Analysis and Modeling of the Continuous Electro-oxidation Process for

Article pubs.acs.org/IECR

Analysis and Modeling of the Continuous Electro-oxidation Process for Organic Matter Removal in Urban Wastewater Treatment Antonio Dominguez-Ramos* and Angel Irabien Departamento de Ingeniería Química y Química Inorgánica, Universidad de Cantabria, Avenida de los Castros s/n, 39005, Santander, Cantabria, Spain ABSTRACT: Electro-oxidation (EO) is one of the most promising forefront technologies to remove pollutants from wastewater. Most of the available references have shown interesting results for biorefractory compounds, especially at laboratory scale. However, the continuous operation of an EO process has not been studied in detail, even less at pilot scale. In this work, the EO technology has been applied to treat an effluent from an Urban Waste Water Treatment plant in a pilot scale with an anodic area of 0.35 m2 (boron-doped diamond electrodes) taking into account the previous results obtained at lab scale (electrode area 7 × 10−3 m2). A model based on the mass balance, leads to the Damköhler number as the main parameter to describe the removal of organic matter expressed as chemical oxygen demand (COD) at lab and pilot scales. COD reduction is possible to manage by adjusting the flow rate to the anodic area ratio. The main difficulties for the process at pilot scale came from the formation of nonconductive films, which increases near two times the energy losses. aquaculture saline water,15 reverse osmosis concentrates,16,17 and other organic toxic compounds.18−23 Currently, a lot of interest is now focused in the decontamination of effluents polluted by hazardous substances,24−26 keeping in mind that the final toxicity of the effluent after the treatment should be taken into account in order to avoid the creation of additional environmental problems.27 However, few references are focused on urban wastewater treatment by electrochemical oxidation technology and even less are available for the continuous operation at pilot scale.9,28,29 Consequently, as the references are available about the removal of emerging pollutants, it is important to understand how an EO process can work under continuous operation such as in a UWWT plant, especially for low flow rate applications and water reclamation objectives, because EO processes allow a disinfection of the reclaimed water. The aim of this work is to perform an analysis and modeling of the continuous removal of organic matter taking glucose (as reference solution) and urban wastewater (as real wastewater) at pilot plant scale by electro-oxidation as first step for latter scale-up. A model based on the mass and energy balances has been applied to describe the experimental results by means of the Damköhler number, which was obtained from the fitting of the experimental results.

1. INTRODUCTION In the European Union the regulation of the Urban Waste Water Treatment (UWWT) by means of the Council Directive 91/271/EEC1 has led to secondary wastewater treatment covering 96% of the load in the EU-15 in 2007−2008.2 This is clearly an improvement over the general application status of the Directive at the beginning of 2003, as only 81% implementation of the cited directive was reported.3 In this way, the Council Directive 91/271/EEC aims to a treatment that allows the receiving waters (fresh-water and estuaries) to meet relevant quality objectives even for agglomerations of less than 2000 person-equivalent. When the technical and economical requirements are fulfilled, and especially for large flow rates to be treated, the common option for a secondary treatment is aerobic biological processes. These processes are usually the preferred option in UWWT plants because of their low energy consumption and proven performance.4 In spite of not being relevant energy consumers, UWWT plants are not free of resource consumption and environmental burdens. Their main cross-media effects are energy for aeration and the sludge generation: a proper management of this biowaste is therefore needed, which is a serious inconvenience. In fact, for the reporting period 2003−2006, almost 8.7 million tonnes of dry matter were generated in the EU-15 as a consequence of wastewater treatment. However, in agglomerations below 2000 population-equivalent, the installation of UWWT plants can be restricted by the economical feasibility of building such a small facility (low flow rates). Advanced oxidation processes such as electro-oxidation (EO) have been recently pointed out as alternatives within the urban water cycle5 with a high potential for effluent containing both biodegradable and biorefractory compounds. This technology does not generate sludge but its energy consumption is about 1 or 2 orders of magnitude higher than the activated sludge processes. Recent references show a wide range of potential applications: landfill leachate,6,7 dyes and textile industry effluents,8−12 high salinity industrial wastewater,13 herbicides,14 © XXXX American Chemical Society

2. MATERIALS AND METHODS A scheme of the electro-oxidation process and the main involved variables are shown in Figure 1, where Q is the continuous flow rate; C1 and C2 are the input and output COD; Aa is the anodic area; VT is the total liquid volume and U−i is Received: November 2, 2012 Revised: May 3, 2013 Accepted: May 9, 2013

A

dx.doi.org/10.1021/ie303021v | Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

Industrial & Engineering Chemistry Research

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was used. The applied current and the voltage were continuously recorded using a data acquisition card (National Instruments) connected to LabVIEW software. A HQ40d unit (Hach-Lange) was used to measure the conductivity and pH of the glucose solution in the storage tank, without any external agitation. 2.2. Pilot-Scale Experimental Setup. The experiments to study the electrochemical oxidation of the urban wastewater were performed in a pilot scale plant commissioned to Adamant Technologies. Once again following the previous flowsheet for the continuous EO process, the final set up of the EO process under pilot scale operation is shown in Figure 3. This pilot scale plant was installed in the facilities of a UWWTP. The plant shown in Figure 3 was equipped with a centrifugal pump, a rectifier, and a set of five serial sorted EO electrodes, Diacell type 108-01, with 10 undivided compartments. Each compartment consists of a BDD anode similar to that presented in the lab scale plant, but included a stainless steel cathode for reducing capital cost (also 70 cm2 and circular shape), keeping the 1 mm interelectrode gap. The ratio of the BDD anodic areas at pilot vs lab scale is 50 times. The centrifugal pumps were responsible to pump the effluent throughout the parallel electrochemical reactors to the reservoir tank (initially loaded with 500 L of process water from the facility whose conductivity was adjusted by adding Na2SO4 to the filtered effluent). At pilot scale, the flow rate in the EO reactors was kept constant at around 270 L·h−1, and the temperature was maintained almost constant (slight variations with outer temperature). A rectifier (Micronic Systems) was used to operate under galvanostatic conditions (up to 750 A at 16 V). Conductivity and pH was measured using a HQ40d unit (Hach-Lange) in the storage tank, ensuring at any time that both probes were completely submerged. A PLC was used to operate the pilot scale plant, which displays the applied current and voltage. 2.3. Experimental Conditions. At lab scale, as long as all the experiments are completed under the same current density i (60 mA·cm−2 to make possible the comparison against other previous references), the same concentration of supporting electrolyte (2500 mg Na2SO4·L−1), anodic area (70 cm2), and total volume (2 L) for a reference substance (glucose), the initial concentration C1 (L1−2) and the flow rate Q (L2−3) are subject to change, therefore their influence in the process can be analyzed. An inlet concentration around 500 mg O2·L−1 at lab scale is selected as a typical mean value for the inlet stream of a UWWT plant.30 A higher value around 1500 mg O2·L−1 was fed to study the influence of the inlet concentration. Sodium sulfate was added as supporting electrolyte of the synthetic solutions in order to increase the conductivity and facilitate the charge transfer in the solution (this salt has been commonly used in electro-oxidation runs at the previous level in similar tests). Glucose and sodium sulfate were provided by Panreac, whereas distillate water from Milli-Q was employed. Continuous flow rates (0.36−0.86 L·h−1) were fixed in order to get an expected conversion of around 50% and 70% in the continuous process. On the other hand, at pilot scale, the current density i was set up at 10 mA·cm−2 (galvanostatic mode) as it was suggested by previous experiments at lab scale. After an initial primary treatment for removing solids and large particles, wastewater from the UWWT plant was withdrawn from a chamber (sampling/withdrawn point was located immediately before a rotating biodisk system in the facility) by means of a submerged

Figure 1. Flowsheet of the continuous EO process.

the applied voltage-current density from the external power supply. The input stream (a glucose solution as reference at lab scale and filtered urban wastewater at pilot scale) characterized by an input concentration C1 and a flow rate Q is continuously pumped to the system from a storage tank. The output stream has an output concentration C2 and the same flow rate, under steady state conditions and the total liquid volume in the system VT remains constant. The stored liquid is recirculated throughout the electrochemical reactor in order to keep a high linear velocity over the electrode surface as needed by the electrochemical cell. The mean residence time τ in the whole system is determined by Q and VT thus τ = VTQ−1. An applied current density i is connected to the electrochemical cell and a total applied voltage U is measured between electrodes. Once the steady state is reached after an initial transient behavior, all the process variables are almost constant. 2.1. Lab Scale Experimental Setup. The experiments related to the electrochemical oxidation of the glucose solution were performed in a lab scale plant following the previous flowsheet of the continuous EO process. Figure 2 shows the general set up for this process under lab scale operation.

Figure 2. Set up of the continuous EO process al lab-scale: (1) electrochemical cell; (2) power supply; (3) process tank; (4, 5) pumps; (6) feed reservoir; (7) flow meter.

Glucose solutions passed through an electrochemical reactor that consists of a single compartment flow DiaCell Type 106: 02-06 (manufactured by Adamant Technologies). The electrochemical cell is equipped with two parallel thin-film borondoped diamond electrodes supported on silicon with an interelectrode gap of 1 mm: both anode and cathode are circular, accounting for a total area of 70 cm2 per electrode. The glucose solution is pumped continuously from the reservoir to the process tank and withdrawn from it as treated effluent by a peristaltic pump (Watson Marlow). A centrifugal pump is responsible for the circulation of the 2 L storage tank solution through the electrochemical reactor. In the experiments, the flow rate through the cell was kept constant at 300 L·h−1, and the temperature was maintained constantly around 23 °C. Under galvanostatic conditions, an Agilent power supply 6554A B



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Figure 3. Set up of the continuous EO process at pilot scale: (1) electrochemical cells; (2) power supply; (3) process tank; (4, 5, 8) pumps; (6) feed tank; (7) ring filter.

Table 1. Experimental Conditions for the Experiments Performed at Lab (L) and Pilot (P) Scale

EXP run

pollutant

scale of the experimental plant

L1 L2 L3 P1

glucose glucose glucose urban waste water

lab lab lab pilot

COD feedwater

flow rate

anodic area

flow rate to available anodic area

residence time

residence time

C1 (mg O2·L−1)

Q (L·h−1)

Aa (m2)

QA−1 a (L·h−1·m−2)

τ (h)

i (mA·cm−2)

0.86 0.86 0.36 10

0.007 0.007 0.007 0.35

123 123 51 29

2.33 2.33 5.56 50

60 60 60 10

1439 482 477 65

± ± ± ±

15 3 5 4

3. RESULTS AND DISCUSSION 3.1. Mass Balances of the Continuous EO Process. Previous references31,32 have shown that the main reaction involved in the electrochemical oxidation of organic matter at BDD anodes can be expressed as in eq 1:

pump. Particles were removed from the untreated effluent by a 5 μm ring filter previous to the feed tank. From the feed tank, the untreated effluent was pumped to the process tank and from that tank to the set of reactors as previously described. The effluent showed a mean value of 65 mgO2·L−1 (relatively low organic matter concentration) and a conductivity around 0.77 mS·cm−1. The process tank was filled with 500 L of process water from the facility, the conductivity being adjusted to the previous value with sodium sulfate in order to keep almost constant the conductivity from the beginning in the process. A continuous flow rate of 10 L·h−1 was chosen at the pilot scale to ensure a conversion greater than 50%. A summary of the experimental conditions of the experiments are presented in Table 1. All the experiments were performed under galvanostatic conditions rather than potentiostatic conditions in order to keep the ratio between the applied current and the available concentration of organic matter corrected by the ratio liquid volume to anodic area. The analytical measurement of the organic matter content in the solution was carried out to estimate the mineralization rate in the continuous EO process at laboratory and pilot scales. COD was obtained by means of a spectrophotometer DR 5000 UV−vis (Hach Lange) using corresponding COD kits (these kits are able to mask the presence of chloride ions). As previously cited, conductivity and pH were also registered continuously. These analytical measures match those completed in a regular basis in the UWWT plant except for biological oxygen demand and suspended solids (the former would not have been needed as solids are filtered before treatment then values would have been very low).

BDD[] + H 2O → BDD[OH·] + H+ + e−

(1)

To complete the mass balance in terms of COD, an equation for the reaction rate should be provided (−rCOD), whose extensive magnitude is the anodic area Aa, considering a first order reaction rate equation according to those mentioned in previous works32 whenever current levels of COD and current density are considered, leading to eq 2: ( −rCOD) = −

1 dm = kC2 A a dt

(2)

where m is the mass of COD, VT is the total volume of liquid, Aa is the anodic area and k is the observed kinetic constant, which has units of length/time (m·min−1). The mass balance assumes (i) the process will be operating under steady state (the accumulation term is removed) and (ii) there is wellmixing conditions for all the liquid processed in the total volume VT; thus the terms of input, output, and reaction of COD are considered: QC1 = QC2 + ( −rCOD)A a

(3)

Consequently, the amount of organic matter from the input stream (QC1) is removed from the system thanks to the reaction term ((−rCOD)Aa) and the output stream (QC2). The Damköhler number referring to the electrochemical oxidation in the electrode area is defined in eq 4: C


Industrial & Engineering Chemistry Research Da =

kC1A a = kA a Q−1 QC1

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(from 2.33 to 5.56 h). In this case it is observed that increasing τ drops the ratio QAa−1 thus leading to a higher conversion of COD in the steady state. In experiment L3, a mean conversion of 72% (C*2 = 0.28) in the steady state is obtained. Therefore, the new value of the flow rate Q modifies effectively the ratio QAa−1 and τ. According to the change in the Damköhler number the kinetic constant agrees well with the previous experiments as it is shown in Table 2. Results of the UWWT at pilot scale are shown in Figure 5, after a transient period t* ≈ 1 in P1, the concentration C*2 continues constant in the steady state. In this experiment P1, a mean conversion of 51% (C*2 = 0.49) in the steady state is obtained. In Table 2 the kinetic constant assuming first order electrochemical oxidation for the urban wastewater is shown. The comparison with other references using BDD anodes shows that UWW has a slower degradation rate than glucose and the previous reported compounds. The degradation of the wastewater from the UWWT plant is slower (k value for glucose is almost 4.4 times the value for the urban wastewater), even more than in the case of a biorefractory substance as lignosulphonate considering the values reported in Table 2 (whose values were around 1.36 × 10−3 m·min−1 and 1.51 × 10−3 m·min−1). On the other hand, the interpretation regarding the ratio QAa−1 is straightforward: the adjustment of the operating conditions (modification of the ratio QAa−1) makes it possible to reach the desired removal rates. If only one pass is completed,29 reducing Q could have a negative influence in the k value as the superficial velocity over the electrode surface will be simultaneously reduced. 3.2. Energy Balance for the Continuous EO Process. To identify the requested energy per unit of removed oxygen demand, the specific energy consumption per unit of COD, SECCOD is defined in eq 6:

(4)

Consequently the mass balance in eq 3 can be rewritten according to eq 5:

C2* =

1 1 + Da

(5)

where C*2 is the dimensionless output concentration which is defined as C2* = C2/ C1 and C1 is the mean concentration at the inlet stream. The outlet concentration C2* will be only determined by the parameter Da. As long as k and Aa remain unchanged, the outlet dimensionless concentration depends only on Da. The volume of liquid VT has no influence in the continuous operation because increasing VT would increase τ, but simultaneously it would increase the ratio VTAa−1 (already included in k); indeed, the extensive magnitude defined for the reaction is the anodic area Aa. Once the steady state of C2* is known from the experimental values, the obtained value of Da can be straightforwardly assessed. As the Da value becomes available (Q and Aa are known), it is possible to assess directly the experimental value of k for the set of the experiments. In this work, a dimensionless time t* is defined as t* = t/τ = t/ (VTQ−1) where t is the absolute treatment time. At lab scale, after a transient period t* ≈ 10 for L1−2 and of t* ≈ 5 for L3, the concentration C2* remains constant at the steady state as it is shown in Figure 4.

SECCOD = =

P UI Ui = = Q (C1 − C2) ( −rCOD)A a ( −rCOD)

Ui U OI ROI 2 = + kC2 ( −rCOD)A a ( −rCOD)A a

(6)

where UO is the minimum voltage requested for generating a faradaic current; RO is the ohmic resistance of the electrochemical system. The energy consumption per unit of volume treated is also used, thus the specific energy consumption per unit of volume SECV can be defined in eq 7: Figure 4. Dimensionless COD outlet concentration C*2 with time in the experiments L1−3: (×) inlet concentration for L1; (+) inlet concentration for L2; (∗) inlet concentration for L3; (▲) outlet concentration for L1; (Δ) outlet concentration for L2; (◊) outlet concentration for L3.

SECV =

UiA a P UI Ui U OI ROI 2 = = = = + Q Q Q Q Q QA a−1 (7)

Constant values are observed in parallel with the evolution of the conductivity (inset of Figure 6). For this experimental set up at lab scale, a minimum voltage UO was necessary at around 5 V32,33 (under L1 experimental conditions, a value of i = 11 mA·cm−2 is estimated for a conductivity λ = 3.5 mS·cm−1) experimentally in L1, SECCOD = 58 kWh·kg O2−1 At pilot scale, the initial applied voltage is much higher, UO = 8.9 V. The measured conductivity λ remains almost constant around a mean conductivity λ ≈ 0.77 mS·cm−1. As opposite to what is observed at lab scale, Figure 7 for P1 at pilot scale shows from t* ≈ 0.16 a linear increase in the total applied voltage about 49 mV·h−1, which represents an increase of 5.5 V

In the experiments L1 and L2 shown in Figure 4, there has not been found any influence of the inlet concentration to the process (≈1500 mg O2·L−1 and ≈500 mg O2·L−1, respectively), it means that a first order electro-oxidation referred to the COD can be assumed at the current density (60 mA·cm−2) and ratio QAa−1 (123 L·h−1·m−2). A mean conversion of 52% (C2* = 0.48) is reached at steady state allowing a direct identification of the Damköhler number and an estimation of the kinetic constants shown in Table 2. In experiment L3 the ratio QAa−1 is modified from 123 to 51 L·h−1·m−2 and the residence time τ is accordingly modified D



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Table 2. Results of Lab and Pilot Scale Experiments and Comparison with Other References

substrate glucose urban waste water formic acid (FA) glucose lignosulphonate urban waste water lignosulphonate

anodic area Aa (cm2)

mode

70 3500 2.7 3500 70 70 70 70

CR CR C BR BR BR BR BR

a

flow rate to anodic area QA−1 a (L·h−1·m−2) 51.4−123j 29 56

current density i (mA·cm−2)

COD kinetic constant k·103 (m·min−1)

60 10 16.7 30−60 10−60 15−60 15−30 30−60

2.19 0.50 2.36 (FA) 1.98b 1.36b 1.80b 1.51c

ref this work L1-3 this work P1 29 24

31

a

Mode: CR, continuous with recirculation (recirculation between the process tank and the electrochemical reactor); C, continuous without recirculation (one pass); BR, batch with recirculation (recirculation between the process tank and the electrochemical reactor). bThe original values of the kinetic constants considering the same experimental setup (operated in BR mode) are presented as TOC removal. As long the removal kinetic for TOC and COD fits a first order dependence versus concentration in the bulk of the liquid phase, it is possible to use these values to complete the comparison. cThe original value is also referred to TOC values. The value corresponds to those operation conditions in which the process is operating under mass transfer controlled conditions.

Figure 7. Dimensionless voltage U* for P1 (○). Inset: Evolution of conductivity λ (mS·cm−1) for P1 (○).

Figure 5. Dimensionless COD outlet concentration C*2 with time in the pilot plant experiment P1: (×) inlet concentration for P1; (▽) outlet concentration for P1.

after 112 h. This increase in the total applied voltage can be explained in terms of the formation of non conductive films over the electrode surface. The recommended procedure for electrode washing (alternating acid/alkaline washes) did recover the initial voltage. Consequently, in order to complete the scale up of the process is essential to know the value of the increase over time of the applied voltage, in order to assess the current energy consumption and determine the optimal point for voltage recovery. The presence of nonconductive films over the electrodes increases the SECV value from 27 to 46 kWh·m−3, which is almost 1.7 times the initial value (a mean value of 35 kWh·m−3 can be considered). In terms of SECCOD, a mean value SECCOD = 1066 kWh·kg O2−1 is obtained. This former value is clearly over the value presented for glucose at lab scale, which can be explained in terms of the increase of voltage with time due to nonconductive films, the presence of a 3.5−4.5 times lower conductivity (related to total applied voltage) and a lower concentration of organic matter in the inlet effluent. A summary of the main results corresponding to the energy balance is presented in Table 3 for lab and pilot scales.

Figure 6. Dimensionless voltage U*. Inset: Evolution of conductivity: (▽) L1; (Δ) L2; (◊) L3.

E



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Table 3. Results of the Energy Balance for Lab and Pilot Scale

pollutant

conductivity λ (mS·cm−1)

glucose glucose glucose urban wastewater

2.8 2.7 3.5 0.77

total applied voltage U (V)

specific energy consumption per unit of volume SECV (kWh·m−3)

specific energy consumption per unit of COD removed SECCOD (kWh·kg O2−1)

ref

8.9 8.7 10.1 10.2 (8.9−13.4)

43 42 118 35 (27−46)

58 171 344 1066 (829−1405)

L1 this work L2 this work L3 this work P1 this work

advances and trends for R&D. J. Chem. Technol. Biotechnol. 2008, 83, 769. (5) Anglada, Á .; Urtiaga, A.; Ortiz, I. Contributions of electrochemical oxidation to waste-water treatment: Fundamentals and review of applications. J. Chem. Technol. Biotechnol. 2009, 84, 1747. (6) Anglada, Á .; Urtiaga, A.; Ortiz, I. Laboratory and pilot plant scale study on the electrochemical oxidation of landfill leachate. J. Hazard. Mater. 2010, 181, 729. (7) Anglada, Á . ; Urtiaga, A.; Ortiz, I.; Mantzavinos, D.; Diamadopoulos, E. Boron-doped diamond anodic treatment of landfill leachate: Evaluation of operating variables and formation of oxidation by-products. Water Res. 2011, 45, 828. (8) Basha, C. A.; Sendhil, J.; Selvakumar, K. V.; Muniswaran, P. K. A.; Lee, C. W. Electrochemical degradation of textile dyeing industry effluent in batch and flow reactor systems. Desalination 2011, 285, 188. (9) Lui, Z.; Wang, F.; Li, Y.; Xu, T.; Zhu, S. Continuous electrochemical oxidation of methyl orange wastewater using a three-dimensional electrode reactor. J. Env. Sci. 2011, 23 (Suppl.), S70. (10) Migliorini, F. L.; Braga, N. A.; Alves, S. A.; Lanza, M. R. V.; Baldan, M. R.; Ferreira, N. G. Anodic oxidation of wastewater containing the Reactive Orange 16 Dye using heavily boron-doped diamond electrodes. J. Hazard. Mater. 2011, 192, 1683. (11) Petrucci, P.; Montanaro, D. Anodic oxidation of a simulated effluent containing Reactive Blue 19 on a boron-doped diamond electrode. Chem. Eng. J. 2011, 174, 612. (12) Jeong, J.; Lee, J. Electrochemical oxidation of industrial wastewater with the tube type electrolysis module system. Sep. Purif. Technol. 2012, 84, 35. (13) Huang, Y.-K.; Li, S.; Wang, C.; Min, Y. Simultaneous removal of COD and NH3-N in secondary effluent of high salinity industrial waste-water by electrochemical oxidation. J. Chem. Technol. Biotechnol. 2012, 87, 130. (14) Bringas, E.; Saiz, J.; Ortiz, I. Kinetics of ultrasound-enhanced electrochemical oxidation of diuron on boron-doped diamond electrodes. Chem. Eng. J. 2011, 172, 1016. (15) Díaz, V.; Ibáñez, R.; Gómez, P.; Urtiaga, A. M.; Ortiz, I. Kinetics of electro-oxidation of ammonia-N, nitrites and COD from a recirculating aquaculture saline water system using BDD anodes. Water Res. 2011, 45, 125. (16) Bagastyo, A. Y.; Radjenovic, J.; Mu, Y.; Rozendal, R. A.; Batstone, D. J.; Rabaey, K. Electrochemical oxidation of reverse osmosis concentrate on mixed metal oxide (MMO) titanium coated electrodes. Water Res. 2011, 45, 4951. (17) Zhou, M.; Liu, L.; Jiao, Y.; Wang, Q.; Tan, Q. Treatment of high-salinity reverse osmosis concentrate by electrochemical oxidation on BDD and DSA electrodes. Desalination 2011, 277, 201. (18) Chen, Y.; Shi, W.; Xue, H.; Han, W.; Sun, X.; Li, J.; Wang, L. Enhanced electrochemical degradation of dinitrotoluene wastewater by Sn−Sb−Ag-modified ceramic particles. Electrochim. Acta 2011, 58, 383. (19) Du, L.; Wu, J.; Hu, C. Electrochemical oxidation of Rhodamine B on RuO2-PdO-TiO2/Ti electrode. Electrochim. Acta 2012, 68, 69. (20) El-Ghrnymy, A.; Arias, C.; Cabot, P. L.; Centellas, F.; Garrido, J. A.; Rodríguez, R. M.; Brillas, E. Electrochemical incineration of sulfanilic acid at a boron-doped diamond anode. Chemosphere 2012, 87, 1126.

4. CONCLUSIONS The analysis and modeling of the continuous electro-oxidation process using boron-doped diamond electrodes for organic matter removal at lab and pilot scales has been completed. At lab scale (electrode area 7 × 10−3 m2), the COD showed that regarding the obtained conversion at steady state (i) there is no influence of the initial concentration (500−1500 mg O2·L−1) under the applied current density (60 mA·cm−2) and (ii) the ratio flow rate/anodic area (51−123 L·h−1·m−2) determines the desired conversion (72−52%). Therefore, a simplified model that relates conversion and the dimensionless Damköhler number allows the modeling and design of continuous processes. The kinetic constant obtained for glucose (2.19 × 103 m·min−1) agrees well with previous references. Consequently, the kinetic constants obtained under batch conditions can be used in the design of the continuous operation. At pilot scale (0.35 m2), the kinetic constant (from the model) for the wastewater had a lower value (0.50 × 103 m·min−1) when compared to the values of glucose at the lab scale, The energy balance revealed the formation of nonconductive films over the electrode surfaces treating UWW, so a continuously increasing applied voltage (49 mV·h−1) has been found, and therefore the specific energy consumption increased because of the ohmic resistance of the nonconductive layer.

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AUTHOR INFORMATION

Corresponding Author

*Tel.: +34 942 20 67 49. Fax: +34 942 20 15 91. E-mail: [email protected]. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS The authors gratefully acknowledge the financial support of the Ministry of Economy and Competitiveness of Spain through the project CTM2006-00317 and the project CENIT Sostaqua.

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REFERENCES

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