Circulating Regeneration and Resource Recovery of Flue Gas

Sep 13, 2012 - Desulfurization residuals (using NaOH sorbent) were regenerated electrochemically, and at the same time sulfur in the flue gas was reco...
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Circulating Regeneration and Resource Recovery of Flue Gas Desulfurization Residuals using a Membrane Electroreactor: From Lab Concept to Commercial Scale Chenglei Yang, Ying Hu, Limei Cao, and Ji Yang* School of Resources and Environmental Engineering, State Environmental Protection Key Laboratory of Environmental Risk Assessment and Control on Chemical Process, East China University of Science and Technology, Shanghai 200237, P.R. China ABSTRACT: Desulfurization residuals (using NaOH sorbent) were regenerated electrochemically, and at the same time sulfur in the flue gas was recovered as H2SO4 and H2 was produced as a clean energy. Since industrialization should always be the final goal to pursue for lab technologies and the evolution of pilot- and full-scale commercial reactors has taken place relatively slowly, this paper is aimed to develop an electroreactor on a sufficiently large scale to evaluate the application potential of the proposed regeneration process. The following key design parameters are discussed: (1) voltage distributions over electrode, membrane, and electrolyte; and (2) scaling up correlation based on lab-scale reactor operation parameters. Thereafter, in the developed reactor, the desulfurization residuals using NaOH sorbent from a semidry flue gas desulfurization (FGD) facility of a power plant in Shandong Province were regenerated and it is significant to note that the electrochemical efficiency of the designed reactor is comparable to that of the chlor-alkali industry, showing that the technology is environmentally friendly and economically feasible. If this technology is to be employed for FGD, the facility could be a profitgenerating manufacturing part instead of a currently money-consuming burden for the plants.

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INTRODUCTION Sulfur dioxide (SO2), formed by the combustion of sulfurcontaining fuels such as coal and oil in most electric power generating units, is known to have detrimental effects on human health (e.g., cardiopulmonary diseases such as asthma and chronic obstructive pulmonary disease).1 Besides the health impacts, SO2 results in acid rain which leads to the acidification of lakes and streams, damage to agricultural crops, and erosion of buildings.2 Therefore, SO2 emission control is critical and strict regulations of SO2 emission standards have been promulgated worldwide. The common methods in practice to remove SO2 from flue gas include wet FGD, semidry FGD, and dry FGD.3 Conventionally, alkaline sorbents, such as limestone or lime, are employed to scrub flue gas due to their ability to form sulfur compounds.4,5 But these calcium based sorbents are usually nonregenerable, and the concentrated waste must be sent to disposal or alternatively processed to produce gypsum as an example of salable byproduct. Sodium alkaline is a regenerable scrubbing agent, by use of which the desulfurization efficiency of the semidry and dry FGD processes could match that of the wet FGD.6 Moreover, compared to wet FGD, semidry and dry FGD processes are more attractive alternatives especially for medium and small scale units because of their much smaller space requirement, no need for reheating energy, and the possibility of easily retrofitting onto existing facilities.7−9 However, these technologies have not yet been widely applied for two major reasons: (1) the high expenses for © 2012 American Chemical Society

sodium alkaline; and (2) compared to the once-through process,10 the physiochemical sorbent regeneration processes (such as the Wellman Lord’s process,11 sodium carbonate eutectic process,12 etc.) are cost-intensive due to the high usage of chemicals, power consumption for heat and pressurization, and the cost of waste treatment.3 Different from the traditional physiochemical processes, electrochemical technology is enjoying a growing acceptance by the industry13 since it can offer an elegant contribution toward sorbent regeneration. During the regeneration, only electron is added as a clean “reagent” and little or no secondary pollution is produced. More importantly, the electrochemical regeneration reactor could always use the surplus electricity of the power plant or the low priced valley electricity,14 which will decrease the operation cost significantly. Furthermore, the technology has lower temperature and capital requirements. It is in this context that current research was undertaken to propose this novel electrochemical process using membrane reactor to regenerate the desulfurization residuals as NaOH sorbent, at the same time to recover sulfur in the flue gas as sulfuric acid and to produce hydrogen as a clean energy as shown in Figure 1. Received: Revised: Accepted: Published: 11273

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analysis of the reactor at this scale shows that the technology proposed by the authors might have a bright future for application and the procedures described could be helpful for similar designs when considering electrochemical reactor scaling up.

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EXPERIMENTAL SECTION Materials and Reagents. The H2SO4, NaOH, and Na2SO4 were purchased from Shanghai Chemical Reagent Company of China and all the chemicals used were reagent grade or better. Deionized water was used throughout the experiments to prepare sample solutions. The desulfurization residuals were obtained from a semidry FGD of a Shandong Zaozhuang power plant using NaOH as the scrubbing agent. The detailed analysis of the residuals shows that it contained 82% Na2SO4 and 11% Na2SO3 by weight. The rest of the sample was composed by the oxides of Si, Fe, and Ca. The desulfurization sludge from the plant was dissolved in DI water to produce a saturated brine solution. The saturated brine solution was then introduced to the primary brine treatment system where most of the calcium and magnesium (dissolved impurities) were removed by chemical precipitation with sodium hydroxide. Then the solution was filtered with a cellulose filter before it was introduced into the electrochemical cell. Electrolytic Reactor. The lab-scale electrochemical reactor setup employed is shown in Figure 2. The reactor was divided into three compartments by the anode (26 cm ×13 cm) and cathode (26 cm ×13 cm). The electrodes were situated 1.5 cm apart from each other and 3.0 cm from the reactor wall. Ion exchange membranes (Wuxi Jianyi Co., China) to separate the anode and cathode in the lab scale reactor were JCM-II (a cation exchange membrane) and JAM-II (an anion exchange membrane), and attached to the electrodes as shown in Figure 2. For each run, the Na2SO4, H2SO4, and NaOH solutions with desired concentrations were introduced into the according compartments and continuously circled by three peristaltic pump (Baoding Longer Precision Pump Co., Ltd., China). A DC voltage stabilized power supply (Shanghai Liyou

Figure 1. Schematic diagram of the membrane electrolytic processes.

Although the increased availability of improved electrode materials and membranes together with off-the-shelf, modular reactors have helped to accelerate the development of industrial electrochemical technology significantly in the past few decades, the evolution of pilot- and full-scale commercial hardware has taken place relatively slowly. Consequently, the design and control of the major parameters that govern the electrochemical cell frequently raise problems in meeting the desired performance. Therefore, the aim of this paper is to develop a membrane electroreactor on a sufficiently large scale to show the potential for application. Several key design parameters for the electrochemical reactor are discussed: (1) voltage distributions over electrode, membrane, and electrolyte; and (2) scaling up correlation based on lab-scale reactor operation parameters. Using the developed reactor, the desulfurization residuals using NaOH sorbent from a semidry FGD facility of a power plant in Shandong Province were regenerated, and it is significant to note that the electrochemical efficiency of the designed reactor is comparable to that of the chlor-alkali industry (>80%).15 The economic

Figure 2. Schematic setup of the laboratory scale electrochemical regeneration process. 11274

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Anode: 2SO4 2 − + 2H 2O → 2H 2SO4 + O2 ↑ + 4e

Electrification Co., Ltd., China) was connected to both the anode and cathode and was used to control the potential. At appropriate time intervals, samples of 10 mL were taken from the reactor and analyzed during electrochemical regeneration. For the scaled up reactor, the DC power supply was HIFB2KA/15 V (Guangdong Liyuang Electrification Co., Ltd., China) Calculations and Analysis. The concentrations of sodium hydroxide and sulfuric acid were determined by acid−base neutralization titration using standard acid and base. The hydrogen generated was also collected and determined using a TYQRD-1102C hydrogen analyzer (GDHR Technology, Beijing, China). The current efficiency of sodium hydroxide production was calculated as follows:16

η=

(Ct − C0)VF It

The cathodic over potential of H2 evolution, which is a function of temperature and current density, could be estimated using the results presented by Rausch and Wendr,17 therefore Ti is employed as the cathode for hydrogen evolution in this work because of the low over potential and stability of the materials. Ti based dimensionally stable anodes (DSA), mostly oxide based anodes, have been widely employed for oxygen evolution in electrochemical industry also due to their low oxygen over potential and inert properties. The anodic over potential on a DSA anode is given by following expression:18

ηA = β log(ip/i0)

2(Ct − C0)VF It

(2)

The energy consumption E (kW h kg−1) was calculated as eq 3:16 E=

∫ CUIdt VM t

(3)

where Ct and C0 (mol/L) are the concentrations of NaOH or H2SO4 at time t and 0, respectively; V (L) is the circulated volume of solution, that is V = 0.70 L in the lab-scale reactor; F is the Faraday constant (96 485 C mol−1); I (A) is the current; U is the voltage drop across the membrane electroreactor, and M (g mol−1) is the molar mass of NaOH or H2SO4. All the experiments were conducted three times to check the reproducibility of results, and the match between successive experiments was within ±5%.

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Table 1. Properties of the Membranes Applied and the Calculated Voltage Drop over Membranes and Electrolyte

RESULTS AND DISCUSSION Lab-Scale Reactor Development. For the successful implementation of this novel concept for desulfurization residual regeneration, electrochemical reactor design is a critical stage. The cell potential and its distribution over the reactor are the most important ones to consider since the potential primarily governs the type of electrode reaction taking place and the current efficiency. The key components of the cell potential are defined in eq 4. ECell = (E eCell + ηA + |ηC |) + IRMembrane + IRSolution

membrane

average thickness (mm)

JCM-II JAM-II ip (A/m2) IR (V, CEM) IR (V, AEM) calculated ip (A/m2) IR (V)

(4)

where EeCell is the equilibrium value of the cell potential (V), η represents over potentials on the anode and cathode, IR is the voltage drop in the membrane and the electrolyte. It could be seen that for certain electrode reactions, the achievement of a low cell voltage requires the following strategies: (1) proper catalytic electrodes for the desired reaction to minimize overpotentials; (2) an appropriate membrane and electrolyte concentration to reduce the ohmic loss terms. Electrodes. Oxygen and hydrogen evolutions are the main reactions during electrochemical regeneration as shown below. Cathode: 2Na + + 2H 2O + 2e → 2NaOH + H 2↑

(7)

where β is a varying coefficient based on different electrode compositions, ip is the current density, and i0 is the exchange current density for oxygen evolution on the DSA. Typical DSAs with low β are RuO2, IrO2, SnO2, and PbO2 electrodes, among which a special interest in IrO2 based metallic oxide catalysts for oxygen evolution has attracted researchers due to the high corrosion-resistant properties, but only slight inferiority in electrocatalytic activities than RuO2. Consequently, Ti−IrO2 was chosen as the anode for the electrochemical reactor. The form of the electrodes is another important factor for reactor design. Despite the great progress in the electrode design and the emergence of a wide array of new electrodes that can efficiently handle diffusion-controlled reactions, the parallel plate electrode has become the most widely used. The success of the parallel-plate reactor is attributed to its merits including simplicity in both design and operation, and uniformity of potential and current distribution. In the following experiments, parallel plate anodes coated with IrO2 and Ti cathodes were employed throughout. Membrane and Electrolyte. Table 1 shows the basic characteristics of the membrane employed. The IR drop across

(1)

Similarly, the current efficiency of sulfuric acid production was calculated as eq 2: η=

(6)

conductivity (S·cm−1)

H2O content (%)

0.20 0.0034−0.0068 33−38 0.20 0.0020−0.0034 22−24 calculated voltage drop over membranes 500 1000 2000 0.15−0.30 0.30−0.59 0.60−1.18 0.30−0.50 0.59−1.00 1.18−2.00 voltage drop over electrolyte (25 °C, 10% Na2SO4) 500 1000 2000 0.10 0.19 0.38

the ion-exchange membranes at various current densities is given by eq 8 18 and also tabulated in Table 1: IR Membrane = [ipσ /K m]1000

(8)

where σ is the membrane thickness, Km is the membrane conductivity which depends on a number of factors including chemical structure of the membrane, ionic form, temperature, pH, water content, and electrolyte concentration in contact with membrane. It could be seen that the membranes selected for this research have good conductivity and low voltage drop,

(5) 11275

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Figure 3. Lab-scale electrochemical reactor performance under different conditions.

Lab-Scale Reactor Performance. Besides the electrode, membrane, and electrolyte, the design and optimization of an electrochemical process are based on a number of parameters such as feed flow rate, feed concentration, and current density, etc.20 The effect of feed flow rate on current efficiency of acid and base generation are shown in Figure 3a and Figure 3b, respectively. The results indicate that both current efficiencies of sulfuric acid and sodium hydroxide generation were higher at the feed flow rate of 12.12 L/h (retention time 105 s) and 18.48 L/h (160 s) than those at the feed flow rate of 5.70 L/h (78 s) and 24.84 L/h (341 s). The reasons of these phenomenon could be as follows: (a) The microbubbles formed by oxygen and hydrogen evolution on the electrodes resulted in the increase of electroreactor’s resistance and decrease of effective reaction area of electrodes when the feed flow rate was slow, while the microbubbles could be removed timely when the flow rate was high;19 (b) Generally, the high

which guarantees the high efficiency of the electrochemical reactor. The IR drop within electrolyte solution is separately obtained by the following expression:19 IR Solution = [ipl /K ]1000

(9)

where l is the distance between the electrodes, and K is the conductivity of the electrolyte. All the following experiments use this expression to minimize the voltage drop over the electrolyte, and the sample calculation is also listed in Table 1. It could be clearly seen that the membranes employed have reasonable IR drop at 500 A/m2, which is within the current range of chlor-alkali industry (500−800 A/m2). Also, it is evident that the increased current will lead to higher IR drop on both membranes and electrolyte, therefore lower energy efficiency. 11276

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Figure 4. Schematic diagram of the scaled-up reactor.

flow rate decreases the thickness of boundary layer, which is beneficial to the reaction. However, the overly high flow rate could make the flow turbulent which results in an apparent barrier of mass transfer at the interface of electrodes.21 Consequently, appropriate flow rate is critical for electrochemical reactor performance. Another important operating parameter for the electrochemical reactor is the feed concentration which determines the ionic strength and conductivity of the solution. The effect of initial acid and base concentrations on current efficiency is shown in Figure 3c and 3d, respectively. Although higher acid and base concentrations mean higher conductivity and lower electrolyte IR drop in the solution, the results indicate that the current efficiency of H2SO4 and NaOH generation decreased with the increase of initial concentrations. This could be explained that as the initial concentration of SO42− was 0.70 mol/L in the middle chamber, the osmotic pressure increased with the increase of the acid and base concentration in the corresponding chambers, which restricted the migration of ions through the membranes. Therefore, appropriate initial concentrations of acid and base were advantageous to ion migration. Current is the driven force of the electromigration of charged species in membrane electrolysis. The effect of current density on current efficiency of sulfuric acid and sodium hydroxide production is shown in Figure 3e and 3f, respectively. The results show that the current efficiency of acid and base

production both increased initially with the increased current density, then slightly decreased or reached a relatively stable value. What is meaningful is that under the optimal conditions, the current efficiency of H2SO4 in the lab scale reactor could reach 87%, and that of NaOH could be as high as 90%, which is higher than the average current efficiency of industrial chloralkali reactor (∼80%). This strongly manifests that the process proposed has the potential for scale-up. Scaling Up of the Electrochemical Reactor. It has been claimed that the reactor scale-up is still not an exact science, but is rather a mix of physics, mathematics, history, and common sense.22 In particular, the proper scale-up of the hydrodynamics and reactions inside the electrochemical reactors is subject to many uncertainties and pitfalls which may drastically deteriorate the reactor’s performance and economy. Therefore, it becomes obvious that the scale-up of a electrochemical reactor to a commercial size is a complex and troublesome endeavor. Current research and design work are based on a well-known fact, that the electrochemical behavior in the reactor is determined by the distribution of the electrochemical process rate along the lines of the electrical current, and is influenced by the hydrodynamic parameters of the system.23 Traditionally, scaling up of hydrodynamic systems is done with the aid of dimensionless parameters which must be kept equal at all scales in order to be hydro-dynamically similar. Dimensional analysis allows one to reduce the number of variables which have to be taken into account for mass transfer 11277

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Figure 5. Performance of the scaled-up reactor under different conditions.

determination. For mass transfer under forced convection, there are at least three dimensionless groups, the Sherwood number, Sh, which contains the mass transfer coefficient, the Reynolds number, Re, which contains the flow velocity and defines the flow condition (laminar/turbulent), and the Schmidt number, Sc, which characterizes the diffusive and viscous properties of the respective fluid and describes the relative extension of the fluid-dynamic and concentration boundary layer. Sh = 1.24Re

0.12

Sc

Table 2. Energy Consumption and Economic Assessment of Process operating parameters current density (A/m2) cell voltage (V) initial concentration of Na2SO4 (mol/L) initial concentration of H2SO4 (mol/L) initial concentration of NaOH (mol/L) retention time (s) energy consumption/T NaOH electric chargea (RMB/kWh) total cost for generation 1t NaOHb

−0.87 ⎛ DW /C ⎞−0.42 1/3⎛ DC / A ⎞

⎜ ⎟ ⎝ L ⎠

⎜ ⎝

L

⎟ ⎠

(10)

The dependence of Sh on Re, Sc, characteristic lengths, DC/L, and DW/L can be described in the form of the power series as shown in eq 10, in which DC/A is gap between cathode and anode, DW/C is gap between reactor wall and cathode, and L is the length of the electrode.24 The lab reactor was scaled up based on eq 10 to maintain the same hydrodynamics and the detailed reactor setup is shown in Figure 4. Actual FGD residuals from a power plant using NaOH sorbent in Shandong Province were regenerated in the scaledup reactor and the results are illustrated in Figure 5. Generally, the current efficiency of the industrialized chlor-alkali reactor is within the range of 80−85%. It can be seen in Figure 5 that under the optimal conditions determined using the lab-scale reactor, the performance of the scaled-up reactor is also satisfying and comparable to that of the chlor-alkali industry. Besides that, the following points should be noted: (1) The current efficiency of the reactor could reach 85%, 83%, and 87% for NaOH, H2SO4, and H2, respectively, while in the lab reactor the current efficiency was 87% for H2SO4 and 90% for NaOH under the optimal conditions. (2) The current efficiency curves showed the same trends in both lab- and pilot-scale reactors and reached the peak point at around 80 min, indicating the scaled-up reactor has similar hydrodynamic parameters as expected. (3) The data obtained from the pilot reactor are more stable than those from the lab-scale reactor because of the large volume of electrolyte involved, manifesting that the scaleup process is successful. Energy Consumption and Economic Assessment. The energy consumptions of the regeneration process using lab reactor and scaled-up reactor are estimated and tabulated in Table 2 under the optimal operating conditions. The result illustrates that the electricity requirement per ton NaOH is 2540 kWh for the lab reactor and 3060 kWh for the pilot one. The data are encouraging since the average electricity

lab-scale

pilot-scale

59.2 3.60 1.00

60.0 3.82 1.06

0.40

0.38

0.40

0.41

160 2540 kWh 0.48 1219.20 RMB ($193.7)

163 3060 kWh 0.48 1468.80 RMB ($233.4)

a

Source: The benchmark price of coal-fired power plants with FGD appliances from Shanghai Municipal Development & Reform Commission, People’s Republic of China, 2011. bThe central parity of exchange between RMB and USA dollars is 629.43 every hundred USA dollars from State Administration of Foreign Exchange, People’s Republic of China, March 31, 2012.

requirement for the chlor-alkali industry to generate one ton NaOH is around 2800−3200 kWh. And the cost of generating 1 Ton of NaOH by the membrane electrolysis process proposed in this study is estimated to be about 1219.20 RMB ($193.7) in the lab-scale reactor and 1468.80 RMB ($233.4) in the scaled-up reactor, which is significantly lower than the average market price (3190.00 RMB per ton of 99% NaOH) in East China in March 2012. However, other parameters should be considered for future industrialization: (1) The reactor capital cost. This reactor should be more expensive than the reactor used in chlor-alkali industry because three chambers instead of two chambers are incorporated. (2) This complexity definitely will lead to higher maintenance cost. (3) Electrode life span. This should be similar to the electrode life span in chlor-alkali industry since the same type of DSA such as IrO2/Ti or PbO2/Ti is employed. 4) Membrane expenses. The membrane expense should be higher in this reactor because both anion and cation membranes are used while only cation membrane is employed in chlor-alkali reactor. In addition, it has to be pointed out that in our reactor, H2 is generated as a clean energy besides NaOH and H2SO4 generated, which indicates that the technology proposed by the authors is environmentally friendly and economically feasible, not to mention that the cost could be significantly 11278

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(14) Molderink, A.; Bakker, V.; Bosman, M. G. C.; Hurink, J. L.; Smit, G. J. M. Domestic Energy Management Methodology for Optimizing Efficiency in Smart Grids; IEEE Bucharest Power Tech: Bucharest, Romania, 2009, (15) Kaveh, N.; Mohammadi, F.; Ashrafizadeh, S. N. Prediction of cell voltage and current efficiency in a lab scale chlor-alkali membrane cell based on support vector machines. Chem. Eng. J. 2009, 147 (2−3), 161−172. (16) Wei, Y. X.; Wang, Y. M.; Zhang, X.; Xu, T. W. Treatment of simulated brominated butyl rubber wastewater by bipolar membrane electrodialysis. Sep. Purif. Technol. 2011, 80 (2), 196−201. (17) Rausch, S.; Wendr, H. Morphology and utilization of smooth hydrogen evolving Raney nickel cathode coatings and porous sintered nickel cathodes. J. Electrochem. Soc. 1996, 143 (6), 304−311. (18) Chandrand, R. R.; Chin, D. T. Reactor analysis of chlor-alkali membrane cell. Electrochim. Acta 1986, 31 (1), 39−50. (19) Jalali, A. A.; Mohammadi, F.; Ashrafizadeh, S. N. Effects of process conditions on cell voltage, current efficiency and voltage balance of a chlor-alkali membrane cell. Desalination 2009, 237 (1−3), 126−139. (20) Lee, H. J.; Sarfert, F.; Strathmann, H.; Moon, S. H. Designing of an electrodialysis desalination plant. Desalination 2002, 142 (3), 267− 286. (21) Barra’gan, V. M.; Rui’z-Bauza’, C. Current-Voltage Curves for Ion-Exchange Membranes: A Method for Determining the Limiting Current Density. J. Colloid Interface Sci. 1998, 205 (2), 365−373. (22) Matsen, J. M. Scale-up of fluidized bed processes: Principle and practice. Powder Technol. 1996, 88 (3), 237−244. (23) Kazdobin, K.; Shvab, N.; Tsapakh, S. Scaling-up of fluidized-bed electrochemical reactors. Chem. Eng. J. 2000, 79 (3), 203−209. (24) Pak, D.; Chung, D.; Ju, J. B. Design parameters for an electrochemical cell with porous electrode to treat metal−ion solution. Water Res. 2001, 35 (1), 57−68.

reduced by using the readily available free surplus electricity of the power plant or low-priced valley electricity.14 If this technology is to be employed for flue gas desulfurization, the FGD facility could be a profit-generating manufacturing part instead of a currently money-consuming burden for the plants.

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AUTHOR INFORMATION

Corresponding Author

*Tel: +86-21-64251668; fax +86-21-64251668; e-mail: yangji@ ecust.edu.cn. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS This research is based upon work supported by the National Natural Science Foundation of China (Project 21177037), National 863 program (2009AA062603), Shanghai Leading Academic Discipline Project (Project B506), and “the Fundamental Research Funds for the Central Universities”. Any opinions, findings, conclusions, or recommendations expressed in this publication are those of the authors and do not necessarily reflect the view of the supporting organizations.

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