Catalytic Seawater Flue Gas Desulfurization Model - Environmental

Nov 18, 2009 - F. Vidal Barrero*, P. Ollero, A. L. Villanueva Perales and A. Gómez-Barea. Department of Chemical and Environmental Engineering, Unive...
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Environ. Sci. Technol. 2009 43, 9393–9399

Catalytic Seawater Flue Gas Desulfurization Model F. VIDAL BARRERO,* P. OLLERO, A. L. VILLANUEVA PERALES, AND ´ MEZ-BAREA A. GO Department of Chemical and Environmental Engineering, University of Seville, Camino de los Descubrimientos s/n, 41092, Sevilla, Spain

Received June 24, 2009. Revised manuscript received October 22, 2009. Accepted October 26, 2009.

A model of a seawater flue gas desulfurization process (SFGD) where oxidation of the absorbed SO2 is catalyzed by activated carbon is presented. The modeled SFGD process is comprised of two main units, an absorption packed scrubber, where SO2 absorption takes place, and an oxidation basin, where the absorbed SO2 is catalytically oxidized to sulfate, a natural component of seawater. The model takes into account the complex physical-chemical features of the process, combining mass-transfer, kinetics and equilibrium equations, and considering the electrolyte nature of the liquid phase. The model was validated with data from a SFGD pilot plant and a sensitivity analysis was performed, showing its predictive capability. The model is a useful tool for designing industrial desulfurization units with seawater.

Introduction Presently, energy conversion is based largely on combustion of fossil fuels, of which coal and oil contain a large amount of sulfur. During combustion, sulfur is converted into SO2, whose release to the atmosphere results in acid rain. Over the years, the SO2 emission limits have been lowered, thereby necessitating increased efficiency in desulfurization plants. In the United States, phase II of the Acid Rain SO2 Reduction Program, established under Title IV of the Clean Air Act Amendments of 1990, which came into force in 2000, imposes more stringent SO2 emission limits in order to accomplish a reduction of 3.5 million tons of SO2 emitted from power stations between 2001 and 2010. With respect to European Union regulations, the revised Large Combustion Plants Directive (LCPD, 2001/80/EC) imposes a 200 mg/Nm3, i.e., 70 ppmv (O2 content 6%, dry base) SO2 emission limit for large power stations (>500 MWth). This means a reduction of 50% with respect to the former limit. Nowadays, different technologies are available to control SO2 emission in fossil fuels power stations. They are classified as wet, semidry, and dry processes. The most widespread is wet limestone flue gas desulfurization (FGD), but for coastal power plants, FGD using seawater is a technologically simpler process. The two main advantages of seawater scrubbing are (i) an alkaline absorbent (Ca(OH)2, CaCO3) is not necessary since the seawater itself is alkaline (pH = 8.1-8.3) and readily available in coastal utility plants; and (ii) the effluent water treatment plant required is much simpler (an oxidation pool). However, this simple process has two main drawbacks: * Corresponding author phone: (+34)954487222; fax: (+34)954461775; e-mail: [email protected] 10.1021/es901863u CCC: $40.75

Published on Web 11/18/2009

 2009 American Chemical Society

(i) the amount of seawater required; and (ii) the large space requirements for the effluent treatment plant. Both factors are partially related to each other and to the oxidation rate from S(IV) to S(VI). Bromley (1), Tokerud (2), and Radojevic (3) have made important contributions to the development of the SFGD technology, ranging from introducing fundamental physicochemical principles, to addressing technical and financial issues. Tilly et al. (4) concluded that, with one possible area of concern, seawater technology was the best practicable environmental option (BPEO). Vidal et al. (5, 6), in a laboratory scale study, determined the oxidation kinetics of S(IV) in seawater. They demonstrated the significant catalytic effect of a commercially available activated carbon, which raised the oxidation rate at pH 4 to similar values at pH 6 for the noncatalytic oxidation. Then, it is not necessary to dilute the scrubber effluent with a large amount of fresh seawater, resulting in a smaller oxidation basin. In a later study, Vidal et al. (7) assessed, at pilot plant scale, an advanced SFGD system comprising an absorption tower filled with highefficiency structured packing, and a catalytic (activated carbon) oxidation tank. The existing bibliography regarding SFGD reveals a lack of specific model development. Abdulsattar et al. (8) modeled the solubility of sulfur dioxide in seawater based on thermodynamic principles. Later, Tokumura et al. (9) studied and modeled the neutralization of the seawater effluent due to the stripping of carbon dioxide dissolved in the effluent. However, this is not certainly the case with respect to studies about limestone wet FGD. Over the past 30 years several mathematical models have been proposed to describe the desulfurization process. These models usually present a detailed description of the chemistry involved, typically including SO2 absorption and oxidation, limestone dissolution and gypsum precipitation, and estimate the global desulfurization efficiency (10-19). The aim of this work was to develop and validate a model of a SFGD process where the oxidation of the absorbed SO2 is catalyzed by activated carbon. The paper is organized as follows. First, the absorber tower and the oxidation basin are modeled independently. Then, each submodel is validated using experimental data from a SFGD pilot plant. Finally, a sensitivity analysis is carried out using the validated model to analyze the effect of main design variables of the SFGD process. Model Description. In the SFGD process, SO2 contained in the flue gas is brought into contact with seawater in a scrubber where SO2 and CO2 absorption takes place (Figure 1).

FIGURE 1. Basic diagram of a SFGD plant. VOL. 43, NO. 24, 2009 / ENVIRONMENTAL SCIENCE & TECHNOLOGY



Both SO2 and CO2 are hydrolyzed in the aqueous phase, as shown in reactions I to IV, and oxidation of sulfite and bisulfite is also possible (reactions V and VI)

CC(IV) ) CCO2 + CHCO3- + CCO2-


CS(VI) ) CHSO4- + CSO2-




SO2 + H2O S HSO3 + H+


+ 2HSO3 S SO3 + H


+ CO2 + H2O S HCO3 + H


+ 2HCO3 S CO3 + H


HSO3 + 1/2O2 ⇒ HSO4


2SO23 + 1/2O2 ⇒ SO4


+ 2HSO4 S SO4 + H


H2O S H+ + OH-


The absorption tower’s effluent is first mixed with fresh seawater in order to raise the pH level. Then, the mixture is sent to an oxidation basin (Figure 1) where air is sparged to oxidize S(IV) to S(VI). In previous papers (6, 7) the authors have demonstrated that commercial activated carbon has a significant catalytic effect on the oxidation kinetics. So, it is possible the presence of this catalyst dispersed as solid particles in the oxidation tank. The simulation model comprises two submodels that correspond to the main units of the process: the Scrubber model for the absorption tower and the Oxidation model for the oxidation tank. Scrubber Model. The SFGD pilot plant scrubber is a counter-current packed tower. The following assumptions are considered for modeling purposes: (i) plug flow for gas (G) and liquid (L) phases; (ii) constant volumetric flow rates of gas and liquid throughout the tower; (iii) atmospheric pressure and isothermal conditions (298 K); and (iv) the gas is saturated with water and also is comprised of N2 (inert), SO2, O2 and CO2. According to these hypotheses, the model is made up of the following equations: Mass balance equations in the gas phase for SO2, O2 and CO2 G/S dPj ) -NjGa RT dh


Mass balance equations in the liquid phase (2)

L dCO2 L a - 1/2(-rS(IV))f ) NO 2 S dh


L dCC(IV) L a ) NCO 2 S dh


L dCS(VI) ) (-rS(IV))f S dh


where concentrations of S(IV), C(IV) and S(VI) are defined as 3



CH+ + Ceq ) CHSO3- + 2CSOO2- + CHSO4- + 3 2CSO2- + CHCO3- + 2CCO2- + COH4





where Ceq represents the net concentration of electric charge equivalent to that yielded by all the ions present in the seawater that do not form part of the reactions considered. The equilibrium constants (see SI) were taken from Abdulsattar et al. (8), and the activities in the equilibrium equations were related to the molar concentrations by means of the activity coefficients (aj ) γjCj), which were calculated using the Debye-Hu ¨ ckel extended equation. (See SI) The mass transfer coefficients in the gas and liquid phases L G ), and the diffusivities of the ionic and and kSO for SO2 (kSO 2 2 molecular species (see SI Table S1) were obtained from the literature. Based on them, we calculated the gas and liquid mass transfer coefficients for O2 and CO2 according to the following expression (Sada et al. (20)): k*j ) k*SO2(Dj /DSO2)2/3,* ) G or L

L dCS(IV) L a - (-rS(IV))f ) NS(IV) S dh

CS(IV) ) CSO2 + CHSO3- + CSO2-

To integrate this model it is necessary to calculate the mass transfer fluxes of SO2, O2 and CO2 between gas and liquid phases from known bulk gas composition (PSO2, PCO2, and PO2) and bulk liquid composition (CS(IV), CC(IV), CS(VI), and CO2) at each integration step. For that purpose, an interphase mass transfer model (see Supporting Information (SI)) based on the following hypothesis was used: (i) no chemical reaction takes place in the gas film; (ii) gas-liquid equilibrium is assumed for SO2, O2, and CO2 in the interphase; (iii) diffusion of molecular and ionic species occurs in the liquid film; (iv) the hydrolysis of CO2 and oxidation of S(IV) to S(VI) (Vidal and Ollero (5)) are very slow and they do not take place in the liquid film; (v) the other reactions are instantaneous and reversible and are assumed at equilibrium; (vi) the condition of zero electric current in the liquid film is forced; and (vii) constant diffusivities of ionic and molecular species. To calculate bulk liquid concentrations of SO2, HSO-3 , SO23 , HSO4-, SO42-, CO2, HCO3-, CO32-, H+ and OH- from known C(IV), S(IV), and S(VI) concentrations we have to solve the Equilibrium model which is composed by a set of 10 algebraic equations: Three equations, 6-8, which relate the concentrations of S(IV), C(IV), and S(VI) with the corresponding concentrations of molecular and ionic species. Six equations for the equilibrium of dissociation reactions I to IV, VII and VIII. And, finally, one equation imposing the electroneutrality of the liquid phase


The Henry constants (Hj) were obtained from the literature (Abdulsattar et al. (8); Chang and Rochelle (21); Degner and Hatzelmann (22)). Oxidation Model. In industrial plants, the oxidation basin is comprised of a set of parallel open-air channels. Air is injected throughout these channels so as to create the appropriate conditions for good transfer of oxygen to the seawater (see SI Figure S1). For modeling purposes, longitudinal piston flow for the liquid phase and transversal crossflow for the gas phase can be assumed in each channel. Because of the internal recirculation caused by injection of air, it is reasonable to assume good mixing for the gas in its transversal flow, ideally backmixing flow, i.e., the flow model for a CSTR. Thus, each channel could be roughly modeled as a set of CSTR reactors in series. The desulfurization pilot plant used in this study (Vidal et al. (7)) has a stirred oxidation

activated carbon) obtained by the authors in previous studies (Vidal et al. 5, 6) The interfacial area per unit of volume (a) is calculated using the results obtained by Bhavaraju et al. (23) and Tumeo and Stephens (24). G G ,NOG2,NCO it is necessary to consider the To calculate NS(IV) 2 diffusion-reaction phenomena taken place in the liquid film as well as the mass transfer exchange between the liquid and the bubbles. For that purpose, the interphase mass transfer model described in the scrubber section is used.

Results and Discussion

FIGURE 2. Comparison of the Equilibrium model with Bromley’s experimental data. tank seawater (see SI Figure S2) provided with an air sparger. For all practical purposes, this tank can be assimilated to one of the perfect mixing stages that comprise the model for an aeration channel. The mass transfer and chemical reaction phenomena that take place in the bubble stirred tank are similar to those that take place in the absorption tower. There will be transfer of O2 from the air bubbles toward the liquid bulk, but desorption of SO2 and CO2 toward the gas phase. Assuming perfect mixing in both phases, the mass balance equations for S(IV), O2, C(IV) and S(VI) are G in out out LCS(IV) ) LCS(IV) + NSO aVT + koxCS(IV) VT 2 in G out + NO aVT ) LCO + LCO 2 2 2

1 k CoutV 2 ox SIV T

(11) (12)

G in out ) LCS(IV) + NCO aVT LCC(IV) 2


in out out ) LCS(VI) - koxCS(VI) VT LCS(VI)


where kox is the first order kinetic constant (kS(IV) for noncatalytic oxidation or kAC for catalytic oxidation with

Equilibrium Model. As it was commented above, we use equilibrium equations for the liquid film and a complete Equilibrium model for the bulk liquid phase both in Scrubber and Oxidation models. To validate the Equilibrium model we calculated the pH of seawater as a function of the amount of S(IV) contained in it and compared the model results with the experimental data reported by Bromley (1) (Figure 2). Increasing the absorption of SO2 in seawater lowers the pH level, due to the higher release of protons during the SO2 hydrolysis reaction. As it can be observed, the comparison indicates that the model approximates the real system quite well. Scrubber Model. The most important variables in the Scrubber model are the partial pressure of SO2 in the gas (which indicates the desulfurization yield), the concentration of S(IV) and O2 in the seawater, and the pH of the acidic effluent from the tower. These parameters are of particular relevance for the oxidation of S(IV) in the seawater treatment plant. Model Validation. For the model to be a reliable tool for designing desulfurization units it should be validated with experimental data obtained on a sufficiently representative scale. In this section, we compare the results generated by the simulation model with the experimental results obtained in a pilot plant. (Vidal et al. (7)). Eleven tests were carried out (Table 1) combining two scrubber packing heights (1.1 and 2.2 m) with several L/G ratios (4-15 × 10-3 m3/Nm3) and inlet gas SO2 concentrations (3000-7500 mg/Nm3). The calculated (x-axis) and the measured (y-axis) desulfurization yields show a very good agreement (y ) 0.99x) as can be seen

TABLE 1. Experimental Results Used to Validate the Scrubber and Oxidation Models Desulfurization Experimental Results test

packing height, m

L/G, m /Nm

Cso2, inlet mg/Nm3

Cso2, outlet mg/Nm3

experimental desulfurization yield, %

D1 D2 D3 D4 D5 D6 D7 D8 D9 D10 D11


4.1 × 10-3 6.2 × 10-3 8.1 × 10-3 4.1 × 10-3 6.1 × 10-3 8.1 × 10-3 10.0 × 10-3 12.3 × 10-3 10.2 × 10-3 12.2 × 10-3 15.1 × 10-3

2963 2990 3027 6469 7244 8014 8290 8193 7143 7334 6086

614 105 12 2,620 1,601 737 166 25 1226 677 237

79.3 96.5 99.6 59.5 77.9 90.8 98.0 99.7 82.8 90.8 96.1





Oxidation Experimental Results test


T, K

O1 O2 O3 O4 O5 O6 O7 O8 O9 O10

2.8 3.7 4.0 4.1 6.1 3.0 4.0 2.8 2.5 4.0

316 309 312 308 307 311 309 309 306 307

WAC, kg/m3

7.1 6.9 7.0 14.1 13.9

DAC, m

experimental kOX, s-1

1.2 × 10-3 1.2 × 10-3 1.7 × 10-3 1.5 × 10-3 1.5 × 10-3

1.40 × 10-4 3.45 × 10-4 6.52 × 10-4 7.70 × 10-4 1.87 × 10-2 6.12 × 10-4 6.27 × 10-3 7.51 × 10-4 9.17 × 10-4 7.57 × 10-3




FIGURE 3. Adjustment between results obtained in the pilot plant and predicted by the Scrubber model.

FIGURE 4. Effect of L/G on the removal of SO2 and pH of the effluent from the tower.

FIGURE 5. Effect of interfacial area and packing height on the SO2 removal yield.

in Figure 3, and therefore, the Scrubber model is a useful tool for analyzing the desulfurization process with seawater in packing towers. Parametric Sensitivity Analysis. Once the Scrubber model was validated, we performed different analyses to study the effect of different variables related to the design and operation of SFGD plants. For example, below we present the effect of L/G ratio, interfacial area (associated with the type of packing) and packing height on the SO2 removal yield and the pH of the liquid effluent. The base case used to carry out the simulations in ) 7250 mg/Nm3; L/G ) 8 × 10-3 m3/Nm3; H ) 2.22 m; is CSO 2 a ) 500 m2/m3; T ) 308 °C. (a) Effect of L/G Ratio. Figure 4 shows the results obtained using the model for various L/G ratios. As L/G increases the desulfurization yield increases, resulting in a moderate decrease of pH of the tower’s effluent, due to higher SO2 absorption. When the desulfurization yield asymptotically approaches 100% (L/G ∼ 10 × 10-3 m3/Nm3), practically no further SO2 is absorbed (relative to the amount of seawater added), resulting in a significant increase in the pH level, due to dilution with fresh seawater. (b) Effect of Interfacial Area and Tower Height. The interfacial area per unit of volume is the packing’s parameter that most affect the absorption efficiency of a packing tower. For example, the commercial provider of the packing used in the pilot plant (Vidal et al. (7)), offers packings with 64 to 750 m2/m3 of interfacial area, although the typical range is between 250 and 500 m2/m3. Figure 5 shows the calculated desulfurization yield as a function of interfacial area and packed height. As expected,

the desulfurization yield increases asymptotically to 100%, as the interface area and/or packing height increases. (c) The Srubber Model As a Useful Tool for Optimal Design of SWFG Towers. The model developed is useful to calculate the desulfurization yield given the interfacial area, the contact height and the L/G ratio of an absorption tower, and can be used as tool to assist its design. The reason is that a given desulfurization yield could be attained with different combinations of interfacial area, contact height, and L/G ratio, but there is a different cost (€/kg of captured SO2) associated with each combination. The model could be used, for instance, to select the cheapest operational window ensuring the aimed desulfurization requirements. Oxidation Model. The most important variables related to the Oxidation model are (1) the oxidation pH, which depends on the fresh seawater added to the scrubber effluent, i.e., the dilution ratio; (2) the S(IV) conversion to S(VI); (3) the concentration of oxygen in the seawater, and (4) the chemical oxygen demand (COD) of the discharged seawater. Due to the low solubility of oxygen in seawater, this parameter is decisive to ensure that the discharge to the sea is a reliable environmental option. Model Validation. To validate the Oxidation model, we compared the results obtained from the model with the experimental one. The validation tests were done at different pH levels and temperatures, without catalyst and with different catalyst loads and sizes (Table 1). In Figure 6 we show the outlet S(IV) concentration obtained experimentally with those predicted by the model. As can be seen, the Oxidation model is capable to predict the seawater oxidation at the stirred tank under different operating conditions.




FIGURE 6. Adjustment between results obtained in the pilot plant and predicted by the Oxidation model.

FIGURE 7. Effect of dilution on the pH in the tank and the oxidation yield.

FIGURE 8. Effect of the concentration and diameter of the catalyst on oxidation rate. Parametric Sensitivity Analysis. By using the validated Oxidation model, we carried out a parametric analysis on dilution ratio, the catalyst load and particle size as important variables concerning the oxidation basin design. (a) Dilution Ratio. The oxidation pH is a very important operation variable when it comes to achieving a high oxidation ratesboth catalytic and noncatalyticsof S(IV) in seawater (Vidal et al. 5, 6). The desired pH is achieved by

diluting the acidic effluent from the absorption tower with fresh seawater. Figure 7 shows the positive effect over the oxidation yield (defined as (CS(IV)in - CS(IV)out)/CS(IV)in × 100) of using activated carbon in the oxidation unit and also the effect of dilution ratio. Without catalyst (dotted lines), it is necessary a high dilution ratio to raise the pH level, and thus to achieve an adequate oxidation yield. This entails larger oxidation basin VOL. 43, NO. 24, 2009 / ENVIRONMENTAL SCIENCE & TECHNOLOGY



and greater pumping power consumption. On the other hand, using the catalyst (solid line) it is possible to oxidize S(IV) at very low pH, thus requiring lower dilution ratios. (b) Effect of Catalyst Load and Size. Figure 8 shows the predictions of the Oxidation model at pH 4 and T ) 307 K when the concentration of activated carbon and the size of the catalyst particles are varied around a base case defined by DAC ) 1.2 mm and WAC ) 10 kg/m3. For the sake of comparison, this figure also contains the prediction of the model when the oxidation is carried out in the absence of a catalyst. It can be observed a large increase in the oxidation rate of S(IV) when activated carbon is used as a catalyst. These results agree with the data obtained by Vidal et al. (5-7) in their experimental works. A higher oxidation rate is obtained using smaller catalyst particles, meaning that the reaction is limited by internal diffusion. Also, the kinetic constant is not lineal with WAC, but proportional to WAC0.86, indicating that there is significant oxygen transfer limitation from the gas phase.

rj (- rS(IV)) S SFGD S(IV) S(VI) T VT WAC z Zj

generation rate of species “j” per unit of volume (kmol/m3 s) kS(IV)CS(IV); S(IV) oxidation rate (kmol/m3 s) tower section (m2) seawater flue gas desulphurization SO32-, HSO3- and H2SO3 species SO42-, HSO4- and H2SO4 species temperature (K) tank volume (m3) catalyst per unit volume (kg/m3) coordinate perpendicular to the film electric charge of ion “j”

Greek Letters δL thickness of liquid film (m) φ electric potential (V) activity coefficient of species “j” γj

Supporting Information Available Acknowledgments This study was carried out as part of a research project funded by the European Coal and Steel Community and by Endesa Ltd.

Appendix A

Interphase mass transfer model; diffusivities and equilibrium constants (Table S1); calculation of the activity coefficients (Table S2); Diagram of an industrial oxidation pool (Figure S1); and basic diagram of the oxidation tank used in the pilot plant tests (Figure S2). This material is available free of charge via the Internet at

Literature Cited Nomenclature a aj C(IV) Cj Cji Cjin Cjout DAC Dj f FGD G Hj h Jj KI to KIV, KVII and KVIII kjG kjL kS(IV) kAC L Nj Pj Pji R 9398


interfacial area per unit of volume (m2/ m3) activity of species “j” (kmol/m3) CO32-, HCO3- and H2CO3 species concentration of species “j” in the liquid phase (kmol/m3) concentration of species “j” in the interphase (kmol/m3) inlet concentration of species “j” in the oxidation tank (kmol/m3) outlet concentration of species “j” in the oxidation tank (kmol/m3) average size of catalyst particles (m) diffusivities of the ionic and molecular species (m2/s) volume of liquid phase per unit of volume of packing (m3/m3) flue gas desulfurization gas flow rate (air or flue gas, Nm3/s) Henry’s law constant for SO2, O2 and CO2 (atm m3/kmol) height of the scrubber (m) diffusive flux of ionic species “j” (kmol/s m2) dissociation constants gas phase mass transfer coefficient for SO2, O2 and CO2 (kmol/m2 s atm) liquid phase mass transfer coefficient for SO2, O2 and CO2 (m/s) noncatalytic kinetic constant (s-1) catalytic kinetic constant (with activated carbon) (s-1) seawater flow rate (tower or oxidation tank, m3/s) flux rate of SO2, O2 and CO2 (kmol/s m2) partial pressure of species “j” in the gas phase (atm) partial pressure of species “j” in the interphase (atm) universal gas constant (atm m3/K kmol)


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