Mass transfer and kinetics study on combined CO2 and SO2 absorption using aqueous ammonia

International Journal of Greenhouse Gas Control 41 (2015) 60–67

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Mass transfer and kinetics study on combined CO2 and SO2 absorption using aqueous ammonia Guojie Qi, Shujuan Wang ∗ , Zhicheng Xu, Bo Zhao, Changhe Chen Department of Thermal Engineering, Key Laboratory for Thermal Science and Power Engineering of Ministry of Education, Beijing Key Laboratory for CO2 Utilization and Reduction Technology, Tsinghua University, Beijing10084, China

a r t i c l e

i n f o

Article history: Received 2 March 2015 Received in revised form 17 May 2015 Accepted 24 June 2015 Keywords: Aqueous ammonia CO2 SO2 Post-combustion Combined capture Mass transfer Kinetics

a b s t r a c t Combined capture of CO2 and SO2 using aqueous ammonia (NH3 ) is a good option to reduce capital and operational costs for post-combustion capture process. The mass transfer and kinetics of combined CO2 and SO2 absorption in aqueous NH3 solvent were determined at various SO2 concentrations, temperatures and NH3 concentrations using wetted wall column. A bi-liquid film assumption was proposed to analyze the mass transfer and kinetic characteristics by dividing the liquid film to SO2 instantaneous reaction region and CO2 pseudo-first-order reaction region. The high SO2 concentration significantly reduces the CO2 mass transfer coefficient in overall liquid phase and the SO2 instantaneous reaction region, but the high temperature and aqueous NH3 concentration benefit the CO2 mass transfer. In the CO2 pseudo-firstorder reaction region, the apparent and second-order reaction rate constants of CO2 absorption were calculated and the second-order reaction rate constant can match very well with the previous work. © 2015 Elsevier Ltd. All rights reserved.

1. Introduction Removal of acid gas impurities, such as sulfur dioxide (SO2 ) and carbon dioxide (CO2 ) from coal-fired flue gas by flue gas desulfurization (FGD) system and aqueous amine based absorption/regeneration process has been well-studied and proven over several decades (Rochelle, 2009; Lunt and Cunic, 2010) and will gain more significance in view of the stricter emission reduction targets expected in the near future (IPCC, 2007; Svendsen et al., 2011). However, both the FGD and CO2 capture processes are in principle inefficient and cost intensive (Black, 2010). One challenge is that the capital and operational costs are very high, particularly, the capital cost of the extra equipment and also energy penalty associated with the aqueous amine based CO2 capture process could significantly reduce a coal-fired plant’s output typically by around 30% (Carter, 2007). Another technical hurdle is that the flue gas from front FGD system normally contains tens of ppm SO2 , continually causing continuous amine solvent degradation in the CO2 capture process (Zhou et al., 2012).

∗ Corresponding author. Fax: +86 62770209. E-mail address: [email protected] (S. Wang). http://dx.doi.org/10.1016/j.ijggc.2015.06.023 1750-5836/© 2015 Elsevier Ltd. All rights reserved.

Aqueous ammonia (NH3 ), as an effective and economic solvent, absorbs CO2 rapidly, has a low regeneration energy, and has less corrosion and degradation issues, in particular, has the potential to capture SO2 and CO2 simultaneously (Liu et al., 2009). A series of publications related to the combined capture of CO2 and SO2 using aqueous NH3 have considered assessment of the feasibility of NH3 based combined CO2 and SO2 capture process development aimed at reducing the cost penalties (Ciferno et al., 2005; Choi et al., 2009; Kozak et al., 2009; Christopher et al., 2009; Qiu et al., 2011; Yu et al., 2011; Qi et al., 2014). Alstom has deployed an integrated chilled ammonia process (CAP) in which CO2 is absorbed by the ammoniated solvent in the absorber at low temperature and residual SO2 from the front FGD system is captured by the escaped NH3 from the washing column on the top of the absorber. This process may achieve 90% CO2 capture efficiency with lower than 2 MJ/kg CO2 regeneration heat requirement (Kozak et al., 2009). Powerspan has developed an aqueous NH3 based ECO2 process integrated with the ECO multi-pollutant control system. It is reported that the ECO system can remove SO2 , NOx, Hg, and PM particles, and the ECO2 system can effectively capture the CO2 and residual SO2 . The SO2 enriched NH3 is used to produce ammonium sulfate fertilizer and the CO2 enriched NH3 is regenerated in the stripper (McLarnon and Duncan, 2009). The US DOE reported that the aqueous NH3 based multi-pollutant control system can effectively capture CO2 , SO2 , NOx (oxidized) and Hg simultaneously and offer 14% lower COE than when the individual pollutants capture by aqueous NH3

G. Qi et al. / International Journal of Greenhouse Gas Control 41 (2015) 60–67

(Ciferno et al., 2005). Qi et al. (2013) also presented that the aqueous NH3 based combined capture process can remove CO2 and SO2 in one absorber and reduce the total cost by 17.4% compared to the individual FGD and CO2 capture system. Hence, aqueous NH3 based combined capture of CO2 and SO2 is a good option for development of cost-effective technologies. However, there is currently insufficient information on combined CO2 and SO2 absorption in aqueous NH3 in the public domain, especially on the mass transfer and kinetics of combined CO2 and SO2 absorption. Choi et al. (2009) reported that the aqueous 2-amino-2-methyl-1-propanol (AMP) and NH3 blend solvent can simultaneously absorb CO2 and SO2 . Aqueous NH3 , as an effective additive, can obviously enhance the mass transfer rate of CO2 . Qiu et al. (2011) carried out a preliminary experimental study on CO2 and SO2 simultaneous absorption in a packed column, and reported that the NH3 concentration has significant impact on CO2 and SO2 mass transfer coefficients. Qi et al. (2012) briefly presented the effect of SO2 concentration and loading on overall CO2 mass transfer coefficient. The overall CO2 mass transfer coefficient reduces by a half when the SO2 concentration increases from 0 to 4000 ppm, and the SO2 loading also has negative impact on CO2 mass transfer. In this study, the mass transfer and kinetics of combined CO2 and SO2 absorption in aqueous NH3 were determined at varied SO2 concentrations, temperatures and aqueous NH3 concentrations in a wetted wall column. A bi-liquid film assumption was applied to analyze the mass transfer and kinetics characteristics, especially in the liquid film, to understand the interaction mechanism of SO2 and CO2 , especially the impact of SO2 on CO2 absorption, within aqueous NH3 in more detail. This work should assist in optimizing the absorber design and specification in the combined capture process development. 2. Experimental apparatus and methods

The reactor pressure was controlled by a needle valve downstream of the gas outlet of the column set by an online pressure transducer. The solvent was circulated by peristaltic pump with a flow rate of 2–3 cm3 /s and preheated by the water bath and coil. An acid wash followed by a drying pipe was used to capture escaping NH3 and water vapor, respectively, before being supplied to the CO2 analyzer. The data points were collected when the temperature, pressure and CO2 and SO2 concentrations reached steady-state. Particularly, the outlet SO2 concentration is not determined in this study because of low outlet SO2 residual. Experiments were performed at varied inlet SO2 concentrations, temperatures and NH3 concentrations. The SO2 concentration was from 0 ppm to 4000 ppm. The experimental temperature range was from 20 to 80 ◦ C. The NH3 concentration varied from 0.579 mol/L to 4.012 mol/L. 2.3. Method The mass transfer process with gas–liquid chemical absorption can be depicted as below based on two-film theory (Whitman, 1923), where the total CO2 mass transfer resistance was assumed as the sum of gas and liquid side resistances as shown in Eq. (1) (Danckwerts, 1970): 1 1 1 = +  KG kg kG

KG =

2.2. Wetted wall column setup In this study, the wetted wall column (WWC) was applied to determine the mass transfer and kinetics of combined CO2 and SO2 absorption. The schematic of WWC is depicted in Fig. 1. The WWC is the same equipment used in our previous work (Liu et al., 2011, 2012). The gas–liquid wetted wall contactor is constructed from a stainless steel tube with effective height of 11 cm, outside diameter of 1.2 cm, and surface area of 41.45 cm2 . It is vertically enclosed in a jacketed glass chamber with an outside diameter of 31.0 cm, separated from a temperature-controlling water bath. Solvent flows up through the hollow tube and forms a uniform liquid film on the outer surface of column. The gas mixture of CO2 , SO2 and N2 feeds at the bottom of column and counter-currently contacts with the falling solvent. The reactions occurring across contact area give a CO2 concentration difference between the gas inlet and outlet. The CO2 concentration at the gas inlet and outlet was measured by an online calibrated IR CO2 analyzer. The gas flow rate was controlled by a series of calibrated mass flow controllers with a flow rate of 4 SLPM. The water bath with water circulation was applied to accurately control the temperature of gas and liquid.

(1)

where KG , overall gas phase mass transfer coefficient, mol/Pa/cm2 /s; kg , gas phase mass transfer coefficient,  , liquid phase mass transfer coefficient, mol/Pa/cm2 /s; kG 2 mol/Pa/cm /s. The reciprocals of the mass transfer coefficients represent the mass transfer resistances. The overall CO2 gas phase mass transfer coefficient KG can be calculated by the mass transfer flux and the driving force, as shown in Eq. (2):

2.1. Materials Aqueous NH3 concentrations were varied by diluting 25 wt% aqueous NH3 (Sinopharm Chemical Reagent Co.) with deionized water. The accurate concentration of aqueous NH3 solvent was measured by titration method with 0.1 mol/L H2 SO4 acid. High purity N2 (≥99.9%), CO2 (≥99.9%) and SO2 (≥99%) from Beijing Huayuan Gas Chemical Industry Co., were supplied to the experiments.

61

Flux ∗ PCO2 ,b − PCO

(2)

2

where Flux is mass transfer flux, mol/cm2 /s; PCO2 ,b is log mean average of operational gas bulk CO2 partial pressure in the wet∗ ted wall column, Pa. PCO is equilibrium CO2 partial pressure, Pa. In 2 ∗ this work, the PCO in aqueous ammonia solvent without CO2 load2 ing was assumed to be 0. The mass transfer flux can be determined by Eq. (3): Flux =

(VG × (CO2 %)in − VN2 /(1 − (CO2 %)in ) × (CO2 %)out ) 22.4 × A

(3)

where VG is total gas flow rate, L/min; VN2 is nitrogen gas flow rate, L/min; (CO2 %)in and (CO2 %)out are inlet and outlet CO2 concentrations, %; A is the reaction area of the wetted wall column, cm2 . The PCO2 ,b can be calculated by Eq. (4): PCO2 ,b =

PCO2 ,in − PCO2 ,out Ln(PCO2 ,in /PCO2 ,out )

(4)

where PCO2 ,in and PCO2 ,out are the operational CO2 partial pressures at gas inlet and outlet of the wetted wall column, respectively.  , In order to calculate the liquid phase mass transfer coefficient kG the correlation for gas phase mass transfer coefficient kg calculation was fitted by the SO2 mass transfer coefficient in 2 mol/L NaOH solvent (instantaneous reaction) with varied gas flow rate at room temperature according to the method described by Bishnoi (2000). The correlation is as shown in Eq. (5):

 Re × Sc × d 0.503

Sh = 6.709

h

(5)

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G. Qi et al. / International Journal of Greenhouse Gas Control 41 (2015) 60–67

Fig. 1. Overall experimental schematic of wetted wall column.

where Sherwood number Sh = RTkg dh /DCO2 ; Reynolds number Re = g Vg dh /g ; Schmidt number Sc = /D; d is the hydraulic diameter of the column and h is the length of the column; R is universal gas constant, 8.314 J/K/mol; DCO2 is diffusion coefficient of CO2 , m2 /s; g is density of gas, kg/m3 ; g is dynamic viscosity of gas, cP; v is kinematic viscosity of gas, m2 /s. 2.4. Reaction mechanism The mass transfer rate of CO2 and SO2 in aqueous NH3 is controlled by diffusion limitation and reaction rate between CO2 , SO2 and NH3 . Caplow (1968) and Derks and Versteeg (2009) reported that the CO2 mass transfer in aqueous NH3 is similar to that in primary amine, which can be explained according to the pseudofirst-order reaction mechanism. Based on the two-film theory, the NH3 concentration in liquid film is expected to be little different from that in the bulk liquid and there is little diffusion limitation when the reaction between CO2 and NH3 occurs. The overall reaction between CO2 and NH3 could be assumed to be a pseudofirst-order reaction dependent on the dissolved CO2 concentration in liquid film. Correspondingly, Qin et al. (2010) and Liu et al. (2012) reported that the CO2 mass transfer process in aqueous NH3 is determined by the reaction rate, and explained using the pseudofirst order-reaction mechanism. Choi et al. (2009) determined that the SO2 mass transfer in aqueous NH3 is following the instantaneous reaction mechanism. When the reaction rate is much higher than diffusion, the diffusion rate of active NH3 in liquid film is lower than its potential consumption, and thus, the mass transfer rate is limited by the diffusion rate of active NH3 . Hence, the SO2 and NH3 concentrations tend to be 0 at the gas–liquid interface area. In this work, a bi-liquid film assumption was proposed based on the two-film theory to describe the combined CO2 and SO2 mass transfer process in aqueous NH3 , in which the liquid film was

divided into a SO2 instantaneous reaction region and a CO2 pseudofirst-order reaction region, as shown in Fig. 2. In this assumption, CO2 and SO2 simultaneously diffuse to gas–liquid interface, primarily, SO2 reacts instantaneously with NH3 diffused from liquid film and form the SO2 instantaneous reaction region with a subinterface in the middle with SO2 and NH3 concentrations of 0. In this region, the CO2 mass transfer is inhibited by the selective absorption of SO2 and reduction of fresh NH3 . Then, the CO2 diffuses through the SO2 instantaneous reaction region and reacts rapidly with the aqueous NH3 and forms the CO2 pseudo-first-order reaction region, in which the NH3 concentration is slightly lower than that in the liquid bulk. Based on the bi-liquid film assumption, Eq. (1) could be described as: 1 1 1 1 = +  +  KG kg k G,1 k G,2

(6)

where 1/k G,1 is CO2 mass transfer resistance in the SO2 instantaneous reaction region, 1/k G,2 is CO2 mass transfer resistance in the CO2 pseudo-first-order reaction region. In Eq. (6), KG could be calculated according to mass transfer flux and driving force as Eq. (2) shown, kg could be determined by the correlation as Eq. (5). k G,1 represents the impact of SO2 on CO2 mass transfer, which could be obtained by the difference of liquid phase CO2 mass transfer coefficient between the absence and presence of SO2 . As discussed above, the CO2 mass transfer in aqueous NH3 could be explained by the pseudo-first-order reaction mechanism. The  mass transfer flux and liquid phase mass transfer coefficient (kG,2 ) in the CO2 pseudo-first-order reaction region could be calculated as: Flux =



kobs DCO2

PCO2 HCO2

(7)

G. Qi et al. / International Journal of Greenhouse Gas Control 41 (2015) 60–67

63

Fig. 2. Two film theory based bi-liquid film assumption.



kobs DCO2



kG,2 =

(8)

HCO2

where kobs is observed reaction rate constant, 1/s; DCO2 is liquid phase CO2 diffusion coefficient, m2 /s; PCO2 is log mean average of CO2 partial pressure, Pa; HCO2 is liquid phase CO2 Henry constant, kPa m3 /kmol. The reaction path of CO2 mass transfer could be described by the zwitterion mechanism proposed by Caplow (1968) and adopted by Qin et al. (2010) in the kinetics study on CO2 absorption in aqueous NH3 . It deems that the formation of carbamate is the mass transfer rate-limited step, NH3 and CO2 react and primarily form a zwitterion NH+ COO− , then the NH+ COO− is deprived of a proton H+ 3 3 by the bases (H2 O and NH3 ) and form carbamate in aqueous NH3 solvent. k2

CO2 + NH3 ←−NH+ COO− 3 →

(R1)

kB

NH+3 COO− + B → NH2 COO− + BH+ ←−

(R2)

k−B

+ CO2 + H2 O → HCO− 3 +H kOH−

CO2 + OH− → HCO− 3

(R3) (R4)

The reaction rate of CO2 and H2 O, (R3) is usually very slow compared to the reaction between CO2 and OH− (R4) when the solution pH value is greater than 9 (Astarita et al., 1981). In this work, the contribution of (R3) on mass transfer is negligible as in the literatures, the reaction rate of (R4) is described as (Pinsent et al., 1956): 2 −OH

13.635 − 2895 T

(10)

The overall reaction rate robs of CO2 absorption is given by robs = rCO2 −NH3 + rCO

(11)

− 2 −OH

where rCO2 −NH3 is the reaction rate of (R1) and (R2). The overall reaction rate constant kobs could be calculated by kobs =

robs [CO2 ]

(12)

Furthermore, the apparent reaction rate constant kapp of the carbamate formation from NH3 and CO2 ((R1) and (R2)) could be obtained by kapp = kobs − kOH− [OH− ]

(13)

kapp = k2 [NH3 ]

(14)

where [NH3 ] is the molecular NH3 concentration, mol/L.

In addition to produce carbamate, the dissolved CO2 also reacts with H2 O and OH− and produce HCO3 − , hence, the overall reaction rate is also contributed by the following reactions (Qin et al., 2010):

−

lg(kOH− ) =

The second order reaction rate constant k2 of carbamate formation is calculated by

k−1

rCO

CO2 , mol/L; kOH− is a reaction rate constant (1/s) for CO2 hydration could be calculated by (Pinsent et al., 1956)

−

= kOH− [OH ][CO2 ]

(9)

where rCO2 −OH− is reaction rate for CO2 hydration; [OH− ] is concentration of hydroxyl ion, mol/L; [CO2 ] is concentration of dissolved

2.5. Physicochemical properties Physiochemical properties of the aqueous NH3 solvent is essential to the calculation of the kinetics described above. Density, viscosity, diffusivity, Henry constant and chemical equilibrium constant were taken from literatures or Aspen Plus. The diffusion coefficient of CO2 in aqueous NH3 was calculated by the solution’s viscosities using a modified Stokes–Einstein equation (Versteeg and van Swaiij, 1988). NH3 −H2 O H2 O DCO = DCO ( 2

2

H2 O

H2 O DCO = 2.35 × 10−6 exp( 2

0.8

)

(15)

−2119 ) T

(16)

NH3 −H2 O

where H2 O and NH3 −H2 O are viscosities of water and ammonia solvent, respectively, Pa s; T, temperature, K. The viscosity data was predicted using Aspen Plus property database.

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G. Qi et al. / International Journal of Greenhouse Gas Control 41 (2015) 60–67

 Fig. 3. Effect of SO2 concentration on liquid film mass transfer coefficient (kG ), gas and liquid mass transfer resistance.

The concentration of hydroxyl ion [OH− ] was estimated by the chemical equilibrium of NH3 –H2 O system. The chemical equilibrium constant KNH3 was chosen from the Aspen Plus which is regressed from Pinsent et al. (1956).





Ln KNH3 = −1.257 −

3335.7 + 1.497Ln(T ) − 0.03706T T

(17)

The Henry constant of CO2 in aqueous ammonia and density of gas and solvent were obtained from the Aspen Plus, in which the NH3 –CO2 –H2 O system is well validated with the experimental results (Qi et al., 2013). 3. Results and discussion The CO2 mass transfer and kinetics data in the aqueous NH3 solvent with NH3 concentration varied from 0.579 mol/L to 4.012 mol/L at different SO2 concentrations from 0 ppm to 4000 ppm over experimental temperature range from 20 to 80 ◦ C are summarized in Table 1. These data are also presented in figures for detailed analysis.

mass transfer coefficient in 0.581 mol/L NH3 solvent at 40 ◦ C. The SO2 concentration is varied from 0 to 4000 ppm representative of the levels of SO2 in the flue gas of the coal-fired power station. As shown in Fig. 3, the mass transfer of CO2 is limited by liquid phase resistance which accounts for more than 90%. The gas phase mass transfer resistance contributes less than 10%. Particu ) decreases from larly, the liquid phase mass transfer coefficient (kG 2 0.44 to 0.22 mol/Pa/cm /s with SO2 concentration increasing from 0 to 4000 ppm. Accordingly, the liquid phase mass transfer resistance increases with SO2 concentration. This indicates that the SO2 absorption in aqueous NH3 competes with CO2 and inhibits the liquid phase CO2 mass transfer. Hence, the bi-liquid film assumption is necessary to properly understand the mass transfer process in the liquid film with both CO2 and SO2 absorption. Fig. 4 displays that the effect of SO2 concentration on liquid  ) and resistance at varied temphase mass transfer coefficient (kG  ) increases with temperature increasing from 20 peratures. The (kG to 80 ◦ C, but decreases with SO2 concentration at each temperature. On the contrary, the liquid film CO2 mass transfer resistance decreases with temperature, but increases with SO2 concentration. This is caused by the higher kinetics at higher temperature accelerating the rate-limited carbamate formation step (R1). Hence, the higher reaction temperature is more beneficial to CO2 absorption than SO2 absorption under the wetted wall conditions with solvent having no CO2 loading. Fig. 5 indicates the liquid phase CO2 mass transfer coefficient  ) and resistance, at different SO concentrations and aque(kG 2 ous NH3 concentrations. Similar to the impact of temperature,  ) increases with aqueous NH concentration from 0.581 to the (kG 3 4.012 mol/L, but reduces with SO2 concentration at same NH3 concentration. The liquid film CO2 mass transfer resistance decreases with aqueous NH3 concentration. Higher concentration of aqueous NH3 offers more opportunity of reaction between CO2 and NH3 and enhances the kinetics of carbamate formation as does the influence of temperature. In summary, higher SO2 concentration reduces the  ), higher temperature liquid phase CO2 mass transfer coefficient (kG and aqueous NH3 concentration benefit the kinetics of CO2 mass transfer. 3.2. CO2 pseudo-first-order reaction region in liquid film

3.1. Effect of SO2 absorption on CO2 mass transfer In this section, the effect of SO2 absorption on CO2 mass transfer at different SO2 concentrations, temperatures and NH3 concentrations is discussed. Fig. 3 shows the effect of SO2 concentration on gas and liquid phase mass transfer resistances, and liquid phase

The kinetics of CO2 absorption in pseudo-first-order reaction region in liquid film are described in this section. The apparent reaction rate constant (kapp ) at different temperatures and aqueous NH3 concentrations, and the second-order reaction rate constant at varied temperatures were calculated based on the reaction mechanism

 Fig. 4. Effect of temperature on (a) liquid phase mass transfer coefficient (kG ) and (b) liquid phase mass transfer resistance at varied SO2 concentrations.

G. Qi et al. / International Journal of Greenhouse Gas Control 41 (2015) 60–67

65

Table 1 Mass transfer and kinetics results summary of combined CO2 and SO2 absorption in wetted wall column. NH3 Conc. (mol/L)

Temp. (◦ C)

CSO2,in a (ppm)

KG,CO2 × 1010 (mol/Pa cm2 s)

kg × 1010 (mol/Pa cm2 s)

k G × 1010 (mol/Pa cm2 s)

k G,1 × 1010 (mol/Pa cm2 s)

k G,2 × 1010 (mol/Pa cm2 s)

kobs × 10−3 (1/s)

kapp × 10−3 (1/s)

k2 × 10−3 (m3 /kmol s)

0.585

20

0.445

1.216

1.136

1.954

0.608

3.504

3.221

5.597

0.579

80

0.735

6.663

5.833

10.267

1.731

40

0.938

5.338

5.201

2.997

2.875

40

1.351

10.942

10.769

3.741

4.012

40

\ 1.065 0.592 0.268 \ 1.238 0.688 0.431 \ 1.412 0.713 0.485 \ 4.255 1.659 \ 1.756 1.385 1.125 \ 3.055 2.365 1.808 \ 3.272 2.221 1.688

0.704

60

0.365 0.272 0.226 0.155 0.445 0.327 0.270 0.219 0.608 0.425 0.328 0.270 0.735 0.627 0.509 0.938 0.611 0.559 0.511 1.351 0.937 0.860 0.773 1.665 1.104 0.952 0.838

0.412

0.587

8.791 8.472 8.526 9.312 8.996 8.991 9.050 9.004 8.686 8.485 8.446 8.510 8.135 8.069 8.130 8.737 8.685 8.609 8.794 8.699 8.691 8.674 8.661 8.697 8.651 8.643 8.625

0.430

40

0.351 0.264 0.220 0.152 0.424 0.316 0.262 0.214 0.568 0.405 0.316 0.262 0.674 0.582 0.479 0.847 0.571 0.525 0.483 1.169 0.846 0.782 0.710 1.397 0.979 0.857 0.764

0.365

0.581

0 2000 3000 4000 0 2000 3000 4000 0 2000 3000 4000 0 2000 3000 0 2000 3000 4000 0 2000 3000 4000 0 2000 3000 4000

1.665

16.430

16.228

4.045

 Fig. 5. Effect of ammonia concentration on (a) liquid film mass transfer coefficient (kG ) and (b) liquid film resistance at varied SO2 concentrations.

Fig. 6. Effect of (a) temperature and (b) NH3 concentration on apparent reaction rate constant (kapp ) of CO2 .

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G. Qi et al. / International Journal of Greenhouse Gas Control 41 (2015) 60–67

Fig. 8. Extrapolated comparison of second-order reaction rate constant (k2 ) of CO2 as function of temperature. Fig. 7. Second-order reaction rate constant k2 of CO2 as function of temperature.

discussed above. Fig. 6(a) shows that the kapp increases from 412 to 5833 1/s with temperature varying from 20 to 80 ◦ C, and Fig. 6(b) shows that the kapp significantly increases from 412 to 16,228 1/s with aqueous NH3 concentration varying from 0.581 to 4.012 mol/L. As shown in Eq. (14), the apparent reaction rate constant (kapp ) of CO2 , which represents the carbamate formation rate is enhanced by higher temperature and aqueous NH3 concentration, which is the same as concluded in the previous section. It should be noted that the CO2 mass transfer process needs both kinetics and driving force to be optimized, since the equilibrium partial pressure of ∗ ) of CO loaded NH solvent, as shown in Eq. (2), reduces CO2 (PCO 2 3 2 at high temperatures, which will limits the CO2 solubility in NH3 solvent. Furthermore, the second-order reaction rate constant k2 of CO2 as a function of temperature was calculated based on the Eq. (14), as shown in Fig. 7. It indicates that the k2 increases with temperature significantly from 704 to 10,267 m3 /kmol s, when temperature increasing from 20 to 80 ◦ C. The Arrhenius expression of k2 was determined as: k2 = 3.121 exp

 −4720.8  T

(17)

Furthermore, the k2 in the present work was compared with several kinetic experimental results published in the literature, as shown in Fig. 8. Particularly, the k2 in this work (red, solid) was studied between 20 and 80 ◦ C, which was extrapolated to 5 ◦ C (red,

dash) in order to compare with the literature results. The k2 is most similar to the experimental data of Liu et al. (2012) (pink, dash) which were performed in the same wetted wall column. Hence, the consistency of the experimental results could be verified by the comparison. The k2 in the low temperature range is also very close to the results from Qin et al. (2010), Hsu (2003), Qin et al. (1983) and Pinsent et al. (1956), but slightly higher than their results at high temperature range, however, the k2 from Puxty et al. (2009) is much higher than the data in the present work, which is probably caused by the reactor and condition differences and the high volatility of NH3 at high temperature. The k2 comparison between MEA (black) and NH3 in this work (red) shows that the reaction rate of NH3 is 10-fold lower than MEA (Versteeg et al., 1996). 3.3. SO2 instantaneous reaction region in liquid film In order to further discuss the effect of SO2 absorption on the inhibition of CO2 mass transfer at different temperatures and aqueous NH3 concentrations, the liquid phase CO2 mass trans fer coefficient (kG,1 ) at SO2 instantaneous reaction region was  analyzed. As shown in Fig. 9, the (kG,1 ) decreases with SO2 concentration, but increases with temperature and aqueous NH3 concentration. The SO2 absorption in the SO2 instantaneous reaction region is considered to form a thin film with low active NH3 and high produced sulfite levels which reduce efficiency and kinetics for CO2 absorption. High temperature and NH3 concentration can enhance the diffusion and kinetics of CO2 mass transfer,

 Fig. 9. Effect of (a) temperature and (b) ammonia concentration on CO2 liquid film mass transfer coefficient (kG,1 ) in the SO2 instantaneous reaction region at varied SO2 concentrations.

G. Qi et al. / International Journal of Greenhouse Gas Control 41 (2015) 60–67

hence, choosing higher temperature and NH3 concentration could be considered in the combined CO2 and SO2 absorption. Comparing  Fig. 9(a) and (b), the effect of NH3 concentration on (kG,1 ) enhancement in the SO2 instantaneous reaction region is more obvious than temperature, because high temperature can improve both kinetics of SO2 and CO2 absorption, but high NH3 concentration ensures the CO2 absorption capacity of NH3 in the SO2 instantaneous reaction region after the NH3 consumed by SO2 absorption. 4. Conclusions In this work, the mass transfer and kinetics of combined CO2 and SO2 absorption in aqueous NH3 solvent were determined at different SO2 concentrations, temperatures and aqueous NH3 concentrations in a wetted wall column. A bi-liquid film assumption was proposed to analyze the mass transfer and kinetics by dividing the liquid film into a SO2 instantaneous reaction region and a CO2 pseudo-first-order reaction region. High SO2 concentration significantly reduces the overall liquid phase CO2 mass transfer coefficient (KG ), but high temperature and aqueous NH3 concentration benefit the CO2 mass transfer. In the CO2 pseudo-first-order reaction region, the apparent (kapp ) and second-order reaction rate constant (k2 ) of CO2 absorption were calculated and the second-order reaction rate constant can match very well with the previous work. In the SO2 instantaneous reaction region, the liquid phase mass trans fer coefficient (kG,1 ) decreases with SO2 concentration, caused by the lower diffusion efficiency and less available NH3 to absorb CO2 . Acknowledgment Financial support from Ministry of Science and Technology of China (Project no. 2013DFB60140) is greatly appreciated. References Astarita, G., Savage, D.W., Longo, J.M., 1981. Promotion of carbon dioxide mass transfer in carbonate solutions. Chem. Eng. Sci. 36, 581–588. Bishnoi, S., 2000. Carbon dioxide absorption and solution equilibrium in piperazine activated methyldiethanolamine. In: Ph.D. Thesis. The University of Texas at Austin, Austin, TX. Black, J., 2010. Cost and Performance Baseline for Fossil Energy Plants V1: Bituminous Coal and Natural Gas to Electricity. DOE/NRTL-2010/1397. Caplow, M., 1968. Kinetics of carbamate formation and breakdown. J. Am. Chem. Soc. 90 (24), 6795–6803. Carter, L.D., 2007. Retrofitting Carbon Capture Systems on Existing Coal-Fired Power Plants. American Public Power Association. Choi, W.J., Min, B.M., Shon, B.H., Seo, J.B., Oh, K.J., 2009. Characteristics of absorption/regeneration of CO2 and SO2 binary systems into aqueous AMP+ ammonia solutions. J. Ind. Eng. Chem. 15 (5), 635–640.

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