Modeling Solid-liquid Equilibrium of the NH3-CO2-SO2-K2SO4-H2O System and Its Application to Combined Capture of CO2 and SO2 Using Aqueous Ammonia

Available online at www.sciencedirect.com

ScienceDirect Energy Procedia 114 (2017) 834 – 839

13th International Conference on Greenhouse Gas Control Technologies, GHGT-13, 14-18 November 2016, Lausanne, Switzerland

Modeling solid-liquid equilibrium of the NH3-CO2-SO2-K2SO4-H2O system and its application to combined capture of CO2 and SO2 using aqueous ammonia Guojie Qi, Shujuan Wang* 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, Beijing 10084, China

Abstract It is of significance to intensively investigate the solid-liquid equilibrium of the NH3-CO2-SO2-K2SO4-H2O system to assess the novel aqueous NH3 based CO2 and SO2 combined capture process proposed in our previous work. In this work, the solubility of K2SO4 in aqueous NH3 at various temperatures, CO2 and SO2 loadings, and NH3 concentrations was predicted using Aspen Plus with a developed thermodynamic model package, which can accurately simulate the vapor-liquid equilibrium and the solid-liquid equilibrium of the combined CO2 and SO2 capture system. The results indicate that the solubility of K2SO4 increases with temperature and CO2 loading, but decreases with NH3 concentration. The precipitation starting point of K2SO4 shift to higher temperature with the increasing of SO2 loading and NH3 concentration. It is favorable to collect K2SO4 precipitates at lower CO2 loading and temperature, and higher SO2 loading and NH3 concentration. © by Elsevier Ltd. This is an open access article under the CC BY-NC-ND license © 2017 2017Published The Authors. Published by Elsevier Ltd. (http://creativecommons.org/licenses/by-nc-nd/4.0/). Peer-review under responsibility of the organizing committee of GHGT-13. Peer-review under responsibility of the organizing committee of GHGT-13. Keywords: Solid-liquid equilibrium; combined capture; CO2; SO2; K2SO4; Aspen Plus;

1. Introduction Carbon dioxide (CO2) and sulfur dioxide (SO2) are two of the main impurities from coal-fired power stations, which could be captured by flue-gas desulfurization (FGD) system and post-combustion CO2 capture (PCC) facility [1,2,3]. However, both FGD and PCC systems are very expensive, and cause significant reduction of the power

* A/Prof. Shujuan Wang. Tel: +86-10-62773154. E-mail : [email protected]

1876-6102 © 2017 Published by Elsevier Ltd. This is an open access article under the CC BY-NC-ND license (http://creativecommons.org/licenses/by-nc-nd/4.0/). Peer-review under responsibility of the organizing committee of GHGT-13. doi:10.1016/j.egypro.2017.03.1225

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station efficiency [4]. Aqueous ammonia (NH3), as an alternative chemical absorbent, has greatly potential to efficiently capture CO2 and SO2 together [5,6]. A novel post-combustion CO2 and SO2 combined capture process (without typical FGD system) using aqueous ammonia was proposed in our previous work, as shown in Figure 1, in that both CO2 and SO2 can be captured in one absorber, the CO2 related content in CO2 and SO2 enriched solvent can be thermally regenerated in a stripper, and the SO2 related content can be removed from the rich solvent by integrated sulfite forced oxidation and sulfate precipitation process [7].

Fig. 1. Proposed combined CO2 and SO2 capture process without FGD system.

In our previous work, the potassium additives (e.g. K2SO4) are proved and selected as the media for SO2 content removal in the combined capture process [5]. However, the detailed study on the solubility of potassium salts in the CO2 and SO2 loaded aqueous NH3 solvent is lacking in literature. Hence, it is of significance to intensively investigate the solid-liquid equilibrium of the NH3-CO2-SO2-K2SO4-H2O system and confirm the feasibility of its application in the combined capture of CO2 and SO2 using aqueous NH3. In this work, a solid-liquid equilibrium model was built using the Aspen Plus with our developed thermodynamic model package for combined CO2 and SO2 capture. The package can accurately calculate the solubility of K2SO4, K2SO3, K2CO3, KHCO3, (NH4)2SO4, (NH4)2SO3, and NH4HCO3 between 0 to 100 ºC, as shown in Figure 2.

Fig. 2. Solid-liquid equilibrium model validation.

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Afterwards, the impact of temperature (0-100 ºC), CO2 loading (0.1-0.3 C/N), SO2 loading (0.1-0.2 S/N), and NH3 concentration (2.5-7.5 wt%) on K2SO4 solubility was studied using the solid-liquid equilibrium model. The K2SO4 precipitation starting temperatures at the various conditions were also presented to identify the operation range of K2SO4 solid formation for SO2 content removal. 2. Solid-liquid equilibrium model In this solid-liquid equilibrium model, the electrolyte-NRTL model was adopted as the aqueous phase activity coefficient model to describe activity coefficients, enthalpies, and Gibbs energies for the liquid phase equilibrium behavior of the electrolyte NH3-CO2-SO2-K2SO4-H2O system. The Redlich-Kwong-Soave (RKS) equation was applied to calculate the vapor phase equilibrium fugacity coefficients. The electrolyte-NRTL model and the RKS equation were already described in detail elsewhere [8]. The reaction scheme of this NH3-CO2-SO2-K2SO4-H2O system, including gaseous species dissolution, liquid species dissociation and solid formation, could be presented as below [9]: eq,R1 2H2O m o H3O+ +OH 

(1)

eq,R2 NH3  H 2O m o NH 4 +OH 

(2)

eq,R3 CO2  2H2O m o H3O+ +HCO3

(3)

eq,R4 NH3  HCO3 m o H 2O+NH 2COO

K

(4)

eq,R5 HCO3  H2O m o H3O+ +CO32

(5)

eq,R6 SO2  2H 2O m o H3O+ +HSO3

(6)

eq,R7 HSO3  H2O m o H3O+ +SO32

(7)

H,R8 CO2 ( g ) m o CO2 (l )

K

(8)

H,R9 SO2 ( g ) m o SO2 (l )

(9)

K

K

K

K

K

K

K

H,R10 NH3 ( g ) m o NH3 (l )

(10)

s,R11 NH4 HCO3 (s) m o NH4  HCO3

(11)

s,R12 (NH4 )2SO3 ( s) m o 2NH 4  SO32

(12)

s,R13 (NH4 )2SO4 ( s) m o 2NH 4  SO42

(13)

s,R14 K 2SO4 (s) m o 2K   SO42

(14)

K

K

K

K

K

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Guojie Qi and Shujuan Wang / Energy Procedia 114 (2017) 834 – 839 s,R15 K 2SO3 ( s) m o 2K   SO32

(15)

s,R16 K 2CO3 ( s) m o 2K   CO32

(16)

s,R17 KHCO3 ( s) m o K   HCO3

(17)

K

K

K

As results of the aqueous phase reactions, there are four molecular species (CO2, SO2, NH3, and H2O), and nine ionic species, H3O+, OH-, NH4+, HCO3-, CO32-, NH2COO-, HSO3-, SO32-, SO42-. In addition, ionic species may form seven solid precipitates, such as NH4HCO3, (NH4)2SO3, (NH4)2SO4, K2SO4, K2SO3, K2CO3, and KHCO3. The chemical equilibrium constants and the Henry’s constant of the reactions above can be calculated with the following equation [10]:

ln K j

a  b / T  c ln(T )  d *T  e *(( P  Pref ) / Pref )

(18)

where Kj is the equilibrium constant of reaction j; T is the temperature in Kelvin; Pref is the reference state pressure. A heat exchanger module embedded in the Aspen Plus was adopted to model the K2SO4 solid formation at specific temperature in aqueous NH3 solvent with various CO2 and SO2 loadings. The heat exchanger temperature can be varied by specifying a sensitivity analysis module in Aspen Plus. The K2SO4 precipitation starting temperatures were determined as the intersection points of the K2SO4 solubility curves and the sulfite concentration curves (It assumes that all the sulfite was converted to sulfate by forced oxidation to simplify the calculation). 3. Results and discussion Figure 3 shows the K2SO4 solubility and the sulfite removal operation ranges in 2.5 wt% aqueous ammonia. The solubility of K2SO4 increases with temperature (0-100 ºC) and CO2 loading (0.1-0.3 C/N). As assumed, the sulfate concentration (mol/100g solvent) in aqueous NH3 can be represented by the sulfite concentration (mol/100g solvent). Hence, the SO2 loading (sulfite concentration) was varied from 0.1 to 0.2 S/N in aqueous NH3 with 0.2 C/N to compare with the K2SO4 solubility curve (blue-dashed line) and identify the K2SO4 precipitation ranges. The K2SO4 precipitation starting temperatures in 2.5 wt% aqueous NH3 with 0.2 C/N are 27 ºC, 12 ºC and slightly below zero when the SO2 loadings are 0.2, 0.15 and 0.1 S/N, respectively. The operable K2SO4 precipitation range (sulfite concentration > K2SO4 solubility) is larger and tends to higher temperature when the SO2 loading increases (highlighted filled areas in Figure 3).

Fig. 3. K2SO4 solubility and sulfite removal operation ranges in 2.5 wt% aqueous ammonia.

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The K2SO4 solubility and sulfite removal operation ranges in 5 wt% and 7.5 wt% aqueous ammonia were also calculated in this work, as shown in Figure 4 and Figure 5. Figure 4 indicates that the K2SO4 precipitation starting temperatures in 5 wt% aqueous NH3 with 0.2 C/N are 36 ºC, 69 ºC and slightly above 100 ºC when the SO2 loadings are 0.1, 0.15 and 0.2 S/N, respectively. Figure 5 presents that, in 7.5 wt% aqueous ammonia with 0.2 C/N, the K2SO4 precipitation starting temperature is 82 ºC when the SO2 loading is 0.1 S/N, and are higher than 100 ºC when the SO2 loadings are 0.15 and 0.2 S/N (out of the temperature range of this work). Similar to Figure 3, the operable K2SO4 precipitation ranges in Figure 4 and Figure 5 also increase with the increasing of SO2 loading.

Fig. 4. K2SO4 solubility and sulfite removal operation ranges in 5 wt% aqueous ammonia.

Fig. 5. K2SO4 solubility and sulfite removal operation ranges in 7.5 wt% aqueous ammonia.

In addition, the K2SO4 precipitation range increases when the NH3 concentration increases from 2.5wt% to 7.5wt%, as shown above, because the SO2 content increases in the solvent with higher NH3 concentration and same SO2 loading. And also, the solubility of K2SO4 significantly decreases when the aqueous NH3 concentration increases, as shown in Figure 6. Hence, the preferable condition for K2SO4 precipitation is at lower CO2 loading and temperature, and higher SO2 loading and NH3 concentration. The K2SO4 precipitation operation ranges presented in this work show that it is feasible to remove the SO2 content in aqueous NH3 by combining sulfite forced oxidation and K2SO4 precipitation (crystallization). For example, as shown in Figure 4, in the 5 wt% NH3 solvent with 0.2 C/N, it is possible to collect half of the sulfite content (from 0.2 to 0.1 S/N) if lower the solvent temperature from around 100 ºC to 36 ºC. It means that the CO2lean solvent from stripper could be partly introduced to the sulfite oxidator and the K2SO4 crystallizer (highlighted

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in Figure 1), when the SO2 loading in aqueous NH3 is highly accumulated to 0.2 S/N, to remove the sulfite content. These results are highly valuable to the design of the innovative combined capture process, especially the sulfite removal section, which will be presented in our future work.

Fig. 6. K2SO4 solubility in aqueous NH3 with different concentration.

4. Conclusions In this work, the solid-liquid equilibrium of the NH3-CO2-SO2-K2SO4-H2O system was studied using Aspen Plus with a developed accurate thermodynamic model package. The solubility of K2SO4 increases with temperature and CO2 loading, but decreases with NH3 concentration. The K2SO4 precipitation starting temperatures increase with NH3 concentration from 2.5 wt% to 7.5 wt%. Especially, in 5 wt% aqueous NH3 solvent with 0.2 C/N, which are 36 ºC, 69 ºC and slightly above 100 ºC when the SO2 loadings are 0.1, 0.15 and 0.2 S/N. Hence, It is preferable to precipitate K2SO4 at lower CO2 loading and temperature, and higher SO2 loading and NH3 concentration, and applicable to remove the sulfite content by splitting CO2-lean and SO2-rich solvent after the stripper to the sulfite oxidator and the K2SO4 crystallizer. The possible system blockage by K2SO4 solid and the compromise with solvent absorption performance will be also concerned in the design of this novel combined CO2 and SO2 capture process, which will be presented in our future work. Acknowledgements Financial supports from National Natural Science Foundation of China (Project No. 51576108) and Ministry of Science and Technology of China (Project No. 2013DFB60140) are greatly appreciated. References [1] IPCC. IPCC fourth assessment report. Geneva: IPCC, 2007. [2] Rochelle GT. Amine scrubbing for CO2 capture. Sci 2009; 325: 1652-1654. [3] Keeth RJ, Ireland PA, Radcliffe PT. Economic evaluation of FGD process. 1991 SO2 Control Symposium: Washington DC USA, 1991. [4] Black J. Cost and performance baseline for fossil energy plants volume 1: bituminous coal and natural gas to electricity. National Energy Technology Laboratory Final Report (revisions 2a): Washington, DC, USA, 2013. [5] Qi G, Wang S, Yu J, Zhao B, Chen C. Modeling and economic analysis on combined capture of CO2 and SO2 in flue gas using aqueous ammonia. Proceedings of the CSEE. 2013; 17: 4-11. [6] Ciferno JP, DiPietro P, Tarka T. An economic scoping study for CO2 capture using aqueous ammonia. National Energy Technology Laboratory Final Report: Washington, DC, USA, 2005. [7] Qi G. Study on combined capture of CO2 and SO2 in aqueous ammonia. Ph. D. thesis, Tsinghua University, China, 2013. [8] Aspen Technology. Aspen Physical Property System: Physical Property Methods. Cambridge MA, USA, 2010. [9] Qi G, Wang S, Lu W, Yu J, Chen C. Vapor-liquid equilibrium of CO2 in NH3-CO2-SO2 -H2O system. Fluid Phase Equilibr. 2015; 386: 47-55. [10] Aspen Technology. Rate-based model of the CO2 capture process by NH3 using Aspen Plus. Cambridge MA, USA, 2009.