactor shape and notations associated with vessel dimensions.

Fig. 2. Reactor structure and notations.

2.2. Solid sorbent The solid sorbent used in this study is a PEI-silica sorbent, 0.37EBPEI, developed by Choi et al. (2016). It is a PEI-functionalized silica sorbent. Unlike other sorbents of the same type, PEI is partially epoxidized to enhance the chemical stability, especially against oxidative degradation. Additionally, a mesoporous silica fume is used as the support to enhance the mass transfer rate. The sorbent has a large working capacity (2.2 mmol g−1) and long-term stability under practical temperature-swing operating conditions. Major characteristics of the sorbent are presented in Table 1. Adsorption isotherm and adsorption/desorption kinetic equations for the sorbent were obtained by Jung et al. (2017). The adsorption equilibrium follows the Langmuir isotherm equation:

qe qm

=

be PCO2 1 + be PCO2

(1)

where PCO2 [bar] is the CO2 partial pressure; be = kf/kb [bar−1]; kf and

Fig. 1. Heat-integrated MB adsorption process: (a) process scheme and (b) prospective temperature profiles.

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and g denote the solid sorbent and gas phases, respectively. The component mass balance for CO2 in the gas phase can be written as follows:

Table 1 Characteristics of the solid sorbent, 0.37EB-PEI (Choi et al., 2016). Pore size Pore volume BET surface area PEI avg molecular weight

20–100 nm 0.37 cm3/g 43 m2/g 1200

Bulk density, ρs Heat of adsorptiona, −ΔHCO2 Bed porosity, ε Nitrogen contents

800 kg/m3 63 kJ/mol 0.27 6.7 mol/kg

d (ug y )

− us

dq =r dz

d (ug Tg ) dz

=

(2)

and T denotes the temperature in K. The kinetic equations were suggested as follows: −sa (q / qm)

(7)

(8)

Heat balance equations can be established as follows:

kb are defined in Eqs. (3) and (4); qe [kg-CO2/kg-sorbent] is the amount of CO2 adsorption under an equilibrium state:

⎧e r = 0.66 × ⎨ ⎩

ρs 1 − ε r C ε

where y represents the mole fraction of CO2 in the gas phase. The component mass balance for CO2 in the solid phase is as follows:

Molar ratio of 1,2-epoxybutane to the nitrogen content in PEI initially used for the reaction is 0.37. a Measured using a DSC.

0.1253 qm [kg-CO2 /kg-sorbent] = 1 + exp(0.06844(T − 353))

=−

dz

Usg a ερg cg

(Ts − Tg )

(9)

Usg a dTs 1 Uhs (Tg − Ts ) + (Th − Ts ) = r (−ΔHco2) + dz cs (1 − ε ) ρs cs 2b (1 − ε ) ρs cs

us

(10)

(kf (qe − q) PCO2 + kb (qe − q) ) when qe ≥ q e−sd kb (qe − q)/ PCO2 when qe < q

uh

(3)

where r [kg-CO2/kg-sorbent sec] represents the reaction rate; and kf [bar−1 sec−1] and kb [sec−1] denote forward and backward reaction rate constants, respectively, with

kf = 1.88 × 10−7e (−28.6/ Rg T ), kb = 5.19 × 10−11e (−82.6/ Rg T )

dTh U W (N + 1) (Ts − Th) = − hs dz ρh ch A

(11)

Eqs. (9), (10) and (11) represent the heat balance for the gas, sorbent, and liquid (CW) phases, respectively. Here, a = 6/dp [m2/m3] means the heat transfer area between gas and spherical solid particles per unit volume; U [W/m2 K], c [kJ/kg K], and -ΔHCO2 [kJ/mol] denote the heat transfer coefficient, heat capacity, and heat of adsorption, respectively; b [m], W [m], N, and A [m2] are variables defined in Fig. 2 and the subscript h is defined as the heat transfer medium. Because the DSB used steam as the heating medium, Eq. (11) was not applied for the DSB. The detailed formula for U is given in Table 2. The following is the momentum balance equation (Gidaspow, 1994):

(4)

sa and sd are determined as follows:

sa = −8.357 × 10−4T 2 + 0.6444 × T − 115.8 for 313 K ≤ T ≤ 393 K sd = −6.071 × 10−4T 2 + 0.5208 × T − 102.2 for 313 K ≤ T ≤ 413 K (5) In Eq. (4), Rg is the gas constant. The kinetic equations were experimentally determined for sorbent particles with an average diameter of 100 μm. In the MB, sorbent particles with average diameter of 4 mm was assumed (see Table 3) to increase the gas velocity without causing fluidization (Wen and Yu, 1966). In this case, the macropore diffusion increases the mass transfer resistance and curtails the apparent reaction rate. The effectiveness factor of 0.66 in Eq. (3) was obtained by approximating the reaction rate equation to a first-order model and calculating the Thiele modulus (Fogler, 1999).

d (ρg ug ug ) dz

150(1 − ε )2μg 1.75ρg ug (1 − ε ) ⎞ dP + − ⎛⎜ ⎟ (us − u g) 2 dz ε dp εdp ⎠ ⎝ + ρg g (12)

=−

where g [m/sec2] is the gravitational acceleration. Eq. (12) is used to estimate the pressure drop in the bed. 3.2. Energy demand models The process requires not only the thermal energy for CO2 desorption but also the mechanical energies for the blower operation, solid transport, and CO2 compression. The energy demand models for the operation of these mechanical devices are given in the following: The blower energy for flue gas transportation was calculated using Eq. (13) (Smith, 1975).

3. Mathematical model of the process 3.1. Reactor performance equations For mathematical modeling of the beds, the following major assumptions were introduced:

Wblower [kWh/t-CO2] =

i The lateral concentration and temperature distributions are negligible. ii The heat loss from the bed is negligible. iii The ideal gas law is assumed in the bulk gas phase. iv The reactor vessel is a hexahedron shape, which indicates that the effects of the lower funnel zone in the Solex® vessel on the bed performance are neglected.

Qflue Pin ⎛ ⎛ PADBin ⎞γ ⎞ −1⎟ ⎜ 17.4ηFs q ⎝ ⎝ Patm ⎠ ⎠ ⎜

⎟

(13)

where γ = cp/cv; cp and cv are isobaric and isochoric heat capacities of the flue gas, respectively; η is the blower efficiency assumed to be 0.75 Table 2 Equations for heat transfer coefficients. Heat transfer coefficients

The total mass balance in the gas phase is described as follows:

d (ug C ) dz

1−ε = −ρs r ε

1 Ui

=

1 hs

+

Δx k

+

1 , hi

hs = − (2k e )/ b ((∂2Ts / ∂x 2)/(Th − Tsav g )) (Kim et al., 2013)

ks / k g = 0.0935 + 0.9065 [2/(1 − k g /ks )][(ks )/(ks − k g )(In ks /k g ) − 1] + 0.11 Rep Prp (-

(6)

Kunii and Smith, 1960; Swift, 1966) Nuh = 0.023 Re0.8 Pr nh, n = 0.4 for heating, 0.3 for cooling (McCabe et al., 1993) h

where C [mol/m3], u [m/s], ρ [kg/m3], and ε represent the total gas concentration, velocity, bulk density, and bed porosity, respectively; z [m] represents the upward axial distance of the bed; and subscripts s

Nug = 2.0 + 1.1 Pr1/3 Re0.6 g g (Ruthven, 1984)

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5. Results and discussion

in this study; Qflue [ton/h], Patm [bar], and PADBin [bar] represent the flow rate of the flue gas, 1 bar, and inlet pressure of the ADB, respectively; and Fs [ton/h] indicates the mass flow rate of sorbent particles. The CO2 liquefaction energy was estimated using the following equation:

Pfin ⎞ Wliq [kWh/t-CO2] = 4.572 ln ⎛ − 4.096 ⎝ Pin ⎠ ⎜

5.1. Internal profiles along the beds Representative temperatures and CO2 absorption profiles along four beds are exhibited in Fig. 3. The profiles were obtained when Fs, Ts,ADBtop, TDSB, FCW, ADB, and FCW, CB are 11.2 ton/h, 323 K, 393 K, 1.96 ton/h, and 0.18 ton/h, respectively. Other fixed conditions are given in Table 3. The minimum temperature approach (MTA) between the solid and CW streams was assumed to be 20 K. The temperature difference between the gas and solid at all axial positions is less than 0.5 K. The CW at 293 K that enters the ADB increases its temperature by taking the heat of adsorption. After exiting the ADB, a large portion of the CW is returned to the cooling tower, and the remaining small portion enters the CB to recover the sensible heat in the sorbent particles at as high a temperature as possible. The high-temperature CW transfers its thermal energy to solid particles in the HB and is returned to the cooling tower. The additional thermal energy required in the DSB is provided by steam. The CB can be operated under different gas-phase conditions. After several attempts, we chose to feed 0.3 ton/h of air with 0.5% CO2 at 323 K, that can be produced by mixing the ADB exhaust gas and air. This flow induces a small amount of CO2 adsorption in the bottom region of the bed and desorption in the upper region, resulting in nearly zero net adsorption or desorption. Generally, a gas flow with a CO2 content larger than 0.5% results in a net CO2 adsorption, and vice versa. Solid particles from the ADB are transferred to the HB increasing the temperature as they move downward. CO2 begins to desorb and flows upward as the solid temperature increases, which induces CO2 adsorption in the upper region of the bed. This phenomenon generates concave profiles in the solid temperature and CO2 adsorption in the upper region of the HB, as shown in Fig. 3. The heated solid particles are moved to the DSB. Steam is used to supply the desorption energy. Unlike in the HB, only CO2 desorption occurs due to the high bed temperature. At the bottom of the DSB, CO2 adsorption reaches a near-equilibrium state.

⎟

(14)

which was obtained using regression of simulation results for four to eight stage compressors (Van Wagener and Rochelle, 2011). For the MB process, the compressor inlet pressure, Pin, is fixed at 1 bar. Wliq is obtained as 118 kWh/t-CO2 when the CO2 is compressed up to Pfin = 150 bar at 313 K. Solid conveying from the ADB bottom to the HB top requires mechanical energy (Munson et al., 1990). The conveying energy can be expressed as follows:

Wconvey [kWh/t-CO2] = κ

gHtotal q

(15)

where Htotal [m] denotes the total height from the ADB bottom to the HB top; and κ is a constant that accounts for the conveyer efficiency and additional horizontal conveying. For κ up to 15, κ gHtotoal/q is less than 10% of Wblower, and Wconvey is neglected in this study. The conversion of the thermal energy demand (by steam) to equivalent mechanical energy was performed using the following Carnot engine equation:

Wstm [kWh/t-CO2] = 0.75QDSB

Tstm − Tsink 1 Tstm 3.6

(16)

where QDSB [kJ/t-CO2] is the thermal energy demand in the DSB. Assuming that steam at Tstm is drawn from the power plant, Wstm represents the loss of electric energy by not using steam for power generation. The conversion efficiency was assumed to be 0.75 (Oyenekan and Rochelle, 2007), and Tsink was assumed to be 313 K.

5.2. Behavioral analysis of the moving bed process with heat integration 4. Simulation conditions and numerical method The desorption temperature mostly influences the process design and operation. Not only the regeneration energy but also the minimum required bed heightss depend on the desorption temperature. Fig. 4 shows the effects of the desorption temperature, TDSB, on individual energy demands and bed heights when Fs, Ts,ADB-top, and MTA are fixed at 11.2 ton/h, 313 K, and 20 K, respectively. The steam temperature is assumed to be TDSB + MTA.

In Table 3, bed parameters, sorbent properties, flue gas, and heat transfer media, and fixed operating conditions are listed. We assumed a process for handling 6.12 ton/h (about 4000 Nm3/h) of flue gas (1 MW equivalent) at 313 K with a CO2 mole fraction of 0.15. Values of W and N were determined so that the superficial gas velocity is 0.3 m/s. The gas velocity is approximately one-third of that in the turbulent fluidized bed for CC (Kim et al., 2014). Therefore, the active cross-sectional area of the ADB becomes coarsely three times that of the fluidized bed adsorber. The particle diameter, dp, is determined so that fluidization does not occur at the prescribed gas velocity (Wen and Yu, 1966, 2013). The space between heat exchanger plates, 2b, is intuitively determined by considering the trade-off between the heat transfer and design complexity/equipment cost. The bed heights were not fixed but were adjusted depending on the operating condition, such that 80% of CO2 recovery can be achieved. The solid circulation rate, Fs, is a dependent variable and is calculated as the minimum value that can carry 80% of CO2 in the flue gas to the HB. The bed model is described by a set of ordinary differential equations (ODEs), but the boundary conditions (BCs) are split by countercurrent flows of gas/CW and solid particles. For example, in the ADB, BCs for the gas and CW equations are defined at z = 0, whereas the BCs for the sorbent equations are given at z = HADB. A shooting algorithm is employed to solve the ODEs (Desai and Abel, 1972). The ODEs are integrated using the Runge-Kutta fourth-order method, and a process simulator is constructed using MATLAB programming software.

Table 3 Model parameters used in this study. MB parameters A (m2) W (m) N b (m)

3.992 1.998 81 0.025

ε Usg (W/m2 K) Uhg (W/m2 K)

0.35 60 250 [CW, steam]

Physical properties cg (kJ/kg K) ch (kJ/kg K) cs (kJ/kg K) dp (m)

1.07 4.18 1.49 0.004

ρh (kg/m3) ρs (kg/m3) kg (J/m K sec) ks (J/m K sec)

1000 800 0.026 0.275

Operating conditions Patm (bar) Pfin (bar) Tsink (K) Tcw,ini (K)* CO2 recovery rate (%)

1 150 313 293 80

η γ Qflue (ton/h) yCO2,ini*

0.75 0.283 6.12 0.15

* Subscript ini refers to a value at the ADB bottom, z = 0.

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Fig. 3. Internal profiles along four beds: (a) temperature, (b) CO2 adsorption. Fs = 11.2 ton/h, Ts,ADB-top = 323 K, TDSB = 393 K, FCW, MTA = 20 K.

ADB

= 78.4 ton/h, FCW,

CB

= 7.2 ton/h,

minimum required bed heights. The energy demand monotonically decreases with a decrease in Ts,ADB-top. Since 313 K is the minimum value of Ts,ADB-top (by assumption), Ts,ADB-top is not decreased any further. The total energy demand reached 238 kWh/t-CO2 under this condition. As Ts,ADB-top decreases, CO2 adsorption is enhanced, and the minimum height of the ADB for 80% CO2 recovery is lowered. However, a larger drop in the solid temperature occurs across the CB, which requires an extended height. From the abovementioned simulation study, Fs = 11.2 ton/h, Ts,ADBtop = 313 K, TDSB = 382 K, MTA = 20 K, and fixed operation conditions in Table 3 are chosen as the nominal operating conditions.

Fig. 4(a) displays the four major energy demand items and their sum in terms of the equivalent work. Desorption energy, WDES, indicates the chemical energy for CO2 desorption in the DSB. The desorption energy is constant in terms of kJ/t-CO2 but increased in kWh/t-CO2 with an increase in TDSB. The sensible energy, Wsen, refers to the energy required to change the sorbent state from TADB to TDSB minus the desorption energy when there is no heat integration. When heat integration is incorporated, Wsen indicates the energy used to only increase the sorbent temperature in the DSB. With the proposed heat integration scheme, approximately 30–40 kWh/t-CO2 can be saved. High desorption temperature causes low lean loading of the solid sorbent after desorption, hence enhances adsorption rate between CO2 and solid sorbent. This shortens the height of the adsorption bed for 80% CO2 recovery (see LADB in Fig. 4(b)) and also decreases the blower energy. The total energy demand reaches a minimum value of 238 kWh/tCO2 at TDSB = 382 K. Fig. 5 shows the effect of Ts,ADB-top on the total energy demand and

5.3. Energy performance of the moving bed process The energy demand assessment described in the previous section did not consider the effect of water on the flue gas. The flue gas contains approximately 7 vol% water, and the PEI-silica-based sorbent is known 15

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Fig. 4. (a) Energy demand and (b) minimum required bed height as a function of the desorption temperature. Fs = 11.2 ton/h, Ts,ADB-top = 313 K, MTA = 20 K.

Fig. 5. Energy demand and minimum required bed height as a function of Ts,ADB-top when Fs = 11.2 ton/h, TDSB = 382 K, and MTA = 20 K.

Fig. 6. (a) Total energy demand and (b) minimum total bed height of the MB process with and without heat integration under nominal operating conditions; TDSB = 383 K, Ts,ADB= 313 K, Fs = 11.2 ton/h, FCW,ADB = 368.8 ton/h, and FCW,CB = 7.6 ton/h.

top

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Fig. 7. Total energy demand of the MB, monoethanolamine (MEA)-based absorption, and piperazine (PZ)-based absorption processes; (a) thermal energy demand and (b) total energy demand.

show reverse tendencies. Overall, improving both the heat transfer (by decreasing MTA and increasing U) and sorbent properties (by increasing qm and decreasing −ΔHCO2 and cs) are key to the energy demand reduction. If all these parameters are improved by 10% (MTA by 5 K), the total energy savings would amount to approximately 23 kWh/t-CO2, and total energy demand can reach approximately 224 kWh/t-CO2. In this case, the minimum total bed height is slightly decreased from the nominal value. Improving the heat transfer capability of the MB process is just as influential or more compared with improving the sorbent properties. Improving the adsorption rate constant, kf, has a minor impact on the energy savings in the MB process.

to adsorb water (Ng et al., 2001). Water enhances the maximum capacity of PEI-silica-based sorbent in the ADB by forming bicarbonate (Sayari and Belmabkhout, 2010). In the DSB, water desorption requires additional regeneration energy; however, it facilitates CO2 desorption by lowing the CO2 partial pressure on the other hand. In this study, we only considered the adverse effect of water on the energy demand and latent heat in the DSB. According to the isotherm suggested by Ng et al. (2001), we assumed that the cyclic capacity of water is 3 g-H2O/kg-sorbent. This capacity increases the regeneration energy by 8 kWh/t-CO2 at 383 K of TDSB. Fig. 6 displays the total energy demand, including water vapor desorption, in the DSB and the minimum total bed height of the MB process with and without heat integration at nominal operating conditions. Heat integration enables a savings of 35.1 kWh/t-CO2, which results in 246.9 kWh/t-CO2 of the total energy demand. At the sacrifice of energy savings, the minimum bed height is increased by 57 cm. In Fig. 7, the total energy demand of the proposed MB process is compared with those of two well-known absorption processes: the 30 wt% MEA process and 7 mol PZ process. Energy demands of the MEA and PZ processes were acquired from (Rochelle et al., 2011). In the thermal energy demand comparison in Fig. 7(a), only the sensible energy and desorption energy are considered in the MB process. The reboiler heat input is considered in the MEA and PZ processes. In the total equivalent work, all energy terms from the blower to liquefaction energies are included. MEA and PZ processes assume 2 and 10 bar of stripper pressure, respectively, and greatly conserve the liquefaction energy compared to the MB process. The proposed MB process requires far less energy than the MEA process but more energy than the PZ process.

6. Conclusions In this study, the MB process for CO2 capture based on a recently developed amine-functionalized silica sorbent named 0.37EB-PEI was proposed. The proposed MB process is composed of four beds for CO2 adsorption, sorbent heating, CO2 desorption, and sorbent cooling, through which sorbent particles are circulated in sequence. To curtail the energy demand of the process, a heat integration scheme was devised such that the sensible heat in the high temperature sorbent particles discharged from the DSB is recovered and used to heat the cold sorbent particles that exit the ADB. The process performance was evaluated using a numerical simulator constructed based on rigorous steady-state balance equations and an accurate kinetic model. The beds were assumed to have a hexahedron shape, and equally spaced heat exchanger panels were installed inside each bed. Water and steam (in the DSB) were used as heat transfer media. As a nominal case for performance evaluation, a flue gas containing 15% CO2 and 7% water vapor, sorbent particles with 4 mm diameter, a 0.3 m/s superficial gas velocity at the adsorber bed inlet, and a minimum temperature approach in the heat exchangers of 20 K were assumed. Other operating conditions were determined to minimize the total energy demand. As a result, the total energy demand of the MB process up to CO2 liquefaction is estimated to be 246.9 kWh/t-CO2, whereas typical values of the energy demand for absorption processes based on MEA and PZ are 290 and 220 kWh/t-CO2, respectively. The sum of four minimum required bed heights for the nominal case under these conditions is 3.38 m. When MTA is decreased to 15 K, and qm, −ΔHCO2, and cs are improved by 10%, the energy demand of the MB process can be reduced to approximately 224 kWh/t-CO2, which is comparable with that of the PZ-based absorption process. Under these conditions, the total bed height remains near 3.3 m, which is far shorter than the absorber column of the absorption process.

5.4. Sensitivity analysis Fig. 8 demonstrates how the energy demand and minimum total bed height vary when key process parameters are changed. All considered process parameters (except MTA and TDSB) change from −15 to +15% from their nominal values, whereas MTA and TDSB change from −10 to +10 K. Regarding the energy demand, the three most influential parameters are MTA, −ΔHCO2, and cs. When MTA is decreased to 15 K, approximately 10 kWh/t-CO2 of energy saving is possible. Conversely, the minimum total bed height is increased by approximately 0.45 m, a part of which is contributed by the ADB. The quantity of energy savings is the net value, which is the sensible heat recovery minus the additional blower energy. Decreases in −ΔHCO2 and cs and an increase in qm can yield an appreciable quantity of energy savings. The above three parameters are thermodynamic properties of the sorbent. Note that the increasing/decreasing tendency of the energy demand and total bed height are the same for all parameters, except for MTA and TDSB, which

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Fig. 8. Variations in the (a,c,e) energy demand and (b,d,f) total bed height as a function of the (a,b) MTA and TDSB, (c,d) U, b, cs, and ε, and (e,f) −ΔHCO2, qm, and kf.

Acknowledgments

The authors are grateful to Professor Minki Choi at KAIST for valuable information and discussion regarding the 0.37EB-PEI sorbent.

This work was supported by the “Human Resources Program in Energy Technology” of the Korea Institute of Energy Technology Evaluation and Planning (KETEP), granted financial resource from the Ministry of Trade, Industry & Energy, Republic of Korea (No. 20174010201150) and the Korea CCS R & D Center (Korea CCS 2020 Project) grant funded by the Korean government (Ministry of Science, ICT & Future Planning) in 2017 (KCRC-2014M1A8A1049261).

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