4.3.1 Anodic oxidation as pre-treatment for a post-biodegradation process 4.3.2 Coupling between EF and bio-treatment 5. Conclusions and perspectives References CHAPTER 2. Electro-Fenton treatment of different pharmaceuticals under study: Kinetic study, energy consumption, mineralization pathway

ACKNOWLEDGMENTS

First, I would like to thank Prof. Stéphanie Rossano for having given me the opportunity to work in her laboratory the LGE. I also would like to acknowledge the director of this thesis Prof. Mehmet A. Oturan for his support during this project as well as for his scientific contributions and wise guidance during research works.

The contribution of Dr. Nihal Oturan during this thesis has also to be highlighted. I want to thank her for all the advice information and technical guidance which were always given as soon as possible.

A specific acknowledgment goes to Dr. Yoan Péchaud who spend a lot of time with me giving me very useful advice.

I wish to express my gratitude to the reviewers of this thesis, Prof. Manuel A. Rodrigo and Prof. Karine Groenen-Serrano for accepting to read and evaluate the present work as well as to the examiners, Prof. Christine Cacher-Vivier, Dr. Florence Fourcade, Dr. Emmanuel Mousset and again Dr. Yoan Péchaud.

I would like to thank the University Paris-Est and the doctoral school Sciences, Ingénierie, Environnement for funding conference fees and providing me financial support to perform this thesis work as well as ECOS-Sud who gave me the opportunity to go to Chili. I want also to thank all the people from the LGE lab and especially Valentina, Douglas, Jules, Théo, Julie, Clément, Anne who make me spent a very nice time in this lab.

Je voudrais remercier ma famille et mes amis pour leur soutien tout au long de la thèse.

Enfin je terminerai par te remercier mon Loulou, tu as toujours été à mes côtés pendant ces 3 années dans les bons moments comme dans les moments de doute. Grâce à ton appui sans faille j'ai su trouver la motivation pour arriver au bout. Encore merci je t'aime.

ABSTRACT

Electrochemical advanced oxidation processes (EAOPs) constitute an efficient technology to treat the pharmaceuticals as they allow the formation of strong oxidants such as hydroxyl radicals able to remove nearly any type of organic contaminants thanks to their very high oxidation power. Among them the electro-Fenton and anodic oxidation processes are environmentally friendly methods as they use no chemical reagent (anodic oxidation) or only oxygen of air and iron ions as a catalyst (electro-Fenton).

In this thesis, four pharmaceuticals from different families and structures were selected based on their toxicity and their occurrence in environmental waters and their removal from water was performed by EAOPs. The objectives of this work were to determine the best operating conditions at lab scale (current and catalyst concentration), investigate the kinetic of degradation and mineralization and finally propose a mineralization pathway based on aromatic intermediates, carboxylic acids and ions released to the solution. As these treatments were successfully applied, a lab scale pilot reactor composed alternately of BDD anodes and carbon felt cathodes with a bottom aeration system and working in the continuous mode was built to scale-up these processes in order to pre-industrialize them. Different configurations of electrodes were tested. The flow rate and the current were found to be more influent on the mineralization rate and on the energy consumption, respectively. To deeper understand the role of the flow rate and the configurations a hydrodynamic study was performed. The hydrodynamic results were gathered with a kinetic model for the mineralization to obtain a model predicting the percentage of mineralization at different position inside the reactor during the steady state. Thus, this model can help to optimize the operating conditions and to size future reactors depending on the mineralization objective of the treatment (high mineralization rate, combined treatment, high flow, …).

To reduce operating cost, the combination of an electrochemical process and a biological treatment was then investigated. In this frame, it was found that electrochemical treatment can (i) degrade the hydrochlorothiazide (ii) reduce significantly the concentration of its aromatic intermediates as they were shown to significantly inhibit the bacterial activity, (iii) promote the formation of biodegradable molecules such as carboxylic acids. The biodegradation of four carboxylic acids formed during the electro-Fenton treatment of the hydrochlorothiazide at lab scale was also studied. It was demonstrated that they were sequentially degraded with different lag phases and kinetics of degradation. Thus to mineralize them, a "plug flow" type reactor is recommended. The combination of treatment was then applied with an electrochemical treatment performed at low current with a BDD anode and a Platine anode. A mineralization degree of 38% and 50% were obtained by the biological treatment enabling to globally reach a mineralization rate of 66% and 85% with the BDD and the Platine anodes respectively. Thus this combined treatment was successful and open the way for the scale-up of these processes.

RESUME

Les procédés électrochimiques d'oxydation avancée constituent une technologie efficace pour traiter les produits pharmaceutiques car ils permettent la formation d'oxydants puissants tels que les radicaux hydroxyles capables d'éliminer presque tout type de contaminants organiques grâce à leur très haut pouvoir oxydant. Parmi eux, l'électro-Fenton et l'oxydation anodique sont des méthodes respectueuses de l'environnement car ils n'utilisent aucun réactif chimique (oxydation anodique) ou uniquement l'oxygène de l'air et des ions fer en tant que catalyseur (électro-Fenton).

Dans cette thèse, quatre produits pharmaceutiques appartenant à des familles différentes ont été sélectionnés en fonction de leur toxicité et de leur présence dans les eaux de l'environnement. Leur élimination de l'eau a été effectuée par électro-Fenton. Les objectifs de ce travail étaient de déterminer (i) les meilleures conditions opératoires à l'échelle du laboratoire (courant et concentration du catalyseur), (ii) la cinétique de dégradation et de minéralisation et enfin (iii) de proposer une voie de minéralisation basée sur des intermédiaires aromatiques, des acides carboxyliques et des ions libérés dans la solution.

Comme ces traitements ont été appliqués avec succès, un réacteur pilote composé alternativement d'anodes en BDD et de cathodes en feutre de carbone, doté d'un système d'aération et fonctionnant en mode continu a été construit pour évaluer la faisabilité d'un changement d'échelle et se diriger vers une pré-industrialisation du procédé. Différentes configurations d'électrodes ont été testées. Le débit et le courant se sont avérés être plus influents sur le taux de minéralisation et sur la consommation d'énergie, respectivement. Pour mieux comprendre le rôle du débit et des configurations, une étude hydrodynamique a été réalisée. Le modèle hydrodynamique a été associé à un modèle cinétique de minéralisation afin d'obtenir un modèle permettant de prédire le pourcentage de minéralisation à différentes positions à l'intérieur du réacteur en régime permanent. Ainsi, ce modèle peut aider à optimiser les conditions opératoires et à dimensionner les futurs réacteurs en fonction de l'objectif de minéralisation du traitement (taux de minéralisation élevé, traitement combiné, flux élevé,…). Afin de réduire les coûts opératoires, la combinaison d'un procédé électrochimique et d'un traitement biologique a ensuite été étudiée. Afin d'obtenir un traitement combiné efficace, il a été constaté que le traitement électrochimique devait (i) dégrader l'hydrochlorothiazide (ii) réduire de manière significative la concentration de ses intermédiaires aromatiques car ils inhibent de manière significative l'activité bactérienne, (iii) favoriser la formation de molécules biodégradables telles que les acides carboxyliques. La biodégradation de quatre acides carboxyliques formés lors du traitement par électro-Fenton de l'hydrochlorothiazide a également été étudiée. Il a été démontré qu'ils étaient dégradés de manière séquentielle avec différentes phases de latence et cinétiques de dégradation. Ainsi, pour les minéraliser, un réacteur de type «piston» est recommandé. La combinaison de traitement a ensuite été appliquée à un traitement électrochimique effectué à faible courant avec une anode en BDD et une anode en Platine. Un degré de minéralisation de 38 et 50% a été obtenu par le traitement biologique permettant d'atteindre un taux de minéralisation global de 66 et 85% avec les anodes en BDD et Platine respectivement. Ainsi, cette combinaison de traitement a été un succès, un changement d'échelle du procédé peut alors être envisagé. Fig. 20: Kinetic of degradation during the elimination of 0.1 mM HCT solution under 500 mA, 50 mM of Na2SO4 and 0.1 mM Fe 2+ with and without additional inorganic ions. The additional ions were 244 mg L - 1 NaCO3, 48 mg L -1 of Cl -, 4 mg L -1 of PO4 3-, 0.9 mg L -1 NO2 -and 1.5 mg L -1 of NO3 - 138 Fig. 27: TOC evolution during the mineralization of 0.1 mM HCT solution at 500 mA in 50 mM of Na2SO4 with and without additional ions using 0.1 mM Fe 2+ as catalyst. The additional compounds were 244 mg L -1 Na2CO3, 48 mg L -1 of Cl -, 4 mg L -1 of PO4 3-, 0.9 mg L -1 of NO2 -and 1.5 mg L -1 of NO3 -150 Fig. 28: Time-course of carboxylic acids during the electrolysis of 0.1 mM HCT at 500 mA constant current. [Fe 2+ ] = 0.1 mM, [Na2SO4]0 = 0.05 M, room temperature Circles correspond to the 14-electrode and squares to the 28-electrode configurations, full points to a flow rate of 20.4 mL min -1 and stripe dots to a flow rate of 42.5 mL min -1 Fig. 44: RTD functions plotted against the time for different flow rates under 0.2 mL min -1 of air with 28electrode configuration

Introduction

Nowadays, the treatment of pharmaceuticals has become a hot topic as they are widely consumed all over the world. A strong increase in the consumption of pharmaceuticals has been observed during the last decades. In the US, IMS Health mentioned an increase of the sales from $152,8 to $280 billion which represents a 83% rise between 2000 and 2006 (Khetan and Collins 2007;Bouissou-Schurtz et al. 2014;Vergili et al. 2019). This increase can be explained by several factors such as: (i) population growth, (ii) population ageing, (iii) discovery of new properties of already existing drugs and (iv) decreasing price of drugs as patents expire [START_REF] Daughton | Cradle-to-cradle stewardship of drugs for minimizing their environmental disposition while promoting human health[END_REF].

As analytical chemistry techniques were improved during last years, especially in the limit of quantification, the presence of pharmaceutics was detected in natural water thus proving their non-effective elimination by wastewater treatment plants (WWTP), since concentrations ranging from ng L -1 to µg L -1 were usually found in natural water bodies (Lindsey et al. 2001;Bouissou-Schurtz et al. 2014). Indeed, in addition to the industrial and hospital effluents, these products are also released to urban wastewater whole or partly degraded after human and veterinary uses (Lindsey et al. 2001;Verlicchi et al. 2010). Pharmaceutical compounds are not efficiently eliminated by WWTP or by natural elimination processes in the environment [START_REF] Laurencé | Anticipating the fate and impact of organic environmental contaminants: A new approach applied to the pharmaceutical furosemide[END_REF]. Several review papers provide a snapshot of the removal efficiency of some selected pharmaceutical molecules achieved in WWTPs (Miège et al., 2009;Gabet-Giraud et al., 2010;Luo et al., 2014).

Pharmaceuticals include lots of different compounds; they can be classified into groups aiming at doing the same action. The most important ones are the nonsteroidal anti-inflammatory drugs and analgesics, the antibiotics, the β-blockers and diuretics, the neuroactive compounds, the steroid hormones and the lipid regulators…

• Nonsteroidal anti-inflammatory drugs and analgesics: this group includes more than one hundred compounds [START_REF] Feng | Removal of residual antiinflammatory and analgesic pharmaceuticals from aqueous systems by electrochemical advanced oxidation processes. A review[END_REF]. They are used as inflammatory reducer and analgesic; their high concentration in natural water can be explained by the fact that they do not need a prescription, allowing very great use.

• Antibiotics: their role is to kill bacteria or/and inhibit their multiplications. This group is composed of a huge number of molecules.

• Beta blockers: they are part of the most important prescription drugs all over the world and used in cardiovascular disease, which is a more and more common problem, and in relieving chest pain (Khetan and Collins, 2007) .

• Diuretic: they are used for anti-hypertensive medications.

Pharmaceuticals are ubiquitous in the environment

As pharmaceuticals have been detected in the environment, it is important to quantify their extent in environmental water. To estimate the concentration of pharmaceuticals in an effluent two main techniques are used: i) direct measure of the pollutant in the effluent (Table 2) and ii) use of PEC (Predicted Environment Concentration), a calculation to predict the concentration of a pollutant in an effluent based on approximations. The PEC is usually based on the fraction of active ingredient excreted in urine and feces but all the formulas are different and linked to their specific situations. For the different types of pharmaceuticals, it appears that for most of compounds, the PEC is higher than the concentration measured in the same condition, indicating that the PEC values are overestimated [START_REF] Besse | Exposure assessment of pharmaceuticals and their metabolites in the aquatic environment: application to the French situation and preliminary prioritization[END_REF]Kümmerer 2009a;[START_REF] Escher | Environmental toxicology and risk assessment of pharmaceuticals from hospital wastewater[END_REF][START_REF] Sim | Occurrence and distribution of pharmaceuticals in wastewater from households, livestock farms, hospitals and pharmaceutical manufactures[END_REF]Bouissou-Schurtz et al. 2014;[START_REF] Gabet-Giraud | Impact of wastewater treatment plants on receiving surface waters and a tentative risk evaluation: the case of estrogens and beta blockers[END_REF]Mendoza et al. 2015;[START_REF] Qarni | Investigating the removal of some pharmaceutical compounds in hospital wastewater treatment plants operating in Saudi Arabia[END_REF].

As can be seen in Table 2, b-blocker, NSAIDs and analgesics and antibiotics are present in higher concentrations than diuretics both in surface water and in hospital wastewater. Among them, Atenolol and Metoprolol, Ibuprofen, Ciprofloxacin and Erythromycin have the highest concentrations. However, these results are very specific and cannot be expanded to other countries or even other regions.

Pharmaceuticals have harmful effect to living species

Besides measuring or predicting the concentration of pharmaceuticals, their toxicity is also an important characteristic. Two types of toxicity have to be distinguished: i) short term effect and/or exposure which is called acute toxicity and ii) long time effect and/or exposure, named as chronic toxicity. For persistent organic compounds (including the pharmaceutical), the chronic toxicity is more suitable, that is why European guidelines (EMEA 2006) recommend to use this term for the pharmaceuticals. However as the chronic toxicity data are often not available for pharmaceutical compounds [START_REF] Crane | Chronic aquatic environmental risks from exposure to human pharmaceuticals[END_REF] , acute toxicity data are in general used instead.

To measure those different toxicities several parameters are often used; the most important ones are LC50, EC50, NOEC and LOEC, defined in Table 1. EMEA 2006). This value can be calculated based on the previous parameters described above (e.g. EC50. NOEC, ...). PNEC is obtained by dividing the given parameter by a safety factor (AF).

Among different class of pharmaceuticals defined above, the antibiotics seem to be the more toxic with PNEC < 1 µg L -1 , the more toxic being ciprofloxacin with a PNEC equal to 0.02 µg L -1 followed by ofloxacin, clarithromycin and erythromycin with 0.04 µg L -1 [START_REF] Kümmerer | Promoting resistance by the emission of antibiotics from hospitals and households into effluent[END_REF].

The second more dangerous category of pharmaceuticals is the anti-inflammatories and analgesics with especially two compounds having a PNEC below 1 µg L -1 : ibuprofen (0.01 µg L -1 ) and diclofenac (0.05 µg L -1 ) (Mendoza et al. 2015). Among the β-blockers and the diuretics only one compound has been found as really hazardous: propranolol with a PNEC of 0.05 µg L -1 [START_REF] Huggett | Toxicity of select beta adrenergic receptorblocking pharmaceuticals (B-blockers) on aquatic organisms[END_REF].

PNEC is a good way to approximate the toxicity but it would be necessary to gather all information approved and to perform some others to have a unique PNEC based on chronic toxicity for a larger number of species. Another point of limitation of this method is that it does not consider the potential effect of the presence of a cocktail of pharmaceuticals which was shown to have a different toxicity than that predicted by the PNEC values of the different pharmaceuticals taken separately. The synergetic effect can be whether antagonist or additive [START_REF] Frédéric | Pharmaceuticals in hospital wastewater: their ecotoxicity and contribution to the environmental hazard of the effluent[END_REF]. Tough a few studies were done on the additive effect, this phenomenon has been illustrated by [START_REF] Cleuvers | Initial risk assessment for three b-blockers found in the aquatic environment[END_REF] and [START_REF] Brian | Accurate prediction of the response of freshwater fish to a mixture of estrogenic chemicals[END_REF]. It is assumed that when different pharmaceuticals have the same action mode, their concentration can be considered as added in evaluating the toxicity of the solution [START_REF] Escher | In vitro assessment of modes of toxic action of pharmaceuticals in aquatic life[END_REF][START_REF] Fraysse | Prediction and experimental validation of acute toxicity of betablockers in Ceriodaphnia dubia[END_REF]. In contrast, a study reported by Kümmerer (2009a) revealed that in presence of 5-fluorouracil, the thresholds effect of cephalosporine could be decreased significantly. As a consequence, to evaluate the real toxicity of effluents, the mixture toxicity must be used.

Even if the methods used for the evaluation of the toxicity of pharmaceuticals are still not very accurate, it clearly appears that these compounds represent a threat to the environment as most of them can be considered as persistent hazardous organic compounds.

Pharmaceuticals are among persistent organic pollutants:

The persistence of a pharmaceutical is its capacity to resist to physical, chemical and biological degradation processes and therefore to maintain a detectable concentration in water. As a consequence, persistent pollutants are not eliminated by a conventional WWTP or by natural elimination in the environment. The elimination processes occurring in a conventional WWTP are well-described by Margot et al. (2015). Pharmaceuticals can be eliminated from wastewater during the sewage treatment whether by being transferred to another phase or by being degraded. In the first case pharmaceuticals are sorbed onto sludge or onto particular compounds thanks to hydrophobic or electrostatic interactions or evaporated. Kümmerer underlines the fact that very little has been done about the sorption of antibiotics and this needs to be studied as it is an important parameter (Kümmerer 2009b). For example, ciprofloxacin was found to sorb well onto activated sludge or sediments [START_REF] Golet | Environmental exposure and risk assessment of fluoroquinolone antibacterial agents in wastewater and river water of the Glatt Valley watershed , Switzerland[END_REF]. Some pharmaceuticals can also be degraded, in the better case, reaching total mineralization of the mother molecule. This process can occur by biological transformation or by abiotic degradation which involves reactions such as photolysis or hydrolysis. Elimination can also happen during wastewater transportation, in this situation the process more likely to occur is photodegradation. This is well-explained in a relevant review article written by Khetan and Collins (2007). Thus, the percentage of removal is hardly ever equal to the degradation rate as some of the pollutants can adsorb themselves on the bacteria. This parameter is indirectly often evaluated thanks to the log(Kow), the octanolwater partition coefficient (Table 2).

Some pharmaceuticals, such as metoprolol, diclofenac and tramadol are hardly removed by sewage treatment plants with removal efficiency less than 40%. Other molecules such as atenolol, ketoprofen and trimethoprim are moderately removed (40-70%) whereas others (acetaminophen, naproxen, salicylic acid, etc.) are more efficiently removed (>70%) (Table 2).

Therefore, according to this criterion: all the β-blockers and diuretics are of big concerns as well as diclofenac, mefenamic acid, naproxen, ketoprofen and tramadol (anti-inflammatories and analgesics) and azithromycin, clarithromycin, erythromycin, metronidazole, sulfamethoxazole and trimethoprim (antibiotics).

Treatment of pharmaceutical wastewater: an issue to be addressed

To improve the removal rate in WWTPs, chlorination and UV irradiation are sometimes used after the biological treatment but the oxidizing power of both methods is not strong enough to degrade most of the pharmaceuticals, their metabolites or degradation products (Khetan and Collins 2007). Other physical and physicochemical methods such as coagulation-flocculation, activated carbon adsorption and membrane filtration can be used to remove pharmaceuticals from water (Luo et al. 2014). However, these treatments are separative methods, the pollutants are not destroyed but separated and transferred to another phase or produce concentrates that need to be further treated by other methods. In order to effectively degrade them, strong and non-selective oxidants are thus necessary.

Alternatively the so-called advanced oxidation processes (AOPs) have been developed to destroy efficiently toxic and/or persistent organic pollutants [START_REF] Pignatello | Advanced oxidation processes for organic contaminant destruction based on the Fenton reaction and related chemistry[END_REF]Oturan and Aaron 2014). Most of the AOPs are aqueous phase oxidation processes based on the generation of highly strong oxidant like hydroxyl radicals, the second strongest oxidant after fluorine, which can destroy nearly any type of organic contaminants thanks to its very high oxidation power (E° = 2.8 V/SHE) (Brillas et al. 2009;Sirés and Brillas 2012;Oturan and Aaron 2014).

Among different AOPs, electrochemical advanced oxidation processes (EAOPs) are of particular interest as they were shown to have a high removal efficiency and the capability to reach complete mineralization of the treated solutions (Rodrigo et al. 2014a;Sirés et al. 2014;Vasudevan and Oturan 2014). These processes are environmentally friendly and generate a great amount of hydroxyl radicals under specific operating conditions (Oturan et al. 2008a).

The electro-Fenton (EF) process is one of the most popular EAOPs coupling the chemical Fenton's process to electrochemistry. It has been used successfully in the treatment of a lot of pharmaceutical active ingredients and real wastewaters (Isarain-Chávez et al. 2011a;Barros et al. 2014a;Feng et al. 2014;Panizza et al. 2014;Yahya et al. 2014;García-Montoya et al. 2015).

However EAOPs become generally expensive when the objective is the quasi-complete mineralization of treated solutions because this operation needs long electrolysis time leading to a significant amount of electrical energy consumption [START_REF] Muñoz | Environmental assessment of different solar driven advanced oxidation processes[END_REF]Faouzi et al. 2006;Oturan et al. 2013). To reduce the operating cost and to make a cost-effective process, the combination of an EAOP with a biological treatment can be an appropriate solution to reduce the electrolysis time and to achieve the mineralization by a biological treatment (Scott and Ollis 1995;Olvera-Vargas et al. 2016a).

In this context, four pharmaceuticals belonging to different drug families (hydrochlorothiazide, ofloxacin, nadolol and tramadol) were selected to study their degradation and mineralization at lab scale by the EF process. The drug family and chemical structure of these molecules are provided in Table 3. They were chosen based on the following criteria: i) their occurrence in surface, hospital and industrial wastewaters, ii) their resistance to conventional treatment by WWTPs and iii) their toxicity. The objective was to find the best operating parameters to efficiently eliminate them and to study their degradation pathway. Then a pilot was built to scale-up the process from 230 mL to 1.7 L. The hydrodynamic and the efficiency of this reactor were studied. Finally, in the objective of decreasing the operating costs, a combination between the electrochemical process and a biological treatment was performed.

mixing and a good aeration. A design of experiment was also carried out on this pilot for the first time based on the liquid flow rate, the current density and the iron concentration and could evidence the main parameters for the mineralization efficiency and the energetic cost. An hydrodynamic study was performed as well as attempting to create a model to predict the percentage of mineralization in several locations of the reactor in the continuous mode. Those findings enable to give hints for further constructions of large scale reactors depending on the objective of the mineralization treatment. Finally, the combined bio-electro-Fenton process was also innovating as dealing with an already unworked molecule, the hydrochlorothiazide, and estimating the amount of energy that could be saved. It also highlights the specific conditions that should be applied in the electrochemical treatment to reach a good mineralization rate by explaining the role of the carboxylic acids, the aromatic intermediates and the pollutant in the biodegradation. A simple model is also presented to predict the inhibiting effects of the aromatic intermediates on the carboxylic acid biodegradation.

Introduction

The purpose of this review paper is first to give an overview of the newest knowledge concerning the EF process and the reactor configurations and then to provide an overview on its applications to the treatment of pharmaceutical compounds. Finally, the feasibility and the cost-effectiveness of the combined process bio-electro-Fenton (bio-EF) will be discussed and the whole literature about this topic will be provided.

Electrochemical advanced oxidation processes

Mechanisms involved in EAOPs

The most popular and commonly used electrochemical advanced oxidation processes (EAOPs) are anodic oxidation (AO), also called (electro-oxidation) and electro-Fenton (EF) processes.

These processes are based on direct (AO) or indirect (EF) catalytic generation of hydroxyl radical ( • OH) in the solution to be treated. Once generated in situ, these strong oxidizing agents will react on organic contaminants to oxidize them until their total or partial mineralization (Eqs. ( 1) and ( 2)).

Pollutants + • OH à intermediates (1)

Intermediates + • OH à à à CO2 + H2O + inorganic ions (2)
These types of processes have been well-described by several review papers (Brillas et al. 2009;[START_REF] Barrera-Dias | Electrochemical advanced oxidation processes : an overview of the current applications to actual industrial effluents[END_REF]Oturan and Aaron 2014;Sirés et al. 2014;[START_REF] Pliego | Trends in the intensification of the Fenton process for wastewater treatment: an overview[END_REF].

Direct oxidation (AO)

In this process the oxidation takes place at the surface of the anode. Heterogeneous hydroxyl radicals M( • OH)) are formed at the anode surface from the oxidation of water according to the following reaction (Eq. ( 3)) (Marselli et al. 2003).

M + H2O ® M( • OH) + H + + e - (3)
where M represents the anode material and M( • OH) is the hydroxyl radical adsorbed on the anode (M) surface.

For this kind of oxidation, two different types of transfer can limit the process: mass transfer or charge transfer. Mass transfer is done by migration (driven by the electrical field), by diffusion (driven by chemical potential gradient) and/or by convection processes. It can be expressed with the Nernst-Planck equation. As it is experimentally difficult to distinguish these three processes, the mass transfer process is generally approached in a semi-empirical way under specific conditions. These conditions are, i) the solution is conductive enough (use of a supporting electrolyte), ii) the solution is well-mixed and iii) the applied potential (or current) is high enough to oxidize efficiently the substrate.

Two different behaviors can occur depending on the anode material (Brillas et al. 2009). Active anodes promote the formation of higher states oxides or superoxides (MO) (Eq. ( 6)) [START_REF] Simond | Theoretical model for the anodic oxidation of organics on metal oxide electrodes[END_REF].

M( • OH) ® MO + H + + e - (6)
This type of anodes form chemisorbed radicals which limit the reaction at the surface of the anode as there is a strong interaction between the radicals and the anode surface. The most common anodes of this type are Pt and mixed metal or DSA (dimensionally stable anode)

anodes.

On the opposite, the non-active anodes enable the formation of physisorbed radicals with a quite weak interaction between hydroxyl radical ( • OH) and the anode surface (M) (Eq. ( 2)).

Hydroxyl radicals thus formed are able to react in the diffusion layer and can conduct to complete mineralization of organics. The most representative non-active anodes are BDD, tin or lead oxides and sub-stoichiometric titanium oxides.

AO process can also imply indirect (or mediated) oxidation occurring in the bulk of the solution thanks to the formation, at the electrode surface, of a mediator (an oxidizing species). Once migrated to the bulk solution, these oxidizing agents can participate in the oxidation of organic pollutants. Usually it concerns active chlorine species and peroxy anions. For instance, the relatively strong oxidants such as hypochlorous acid/hypochlorite anion (formed in chloride medium), peroxydisulfate anion (formed from sulfate medium) and peroxyphosphate anion (formed in phosphate medium) can be created [START_REF] Barrera-Dias | Electrochemical advanced oxidation processes : an overview of the current applications to actual industrial effluents[END_REF]) at the anode surface and contribute to the oxidation process in the bulk solution.

Indirect electro-oxidation: Electro-Fenton (EF) process

EF process is an indirect EAOP since hydroxyl radicals are not generated directly from charge transfer at the electrode level but in the solution from the well-known Fenton's reaction (Eq. ( 8)) [START_REF] Fenton | LXXIII.-Oxidation of tartaric acid in presence of iron[END_REF][START_REF] Haber | The Catalytic Decomposition of Hydrogen Peroxide by Iron Salts[END_REF]. In the Fenton's process, homogeneous hydroxyl radicals ( • OH) are generated from Fenton's reagent, a mixture of H2O2 and Fe 2+ , added externally to the solution to be treated. In contrast, for the EF process, the Fenton's reagent is electrochemically produced at the cathode. H2O2 is formed by a 2-electron reduction of dissolved O2 (Eq. ( 9)) and Fe 2+ by a single electron reduction of ferric cation (Eq. ( 10)). At the anode, water is oxidized to O2 (Eq. ( 11)) and, depending on the anode material, adsorbed hydroxyl radicals can be formed (Eq. ( 3)).

Fe 2+ + H2O2 ® Fe 3+ + OH + OH - (8) O2 + 2 H + + 2 e -® H2O2 (9)
Fe 3+ + e -® Fe 2+ (10)

2 H2O ® O2 + 4 H + + 4 e - (11)
The Fenton's process was a major step forward as it is easy to implement and uses relatively inexpensive and harmless chemicals [START_REF] Bautista | An overview of the application of Fenton oxidation to industrial wastewaters treatment[END_REF]Oturan and Aaron 2014). It was first applied in 1960 for the treatment of organic pollutants [START_REF] Brown | The reaction of Fenton's reagent with phenoxyacetic acid and some halogen-substituted phenoxyacetic acids[END_REF], since then a large application to the treatment of effluents was reported, such as the discoloring effluents of dye industries [START_REF] Kuo | Decolorizing dye wastewater with Fenton's reagent[END_REF] or the treatment of different kind of wastewaters [START_REF] Gogate | A review of imperative technologies for wastewater treatment I: Oxidation technologies at ambient conditions[END_REF].

Despite its interesting use in the treatment of effluents contaminated by organic pollutant, this process suffers from some drawbacks such as needing high amounts of Fenton's reagent to generate enough quantity of OH. The introduction of a high amount of H2O2 and Fe 2+ implies economic aspect and also the formation of ferric hydroxide sludge needing a supplementary processing. Moreover this situation favors the rate of wasting reactions (Eq. ( 12) and ( 13))

resulting in low efficiency.

OH + H2O2 ® H2O+ HO2 ( 12)

OH + Fe 2+ ® Fe 3+ + OH - (13)
The EF process, developed and popularized by Oturan's and Brillas' group at the beginning of the 21 st century has been designed to avoid the disadvantages of the classical Fenton's process by coupling between electrochemistry and Fenton's chemistry [START_REF] Brillas | Mineralization of 2,4-D by advanced electrochemical oxidation processes[END_REF][START_REF] Oturan | An ecologically effective water treatment technique using electrochemically generated hydroxyl radicals for in situ destruction of organic pollutants: Application to herbicide 2,4-D[END_REF]. Electrochemistry offers an amazing help in this respect (Vasudevan and Oturan 2014).

Indeed, in contrast to the classical Fenton's process, in the EF process the Fenton's reagent is in situ generated (H2O2) or regenerated (Fe 2+ ): H2O2 is generated on a suitable cathode by a 2electron reduction of dissolved O2 (Eq. ( 9)) and regeneration of Fe 2+ from Fe 3+ formed in Fenton reaction (Eq. ( 10)). The electrocatalytic regeneration of ferrous iron allows the use of a catalytic amount of this ion as catalyst thus avoiding the ferric hydroxide sludge formation and also preventing the wasting of generated OH (Eq. ( 13)) (Brillas et al. 2009;Oturan and Aaron 2014). Therefore, the use of EF process instead of Fenton's process allows the following advantages: (i) minimizing reagent cost by in situ generation of Fenton's reagent, (ii) avoiding risks related to transportation and storage of H2O2, (iii) avoiding (or minimizing) reactions wasting OH (Eqs. ( 12) and ( 13)) because H2O2 and Fe 2+ concentrations are quite low since they are consumed as soon as they are formed, (iv) eliminating sludge formation and (v)

generally providing high removal rate of organics due to the electrochemical regeneration of Fe 2+ which catalyzes Fenton's reaction [START_REF] Zhu | Optimization of Fenton and electro-Fenton oxidation of biologically treated coking wastewater using response surface methodology[END_REF]Rodrigo et al. 2014a;Sirés et al. 2014;Oturan and Oturan 2018).

The use of an appropriate cathode material is essential in the EF process; it should be able to generate efficiently H2O2 and have a low catalytic activity for its reduction. Moreover the cathode must have high overpotential for H2 evolution since the process is optimal at acid medium (Oturan and Oturan 2018). These characteristics are generally satisfied by carbonaceous materials that are often used in EF process such as three-dimensional carbon materials (carbon felt, graphite felt, carbon sponge, activated carbon fiber) (Oturan et al. 2000;[START_REF] Wang | Mineralization of an azo dye Acid Red 14 by electro-Fenton's reagent using an activated carbon fiber cathode[END_REF][START_REF] Özcan | Carbon sponge as a new cathode material for the electro-Fenton process: Comparison with carbon felt cathode and application to degradation of synthetic dye basic blue 3 in aqueous medium[END_REF][START_REF] Oturan | Metomyl degradation by electro-Fenton and electro-Fenton-like processes: a kinetics study of the effect of the nature and concentration of some transition metal ions as catalyst[END_REF][START_REF] Panizza | Degradation of Alizarin Red by electro-Fenton process using a graphite-felt cathode[END_REF]) and carbon-PTFE-O2 gas diffusion cathode [START_REF] Brillas | Electrochemical destruction of aniline and 4chloroaniline for wastewater treatment using a carbon-PTFE-O2-fed cathode[END_REF]. Other carbonaceous materials like graphite, reticulated vitreous carbon and carbon nanotubes have also been tested. More recently BDD and Ni foam have been also shown to be able to produce H2O2 [START_REF] Oturan | Metomyl degradation by electro-Fenton and electro-Fenton-like processes: a kinetics study of the effect of the nature and concentration of some transition metal ions as catalyst[END_REF][START_REF] Panizza | Degradation of Alizarin Red by electro-Fenton process using a graphite-felt cathode[END_REF][START_REF] Saussereau | Determination of levels of current drugs in hospital and urban wastewater[END_REF]Bocos et al. 2016a).

On the other hand, the nature of the anode is also very important in EF process. In the case of the use of a non-active anode, M( • OH) are produced on the anode surface (Panizza and Cerisola 2001). In this case EF process also includes AO process since homogeneous ( • OH) and heterogeneous M( • OH) are produced simultaneously in the bulk solution and on the anode surface, respectively, thus enhancing strongly the oxidation power of the process (Oturan et al. 2012). The best non-active anode material known is the BDD. However, its high cost can constitute an obstacle to its scale up for large-scale applications. Therefore, research is done on the use of ceramic electrodes based on sub-stoichiometric titanium oxides (TinO2n-1) [START_REF] Smith | Reviews in applied electrochemistry. Number 50 -Electrodes based on Magneli phase titanium oxides: the properties and applications of Ebonex (R) materials[END_REF][START_REF] Chaplin | Critical review of electrochemical advanced oxidation processes for water treatment applications[END_REF]. Recent studies have shown that this new anode material is able to provide results close to that of BDD in AO and EF processes [START_REF] Ganiyu | Use of Sub-stoichiometric Titanium Oxide as a Ceramic Electrode in Anodic Oxidation and Electro-Fenton Degradation of the Betablocker Propranolol: Degradation Kinetics and Mineralization Pathway[END_REF][START_REF] Oturan | Sub-stoichiometric titanium oxide as a new anode material for electro-Fenton process: Application to electrocatalytic destruction of antibiotic amoxicillin[END_REF].

During the EF treatment, numerous reactions are involved and the mass transfer constitutes an important parameter in the degradation of pollutant as some reactants have to reach the surface of the electrode (O2, Fe 3+ ) when others have to reach the bulk from the electrode surface (H2O2, Fe 2+ ). For instance, the oxygen required for the formation of H2O2 (Eq. ( 9)) has to be transferred from the gas phase to the aqueous phase and then needs to be adsorbed onto the electrode surface before its reduction process (Brillas et al. 2009). The required dissolved oxygen concentration is reached usually by aeration. If oxygen is not provided at an adequate rate, it becomes a limiting factor, controlling H2O2 generation rate and consequently the EF process.

As the diffusivity of oxygen in the gas phase is significantly higher than in the liquid phase, it is considered that the gas phase offers no transfer resistance. Thus, it can be suggested that the overall oxygen transfer rate (OTR) is controlled by the liquid phase, expressed as follows (Eq. ( 14)):

OTR = kLa (C * bulk, O2 -Cbulk,O2) ( 14
)
where OTR is the oxygen mass transfer rate (mol m -3 s -1 ), kLa is defined as the volumetric oxygen mass transfer coefficient (s -1 ), C * bulk, O2 and Cbulk,O2 are the equilibrium concentration and the oxygen concentration in the bulk (mol m -3 ), respectively.

It is worthy to note that the oxygen mass transfer has been well studied for water and wastewater treatment systems but not in EF process. The value of kLa depends mainly on the configuration of the reactors, on the air flow rate, on the airflow diffuser which will influence the bubbles size and on the hydrodynamics in the reactors.

The transfer from the bulk to the anode (pollutants) and the transfer from the cathode to the bulk (H2O2 and Fe 2+ ) are key points for the efficiency of the process. That is why, understanding the controlling parameter is important. As it was explained previously for the AO (Eqs. ( 4) -( 6)), the expression of the current density in EF follows the same behavior. But, the system is more complex as for the EF, the oxidation of pollutants can occur on the anode surface (AO)

and in the bulk (involving both homogenously generated • OH from Fenton reaction and mediated oxidation from oxidants formed at the anode).

There are several operating parameters involved in EF process efficiency which have now long been studied (Brillas et al. 2009;Nidheesh and Gandhimathi 2012;Oturan and Aaron 2014;Martínez-Huitle et al. 2015). The principal parameters are: solution pH, applied current, catalyst nature and concentration, supporting electrolyte (nature and concentration), electrode gap, oxygen (or air) supply rate and temperature.

The pH is a very important parameter of this process, it is usually recommended to work at a pH around 3 [START_REF] Boye | Degradation of herbicide 4-chlorophenoxyacetic acid by advanced electrochemical oxidation methods[END_REF]. The acidic medium is required for the formation of H2O2 (Eq. ( 9)), however pH values lower than 2 lead to the transformation of H2O2 to peroxonium (H2O3 + ) ion which is less reactive toward Fe 2+ in the formation of • OH via Fenton reaction and at higher pH this compound decomposes into water and oxygen [START_REF] Wang | Kinetic model and optimization of 2,4-D degradation by Anodic Fenton treatment[END_REF][START_REF] Feng | Degradation of azo-dye orange II by a photoassisted Fenton reaction using a novel Composite of Iron Oxide and Silicate nanoparticles as a catalyst[END_REF]Wang et al. 2010b). Moreover, at pH above 4, Fe 3+ precipitates as Fe(OH)3 and at pH below 1, Fe 2+ forms complexes with H2O2 [START_REF] Pignatello | Dark and photoassisted iron(3+)-catalyzed degradation of chlorophenoxy herbicides by hydrogen peroxide[END_REF]).

The current is another key factor in EF process since the rate of electrochemical reactions is promoting the generation of • OH (Eqs. ( 9), (10) and consequently Eq. ( 8)) and M(OH) (Eq.

(3)). Therefore this parameter influences significantly the oxidation and/or mineralization efficiency of the process [START_REF] Narayanan | Degradation of o-chlorophenol from aqueous solution by electro-Fenton process[END_REF]Oturan and Aaron 2014). But a too high current promotes competitive reactions consuming electrical energy in side reactions (reduction of water to H2 or its oxidation to O2) or wasting of • OH/M( • OH) (recombination or oxidation reactions) in the electrolytic cell (Zhang et al. 2007;Brillas et al. 2009;Sirés et al. 2014). The current value determines also the type of limitation for the system whether by the mass transfer or by the charge transfer as explained previously. Usually this parameter is experimentally determined for each compound/equipment before treatment.

Another highly important parameter is the nature or concentration of the catalyst. Among different catalysts tested for EF process, it was evidenced that iron (Fe 2+ /Fe 3+ ) is one of the best since it is not harmful, has low cost and requires low concentration (Brillas et al. 2009). With regard to the Fe 2+ concentration, the optimal value depends on the cathode used; optimal values about 0.1 -0.2 mM for carbon-felt cathode and 0.5 mM for gas diffusion cathode were reported many times in literature (Brillas et al. 2009;Oturan and Aaron 2014;Sirés et al. 2014). High Fe 2+ concentration is generally avoided since it promotes the wasting reaction given in Eq. ( 13) (Panizza and Cerisola 2001;Oturan and Aaron 2014). This factor is also usually experimentally adjusted. Meanwhile the Cu 2+ /Cu + couple was also investigated as catalyst in the EF process and contradictory results have been published. For instance a recent study, demonstrated that copper could be more efficient and more cost-effective than iron with a higher TOC removal rate [START_REF] Santana-Martínez | Electro-Fenton and electro-Fenton-like with in situ electrogeneration of H2O2 and catalyst applied to 4chlorophenol mineralization[END_REF]. But, other researchers highlighted that iron catalyst provides a better mineralization efficiency than cupper ions [START_REF] Oturan | Degradation of three pesticides used in viticulture by electrogenerated Fenton's reagent[END_REF].

Airflow and air diffuser are also important parameters that influence significantly both oxygen transfer and hydrodynamic conditions in the reactors. In the case of low oxygen flow rate the solution cannot be saturated in O2 and oxygen transfer can become the limiting step affecting the efficiency of H2O2 production rate and accordingly • OH generation rate through Fenton's reaction.

In addition, temperature is an influent parameter that impacts all the process. It affects the kinetics of chemical and electrochemical reactions and mass transfer parameters (diffusion coefficient, etc.). In general, temperatures ranging between 20 and 30 °C are applied in the literature [START_REF] Umar | Trends in the use of Fenton, electro-Fenton and photo-Fenton for the treatment of landfill leachate[END_REF]) since higher temperatures decrease O2 solubility as well as promote H2O2 decomposition. For the nature of the electrolyte, good conductivity ions must be used.

Indeed, Ghoneim et al. [START_REF] Ghoneim | Electro-Fenton oxidation of Sunset Yellow FCF azo-dye in aqueous solutions[END_REF]) performed an experiment comparing the efficiency of the EF with three different electrolytes: Na2SO4, NaCl and KCl. The better efficiency of Na2SO4 have been explained by the higher conductivity of SO4 2-compared to Cl - [START_REF] Gil | Electronic tongue for qualitative analysis of aqueous solutions of salts using thick-film technology and metal electrodes[END_REF]). They also obtained a higher degradation rate using KCl than NaCl explained again with the greater conductivity of K + compared to Na + . On the other hand, the degradation of methyl parathion by EF with sodium sulfate was slower than with NaClO4 or NaNO3 because of the formation of sulfato-complexes with iron [START_REF] De Laat | A comparative study of the effects of chloride, sulfate and nitrate ions on the rates of decomposition of H2O2and organic compounds by Fe(II)/H2O2and Fe(III)/H2O2[END_REF][START_REF] Diagne | Removal of methyl parathion from water by electrochemically generated Fenton's reagent[END_REF]). The electrolyte allows to increase the solution conductivity by avoiding an ohmic potential drop, responsible for a higher energetical consumption. Meanwhile it was demonstrated that a too high concentration of Na2SO4 can produce a great decrease in the TOC abatement which can be due to the side reaction of hydroxyl radicals with sulfate ions [START_REF] Zhou | Electro-Fenton method for the removal of methyl red in an efficient electrochemical system[END_REF][START_REF] Gökkuş | Oxidative degradation of basic black 3 by electro-generated Fenton's reagent using carbon fiber cathode[END_REF].

Reactor configuration

Recent ways of enhancing efficiency and decreasing the costs in EAOPs consist in choosing a specific reactor configuration to enhance mass transfer and current efficiency. The reactor configuration is a key point for the efficiency of the pollutant degradation by electro-Fenton.

Some review papers have given an overview of the different cell configurations used until 2012 (Brillas et al. 2009;Nidheesh and Gandhimathi 2012). Most of the studies have been performed

at lab-scale with undivided cell reactors with volumes lower than 300 mL. Another system, quite similar to the previous one (undivided cell), is the use of divided cells with two or three electrodes. The undivided system has the advantage of lower operating costs due to the possible use of a lower cell voltage in comparison with the divided cells. Moreover the latter has the advantage of avoiding H2O2 to reach the anode and to be oxidized there (Eq. ( 15)) thus preventing the following side reaction (Brillas et al. 2009).

H2O2 ® HO2 • + H + + e - (15)
Two other systems used to a lesser extent were the filter press reactors operating in batch and recirculation mode [START_REF] Prabhakaran | Treatability of resin effluents by electrochemical oxidation using batch recirculation reactor[END_REF]) and the bubble reactor operating in continuous mode [START_REF] Rosales | Electro-Fenton decoloration of dyes in a continuous reactor: a promising technology in colored wastewater treatment[END_REF]). However, in the last 5 years, new reactor configurations or enhanced configurations of formerly used reactors were developed that allows having a greater efficiency and/or lower costs (Martínez-Huitle et al. 2015). Therefore, the aim of this section is to give an overview of those new studies which are sum-up in Table 4. Some of these configurations are also illustrated in Fig. 1. The novelties include: i) the way of mixing in order to favor mass transport and transfer to the anode, ii) the manner in which the oxygen is transferred to the solution, iii) the shape and distance between electrodes and their number, iv) the shape and the volume of the cell and v) the type of out streaming.

Geometric parameters

Recent research has most specifically focused on geometric changes in electrodes. Among them, Su et al. [START_REF] Su | Comparison of aniline degradation by Fenton and electro-Fenton reactors using plate and rod electrodes[END_REF]) studied the impact of using plate or rod electrodes for aniline removal (Fig. 1). The experimental conditions are described in Table 4. They found out that the plate electrodes have a better aniline degradation rate than the rod electrodes for added [H2O2] > 29 mM. This phenomenon is explained as the distance between the electrodes is lower (1.5 cm) for plate electrodes while it is significantly higher (5.5 cm) for rod electrodes. This allows the plate electrode system to have a lower electrical resistance and thus a better electrical efficiency. On the contrary, in the case of lower [H2O2] added to the solution, the rod electrode cell is more efficient at degrading the pollutant [START_REF] Su | Comparison of aniline degradation by Fenton and electro-Fenton reactors using plate and rod electrodes[END_REF]. The area of the working calculation. The parallel configuration of the electrodes in reactor 1 can explain the lower energy consumption. Indeed, this configuration reduces by two the current intensity which results in a lower voltage for the cell and thus in a lower energy consumption compared with reactor 2 where the current is not divided. The surfaces of the electrodes facing each other in reactor 3 are smaller causing a lower conductance and thus an increase of the cell voltage.

Another cell geometry has been proposed by [START_REF] Lei | Development of a trickle bed reactor of electro-Fenton process for wastewater treatment[END_REF]): the trickle-bed reactor (Fig. 2) to enhance the efficiency of the electro-Fenton process. This system is made of several pieces which constitute the fixed part of the packed bed of the reactor using the downward movement of a liquid and the upward movement of gas. Details about the reactor configuration and the treatment conditions are given in Table 4. The air and the solution are both mixed before entering the cell. This configuration enables to dissolve the oxygen first in the liquid allowing the easy transfer of oxygen to the electrode surface and therefore enhancing the production of hydrogen peroxide. The authors found out that this system is as efficient as GDE reactor but avoids electrolyte leakage and gas bubbles which makes its application feasible at a larger scale [START_REF] Lei | Development of a trickle bed reactor of electro-Fenton process for wastewater treatment[END_REF]. Finally, Xu et al. [START_REF] Xu | Electrogeneration of hydrogen peroxide using Ti/IrO2-Ta2O5 anode in dual tubular membranes Electro-Fenton reactor for the degradation of tricyclazole without aeration[END_REF] suggested the use of tubular membranes for both anode and Fig. 3: Scheme of the dual tubular membranes EF reactor. Reprinted with permission from ref. [START_REF] Xu | Electrogeneration of hydrogen peroxide using Ti/IrO2-Ta2O5 anode in dual tubular membranes Electro-Fenton reactor for the degradation of tricyclazole without aeration[END_REF]. Copyright 2016, Elsevier.

Another tubular system was recently created. It is based on two driving forces: i) the transmembrane pressure and ii) the current. The authors demonstrated the efficiency of their system, described in Table 4 (tubular membrane A) by the higher H2O2 production of their system by finding optimums current and transmembrane pressure (Olvera-vargas et al. 2018).

Out streaming

As in the study on tubular membranes, research was done on the effect of the direction of the flow whether parallel to the electrodes (flow-by) or perpendicular to them (flow-through or vertical-flow). Two recent studies analyzed the degradation of methylene blue and tartrazine, respectively, under those different flowing conditions. The single compartment system working on methylene blue is composed of perforated DSA (dimensionally stable anode) electrode as the anode and a modified graphite felt as a cathode (Fig. 4) [START_REF] Ma | A highly energy-efficient flow-through electro-Fenton process for organic pollutants degradation[END_REF]. The second multiple cell reactor, used to degrade tartrazine, is equipped with PbO2 anodes and modified graphite felt mesh cathodes (Ren et al. 2016). Results obtained showed the superiority of the flow-through reactor over the flow-by reactor in the same conditions (Table 4). This result can be explained by the enhancement of the mass transport in the vertical-flow compared to the parallel flow [START_REF] Santos | Characterization of fluid dynamics and mass-transfer in an electrochemical oxidation cell by experimental and CFD studies[END_REF][START_REF] Liu | Reactive transport mechanism for organic oxidation during electrochemical filtration: Mass-transfer, physical adsorption, and electron-transfer[END_REF]. The higher efficiency is related to the greater accumulation of H2O2 in the flow-through (57.8 mg L -1 ) than in the flow-by system (51.9 mg L -1 ) [START_REF] Ma | A highly energy-efficient flow-through electro-Fenton process for organic pollutants degradation[END_REF].

closer the electrodes, they could increase the COD abatement and decrease the energetical consumption. The specific energy consumption was of 50 and 30 kWh (kgCOD) -1 for 1 pair of electrode separated by 10 and 2 mm, respectively, after 6 h electrolysis [START_REF] Hu | Effect of the structure of stacked electro-Fenton reactor on treating nano filtration concentrate of landfill leachate[END_REF].

Fig. 5: Scheme of the 10-compartment reactor. Reprinted with permission from ref. (Ren et al. 2016).

Copyright 2016, Elsevier.

Special way of supplying O2

Recently an interesting idea appeared in numerous publications: a special way of providing oxygen to the system without supplying an air or oxygen flow rate. This system is based on the use of two parallel rotating cathodes with an anode in the middle (Fig. 6) [START_REF] Yu | A novel electro-Fenton process with H2O2 generation in a rotating disk reactor for organic pollutant degradation[END_REF]. The operating parameters related to this reactor are described in Table 4. This system enables to bring oxygen without blowing air or O2 in the solution making it more cost-effective and providing a good mass transfer between the solution and the electrode surface as demonstrated by comparing with the same electrodes but without rotating cathodes and providing O2 (Table 4) [START_REF] Yu | A novel electro-Fenton process with H2O2 generation in a rotating disk reactor for organic pollutant degradation[END_REF]. Another two-compartment reactor avoiding the use of oxygen is based on the catalytic generation of H2O2 (Eq. ( 16)) thanks to Pd/C particles. The design of this novel reactor is based on the work reported by Liu et al. [START_REF] Liu | A novel electro-Fenton process for water treatment : reaction-controlled pH adjustment and performance assessment[END_REF]) reporting an electrolytic system able to generate different pH environment without adding any chemicals by using water electrolysis to create O2 and H2. A low pH was first used to produce H2O2 and then a high pH was employed to neutralize the treated solution before its release to the natural water stream [START_REF] Liu | A novel electro-Fenton process for water treatment : reaction-controlled pH adjustment and performance assessment[END_REF]).

The characteristics of the reactor are described in Table 4. The first compartment enables the formation of H2, O2 and H + by water electrolysis, of H2O2 (Eq. ( 16)) by the catalytic reaction thanks to Pd/C and finally leading to hydroxyl radicals from the Fenton's reaction.

H2 + O2 ® H2O2 (16)

The second compartment accumulates OH -which is utilized to neutralize the solution when needed. The main advantage of this system is to avoid chemicals to set the solution pH and to obtain a neutral solution after treatment that can be directed directly to a bioreactor. To illustrate the efficiency of their reactor, the authors compared the production of H2O2 and its accumulation to some values reported in the literature and concluded that it could produce more H2O2 than carbon fiber or carbon nanotubes cathodes but less than a GDE cathode. This reactor presents some drawbacks such as a high voltage required due to the electrical resistance of the salt bridge and the high price of the catalyst; however it was shown to be reusable [START_REF] Yuan | Pd-catalytic in situ generation of H2O2 from H2 and O2 produced by water electrolysis for the efficient electro-fenton degradation of rhodamine B[END_REF]. In order to overcome such drawbacks another system was created by Zhang et al. Their system produced O2 and an acidic medium thanks to the use of a second anode where water is oxidized in O2 and H + . This system is promising by avoiding the use of O2 and expensive material such as Pd [START_REF] Zhang | A three-electrode electro-Fenton system supplied by self-generated oxygen with automatic pH-regulation for groundwater remediation[END_REF].

One of the latest study on a new way of providing O2 was proposed by Perez et al., [START_REF] Pérez | Electrochemical jet-cell for the in-situ generation of hydrogen peroxide[END_REF]) which consisted of a combination of a flow-through system and a Venturi-based jet aerator (Fig. 7) which allows an efficient production of H2O2 and a very good oxygen supply with low energy costs. This system is based on the Venturi effect as in the thinner part of the tube a depression is created which sucks the air into the system leading to a higher liquid velocity which enables to break the oxygen gas bubbles and thus helps their dissolution in water. This system is compared to a similar reactor using GDE cathode and a flow-by cell system. The reactor with Venture-based jet aerator can produce 960 mg H2O2 L -1 whereas the reactor with GDE cathode produces 700 mg H2O2 L -1 (Table 4). It is also showed that the efficiency of such aeration is higher than the same reactor without oxygen supply or with air sparging to the jet aerator. The concentration of H2O2 obtained at a current density of 75 mA cm -3 after 30 min, is very low for no oxygen supply and, 8 and 400 mgH2O2 L -1 for air sparging and jet aerator, respectively [START_REF] Pérez | Electrochemical jet-cell for the in-situ generation of hydrogen peroxide[END_REF].

Fig. 7: Scheme of the jet-cell reactor. Reprinted with permission from ref. [START_REF] Pérez | Electrochemical jet-cell for the in-situ generation of hydrogen peroxide[END_REF]. Copyright 2016, Elsevier.

Micro-fluidic

A different type of reactor was recently used for electro-Fenton processes: microfluidic reactors [START_REF] Scialdone | Electro-generation of H2O2 and abatement of organic pollutant in water by an electro-Fenton process in a microfluidic reactor[END_REF][START_REF] Scialdone | Abatement of acid orange 7 in macro and micro reactors. Effect of the electrocatalytic route[END_REF][START_REF] Sabatino | Electrochemical abatement of organic pollutants in continuous-reaction systems through the assembly of microfluidic cells in series[END_REF]. It was first used by [START_REF] Scialdone | Electro-generation of H2O2 and abatement of organic pollutant in water by an electro-Fenton process in a microfluidic reactor[END_REF] who studied the capacities of this type of reactor compared to a conventional undivided macro reactor. The two different reactors are described in Table 4. The micro-reactor had the following advantages: i) no use of electrolyte as there is a very small transfer resistance due to the very small interelectrode gap ii) no need of gaseous stream as oxygen is provided by water oxidation at the anode. The use of a graphite cathode provides a cheap and easy to handle system. The chemical oxygen demand (COD) is used as comparison parameter. As can be seen in Table 4, the abatement is more than two times higher for the microfluidic reactor compared to the macro reactor. The better efficiency of the micro reactor is explained by a comparison of H2O2 production which is 0.6 mM and 6 mM for the macro and microreactor, respectively.

Besides, the calculated current efficiency (CE) values also highlight better performance of the microreactor (8% for the macro-reactor against f 33% for the micro-reactor at 10 A m -2 ). This fact can be explained by high solubility of O2 in micro conditions: it is formed by water oxidation at the anode and dissolved close to its maximum solubility. As the system is small, the O2 is present at a high concentration at the cathode (both electrodes being close to each other) and leads to the formation of a higher concentration of H2O2 [START_REF] Scialdone | Electro-generation of H2O2 and abatement of organic pollutant in water by an electro-Fenton process in a microfluidic reactor[END_REF].

However, the very small size of the channels could be problematic since they can be plugged when real effluents with organic and mineral particles are treated. The group of Scialdone investigated also the performances of the different advanced oxidation processes for a microfluidic treatment and concluded that combining EF and electro-oxidation could significantly be enhanced by microfluidic [START_REF] Scialdone | Abatement of acid orange 7 in macro and micro reactors. Effect of the electrocatalytic route[END_REF]). The last study on microfluidic system applied to EF dealt with the use of micro-reactor cells in series. The authors concluded that using three EF reactors in series was not useful compared to the use of one micro-reactor as the energy consumption value increases much more than the TOC abatement rate for the reactors in series [START_REF] Sabatino | Electrochemical abatement of organic pollutants in continuous-reaction systems through the assembly of microfluidic cells in series[END_REF].

Electrical energy consumption (EC)

In order to compare the cost effectiveness of different systems, their energetical treatment costs were estimated. For instance, the EC were of 0.5 and 20 kWh (gTOC) -1 for the system with rotating cathodes [START_REF] Yu | A novel electro-Fenton process with H2O2 generation in a rotating disk reactor for organic pollutant degradation[END_REF]) and for the system with the two cathodes in the center [START_REF] Yu | A novel dual gas diffusion electrodes system for efficient hydrogen peroxide generation used in electro-Fenton[END_REF], respectively. This can suggest that the rotating cathode is a more cost-effective system. Meanwhile this treatment is done in batch so to well-evaluate the real cost of the treatment, experiments in continuous mode are required. In the case of flow-through reactors [START_REF] Ma | A highly energy-efficient flow-through electro-Fenton process for organic pollutants degradation[END_REF]Ren et al. 2016), the EC could be evaluated using a modified formula (Eq. ( 17)) in order to take into account the continuous mode.

EC = ! × $ (&'( ) * &'( + ) ×- (17)
with E, the cell voltage in Volt, I the current in Ampere, TOC0, the initial TOC, TOC(t), the TOC at the steady state, D the flow rate in L h -1 .

Using this equation, it was found that the flow-through multiple system (Ren et al. 2016) is consuming 140 kWh (g TOC) -1 whereas the flow-through single system [START_REF] Ma | A highly energy-efficient flow-through electro-Fenton process for organic pollutants degradation[END_REF]) is consuming 4 kWh (g TOC) -1 so 35 times less. Meanwhile an important point has to be underlined, the amount of treated pollutant was of 20 mg per h for the single system against 240 mg per h for the pilot (multiple system). So the gap is reduced by taking this point into account. Nevertheless the consumption is higher in the pilot which can be explained by the scale up of the treatment as the single system is done in 150 mL. However, it is at this step impossible to compare experiments done on the different reactor configurations (different pollutants, different initial TOC configuration, different flow mode, etc.). Much more researches are needed on this aspect and the EC should be estimated for each configuration under comparable conditions and considered as a key parameter. After 120 min of treatment, the TOC removal was 57.9% and 39.1% for the flow-through and flow-by system respectively with 50 mA, a pH of 3, a flow rate of 7 mLmin -1 , 0.3 mM of Fe 2+ and an initial concentration of pollutant 50 mg L -1 H2O2 was 566, 531 and 487 mg L -1 , the CE was 85.4, 75.8 and 70.2 % in 180 min and the energy consumption was 8.6, 14.7 and 24 kWh kg -1 H2O2 for reactor 1, 2 and 3 respectively under 0.05 M of Na2SO4, at a current intensity of 100 mA and at a pH of 7.

(Yu et al.

2015)

Venturibased jet aerator Rhodamine B initial concentration (10 mg L -1 ) decreased to 1.6 mg L -1 after 60 min, with 1 mM of Fe 2+ using 50 mA for each cathode.

(Yuan et al.

2011)

Reactor with rotating cathodes Fig. 6 100 mL (reactor) 1.5 cm 1 anode 2 cathodes Anode: Pt or DSA Cathodes: graphite felt Batch mode

Rotation of the cathodes

No A solution of 50 mg L -1 of methyl orange, 0.05 M Na2SO4, 0.2 mM Fe 2+ , pH of 3, after 20 min the pollutant was degraded using 2.37 kWh m -3 for the rotating cathodes at 10 rpm and in 30 min using 3 kWh m -3 when using the rotation of the cathodes and providing 0.2 L min -1 of O2. [START_REF] Yu | A novel electro-Fenton process with H2O2 generation in a rotating disk reactor for organic pollutant degradation[END_REF] a : MF, micro-fluidic reactor b : MR, macro reactor

Hints for scale-up

The different parameters stated in this section are also very important in order to scale-up the process to treat real effluent at an industrial scale. The already published research and review articles gave some hints to scale-up the process concerning i) the size and organization of the electrodes, ii) the temperature, iii) the mixing, iv) the out streaming and v) the continuous mode.

In order to treat larger volumes, the system must be larger. Regarding the electrodes, two options are possible; either using high surface electrodes or using a stack system. The use of bigger electrodes produces a non-uniform repartition of the current on the electrode surface which reduces the yield. By using stacking electrodes, the cost is higher so a compromise has to be reached. When stacking is used the electrode configuration is also important, the configuration will be either in monopolar or bipolar. In monopolar connection, all anodes and with a current of 15 mA cm -2 and 20 mA cm -2 obtaining 50.3% and 76.2% of COD removal with a BDD anode for the mixed cell and flow cell, respectively. To improve the mass transport of mixed tank reactors, a recirculation can be added; this will not only improve mass transport efficiency but will also drag gas bubbles which otherwise would accumulate increasing the Ohmic resistance of the bulk and finally rising the energy consumption (Martínez-Huitle et al.

2015)

. The type of flow is also a key point for the scaling-up as it is directly related to the efficiency of the process, as seen before, the flow-through reactor provides better performance than the flow-by reactor. [START_REF] Radjenovic | Challenges and opportunities for electrochemical processes as next-generation technologies for the treatment of contaminated water[END_REF] proposed the combination of this type of reactor with three-dimensional electrodes in order to obtain the better cell system for a large scale treatment. Computational fluid dynamics are recommended to optimize the flow system, especially at the entrance and at the exit of the cell where turbulences often happen.

Nonetheless, only a few numbers of studies were reported on this tool [START_REF] Frías-Ferrer | Optimized design of an electrochemical filter-press reactor using CFD methods[END_REF][START_REF] Cruz-Díaz | The FM01-LC reactor modeling using axial dispersion model with a reaction term coupled with a continuous stirred tank (CST)[END_REF][START_REF] Rivero | Numerical simulation of mass transport in a filter press type electrochemical reactor FM01-LC: Comparison of predicted and experimental mass transfer coefficient[END_REF][START_REF] Vázquez | CFD evaluation of internal manifold effects on mass transport distribution in a laboratory filter-press flow cell[END_REF]. Temperature belongs to the influent parameters in the reactor design, but it is often not studied as at bench scale there is no need to cool the system. However, at large scale applications, this parameter has to be taken into account and a cooling system can be involved.

For industrial applications level, the continuous mode has to be reached to enable the treatment of large amounts of wastewater. The stability of the electrodes is to be considered as this can directly pollute the solution or can decrease the efficiency of the process. Some studies evaluated the steady efficiency through several cycles (Rosales et al. 2012b;Bocos et al. 2016c;[START_REF] Plakas | Removal of organic micropollutants from drinking water by a novel electro-Fenton filter: Pilot-scale studies[END_REF] or by inspecting the electrode surface thanks to Scanning Electron Microscopy (SEM) after the treatment [START_REF] Iglesias | Electro-Fenton decolourisation of dyes in an airlift continuous reactor using iron alginate beads[END_REF]. Most of the works describing a continuous electro-Fenton process are given in Table 5. Most of the experiments summarized in this table are carried out in heterogeneous system which avoids the formation of sludge and requires of a filtration post-treatment to eliminate or recover the catalyst (Brillas et al. 2009).

Iron alginate gel beads are often used as solid catalysts aiming at degrading for example 81%

of Lissamine Green B [START_REF] Iglesias | Electro-Fenton decolourisation of dyes in an airlift continuous reactor using iron alginate beads[END_REF]), 90% of Imidacloprid [START_REF] Iglesias | Electro-Fenton oxidation of imidacloprid by Fe alginate gel beads[END_REF]) and 99%

of Azure B (Rosales et al. 2012b) (Table 5). The economic aspect of these reactors is often left aside, only a few studies considered it. With 20 A, 25.5 L of industrial effluent is degraded consuming 0.14 kWh (g TOC) -1 [START_REF] Casado | Pilot-scale degradation of organic contaminants in a continuous-flow reactor by the helielectro-Fenton method[END_REF], with 5 V, 0.15 L of Lissamine Green B is partly mineralized (COD removal of 86%) costing 15.75 kWh (kg dye) -1 (Bocos et al. 2016c), with 3.5 V, 3 L of Rhodamine B is treated reaching 98% of dye degradation consuming 50 Wh (g dye) -1 (Nidheesh and Gandhimathi 2015) and with 0.14 A, 1 L of C.I.

Acid red and blue are degraded using 0.024 kWh (g COD) -1 [START_REF] Pajootan | Discoloration of wastewater in a continuous electro-Fenton process using modified graphite electrode with multi-walled carbon nanotubes/surfactant[END_REF]) (Table 5).

However, only a few articles reported the continuous mode; more work has to be done and especially treatments with real wastewaters need to be studied. Industrial water from a chemical firm and a pulp and paper company [START_REF] Casado | Pilot-scale degradation of organic contaminants in a continuous-flow reactor by the helielectro-Fenton method[END_REF], landfill leachate [START_REF] Zhang | Electro-Fenton treatment of mature landfill leachate in a continuous flow reactor[END_REF]) and real wastewater from Mexico (Durán [START_REF] Moreno | Electro-Fenton as a feasible advanced treatment process to produce reclaimed water[END_REF]) are successfully degraded by EF in continuous mode. Thus the TOC removal is 80 and 60% for the effluents from chemical firm and the pulp and paper company, respectively, the COD removal was 60%

for the landfill and 92.3% discoloration is obtained for the wastewater from Mexico (Table 5).

In synthetic effluents, the concentration of pollutants is indeed higher than in real wastewaters, the interaction of different pollutants together and the effect of mineral ions as well as the effect of natural substances such as humic acids are not yet evaluated.

EF applied to the removal of pharmaceuticals from water

EF applied to β-blockers

Table 6 gives an overview of EF applied to β-blockers for the last few years. Only three βblockers have been studied, atenolol, metoprolol and propranolol. They are always degraded in synthetic aqueous solution, with only one study evaluating the synergetic effect of using a solution composed of several β-blockers (Sirés et al. 2010). The concentration of the pollutants ranges from 26 mg L -1 to 150 mg L -1 . Most of these molecules require between 5 and 30 min to be completely oxidized but the mineralization of their solutions takes several hours of electrolysis. This difference of treatment time can be explained by involving several steps in the mineralization procedure while the oxidative degradation occurs in one step. Another reason can be the formation of intermediates products more difficult to oxidize than the initial pollutant, in the latter case.

The mineralization current efficiency (MCE) reported for these experiments is depending on the treatment time and is always less than 50%; it is partly due to side reactions occurring during the process. A link between the chemical structure and the apparent kinetic rate constant of degradation can be underlined thanks to Isarain-Chávez et al. (2011c) who performed the degradation of three beta-blockers under the same conditions. Atenolol has the higher degradation rate constant which is about 10 times greater than the value reported for metoprolol and propranolol. This difference can be explained by the chemical structure of atenolol containing an amide group which is more electrophilic than the groups contained in metoprolol and propranolol. One of the best pseudo first order kinetic constant is obtained for atenolol degradation with a value of 9.5 10 -3 s -1 whereas the other constants are around 10 -3 s -1 (Fig. 8).

In spite of this high rate constant value, the electrical energy consumption is significantly higher

(1.08 kWh (g TOC) -1 ) than other b-blockers having a value of around 0.30 kWh (g TOC) -1 . This can be related to the formation of hardly oxidizable intermediates in the former case. A compromise should then be reached between an efficient degradation and a cost-effective operating process. More studies taking into account the cost of the process are required in order to better evaluate the feasibility of this process.

L h -1 3 A 3-0.4 A 3-0.4 A - - -
in 360 min:

(1) 54% ( 2) 54% (3) 56%

(1) 26%

(2) 22%

(3) 35%

(1) 61%

(2) 66%

(3) 65%

(1) 2.24 10 -3 s -1

(2) 0.22 10 -3 s -1

(3) 0.13 10 -3 s -1

(1) 0.36 10 -3 s -1

(2) 0.18 10 -3 s -1

(3) 0.15 10 -3 s -1

(1) 2.43 10 -3 s -1

(2) 0.74 10 -3 s -1

(3) 0.91 10 -3 s -1

Between:

(

EF applied to anti-inflammatories and analgesics

Seven anti-inflammatories and analgesics are studied by the EF process. Table 7 gives some details about these experiments and their treatment efficiency. The efficiency of the treatment measured by the TOC removal degree varies significantly as it was 57% in 8 h for salicylic acid with graphite plates as electrodes whereas it reached 95% in 1.5 h for ketoprofen using BDD and carbon felt as anode and cathode, respectively. Concerning the pseudo first order kinetics, the apparent rate constant (kapp) values range from 10 -4 to 10 -3 s -1 (Fig. 8). Finally, as for betablockers, the best way to identify the usefulness of the process is the energetic consumption which is available for only five studies, the more cost-effective among them appears to be dipyrone with the DSA-Cl2/GDE system with 0.049 kWh (g TOC) -1 and with a kapp of 1.2 10 -

EF applied to antibiotics and diuretics

More studies were done on antibiotics as they are a large group of pharmaceuticals widely used all over the world. The kapp values are in the same range for beta-blockers and antiinflammatories which demonstrates the ability of this process to treat pharmaceuticals (Fig. 8).

The better value (26 10 -3 s -1 ) is obtained for flumequine degradation using a BDD/carbon felt cell. 17 different pharmaceuticals were degraded through an EF process (Table 8). The treatment time for complete oxidative degradation of pollutant is quite different from one system to another; it can be of 4.5 h for cefalexin or of 3 min for flumequine; it is around 30 min for most cases. The TOC or the DOC (dissolved organic carbon) removal is around 6 h for most of the experiments. Concerning the MCE, for almost all the experiments the values are around 15%, only in one study, this value reached up to 55% for chloramphenicol, due probably to the very high initial pollutant concentration (245 mg L -1 ). The energy consumption is also in the same range as for the other pharmaceuticals being around 1 kWh (g TOC) -1 or 1 kWh (g DOC) -1 .

To sum up, the EF is a clearly efficient process able to degrade lots of different groups of pharmaceuticals. Some points have yet to be studied more deeply. First in order to apply this process at the industrial scale, pilot studies are required to treat issues that have been explained

previously like, for instance, mass transfer. Our group is actually working on a pilot in order to evaluate the scaling-up of the EF. To prepare the transition to industrial scale, the EF needs to be studied in continuous mode; this has only been done a few times (Table 5). The concentration of the pollutant ranges from around 50 mg L -1 to 150 mg L -1 which is significantly higher than the concentration in wastewater as discussed in the introduction section. In order to simulate the efficiency of the process better and in a more similar context of real wastewater treatment, smaller concentrations should be involved or real effluent should be used which has only been done a few times (Durán [START_REF] Moreno | Electro-Fenton as a feasible advanced treatment process to produce reclaimed water[END_REF][START_REF] Casado | Pilot-scale degradation of organic contaminants in a continuous-flow reactor by the helielectro-Fenton method[END_REF]. Indeed, in a real effluent, due to the presence of several pharmaceuticals, some interactions between them can slow down the process. As the cost remains generally high for mineralization process, combining this process with a biological treatment could decrease significantly the operating costs. Therefore the cost of the process is to be considered more deeply. The bio-treatment being more cost-effective than the EF, a comparison of their efficiency was done by analyzing their apparent kinetic constant for the same pharmaceuticals (Fig. 8). For all the data given in Fig. 8, the EF is significantly more efficient. However, the scale of the process is not the same as for the bio-treatment an industrial scale is used whereas for the electro-Fenton the volume is around 300 mL. But, for the biological degradation, the limiting parameter is the degradation kinetic, this implies that the scale-up is not impacting too much on the process; this enables the comparison between both processes. Because of this difference of efficiency, the solution to degrade efficiently and at a lower cost can be to combine both processes.

2.E+03

A (2) 75% (1) 95%

T E N O L O L M E T O P R O L O L P R O P R A N O L O L D I P Y R O N E I B U P R O F E N K E T O P R O F E N N A P R O X E N S A L I C Y L I C A C I D A M O X I C I L L I N C H L O R A M P H E N I C O L C H L O R O X Y L E N O L C I P R O F L O X A C I N E N R O F L O X A C I N L E V O F L O X A C I N R A N I T I D I N E S U L F A M E T H A Z I N E S U L F A N I L A M I D E Apparent kinetic constant (d-1) EF treatment Bio-treatment
(3) 73% in s -1 4.2*10 -3 (1) 4.5 10 -3 (2) 7.2 10 -3 (3)

1.2 10 -3 (1) 1.0 10 -3 (2) 3.7 10

(2) 81% (3) - - 3.0 10 -3 s -1 (1) 2.0 10 -3 s -1 (2) 2.3 10 -3 s -1 (3)
After 6 h in kWh (gTOC) -1 : 1.30 (1); 3.20 (2); 1.10 ( 3) -

Bioelectro-Fenton: A combined process for application to treat pharmaceuticals

Conventional biological processes in WWTP do not always provide satisfactory results in terms of pharmaceuticals removal (Miège et al. 2009;Luo et al. 2014;Margot et al. 2015), since many of these organic substances are toxic or resistant to microorganisms. Therefore, one feasible option to remove the persistent organic compounds is to combine biological treatment with EAOPs such as EF. This attractive potential alternative is a solution to reduce the global cost which increases with treatment time in EF.

AOPs as a pretreatment for biodegradability improvement

Two properties of the solution are considered in the literature to assess the feasibility of a biological treatment: biodegradability and toxicity.

Biodegradability of the solution is a parameter of great importance when combined treatments are considered. It is often defined thanks to the ratio of BOD5/COD, with BOD5 the biochemical oxygen demand at 5 d. Usually, when this ratio is greater than 0.4, the solution is considered as biodegradable [START_REF] Pulgarin | Strategy for the coupling of photochemical and biological flow reactors useful in mineralization of biorecalcitrant industrial pollutants[END_REF]. Less commonly, one can also find the ratio of BOD14/COD in the literature, BOD14 being the biochemical oxygen demand at 14 d. The limitation of the BOD test is mainly due to the fact that micro-organisms used as inoculum in the BOD tests may not be initially adapted to transform the organic compounds present. Thus, a lag phase is often observed during the first days in the BOD curve that can lead to very low values of the BOD after 5 d. However, after this lag phase, degradation of the organic compounds can occur leading to a strong increase in the BOD curve. On the contrary, this value can also be under estimated when they are different compounds with different biodegradability levels. So, to overcome this potential underestimation of the biodegradability of the solution, the ultimate BOD (BODu) can be used (Trellu et al. 2016a). Another bias of the use of BOD test is that in many papers, the concentration used in the BOD test is not the same as in the EF process. Thus, as the rate of biodegradation appears to vary with concentration, the inhibiting effect of the organic compounds can be underestimated. For these two main reasons, results of BOD tests should be interpreted with caution.

Another way to indirectly measure the biodegradability of organic effluents which was used in the most recent studies is the ratio between COD and TOC. Indeed, a decrease in COD (including oxidation of the pollutant) can promote the decrease of the solution toxicity. A decrease of the ratio COD/TOC is a favorable trend for biological treatment as there are fewer oxidative products (pollutant) and the amount of organic carbon is sufficient to feed bacteria [START_REF] Chebli | Supported photocatalysis as a pre-treatment prior to biological degradation for the removal of some dyes from aqueous solutions; acid red 183, Biebrich Scarlet, methyl red sodium salt, orange II[END_REF][START_REF] Annabi | Degradation of enoxacin antibiotic by the electro-Fenton process: optimization, biodegradability improvement and degradation mechanism[END_REF]. From these two parameters, the AOS (average oxidation state) can be estimated which is given in Eq. 18. This parameter enables to approximate the degree of oxidation of a solution. The lowest value is -4 for CH4 and 4 for CO2 (Scott and Ollis 1995).

!"# = % (()*+*),)

()* (18)
with TOC in mg C L -1 and COD in mg O2 L -1 .

Measuring the toxicity can also be interesting. However, there are usually errors in the interpretation of results. Several methods are used to measure the toxicity of solutions electrochemically pre-treated. In general, ecotoxicological bioassays standards are used. The most commonly used is the bacterial assay Microtox® which is based on natural or provided ability of bioluminescence of a bacterial strain. However, as it is a marine strain, doubts are arising due to its application as a reference when a combined treatment is considered. It should be highlighted that the toxicity measured is not representative of the toxicity of adapted consortia that are usually employed when biological treatments are considered. Thus with this method, a high toxicity value does not mean that all the micro-organisms from a consortia will be inhibited or killed by the solution.

The choice of the relevance of the combined process and the order in which to apply it, whether the biological treatment first and the AOP after or the contrary, are in general based on the biodegradability and toxicity of the solution to be treated. A first methodology was proposed by Scott and Ollis (1995) who divided the wastewaters into four types. The first group gathers large-size molecules that are not biodegradable (polymers, pesticides, herbicides, emulsifying agents,…). In this case, a chemical treatment is advised first to hydrolyze the big molecules into smaller biodegradable molecules to allow their biodegradation. [START_REF] Mantzavinos | Enhancement of biodegradability of industrial wastewaters by chemical oxidation pre-treatment[END_REF] and [START_REF] Annabi | Degradation of enoxacin antibiotic by the electro-Fenton process: optimization, biodegradability improvement and degradation mechanism[END_REF] explained this phenomenon: bigger molecules are less likely to be degraded by a biological treatment as their size prevents them from permeating cell walls.

The second group is formed of molecules with a high biodegradability rate but also containing a small amount of non-biodegradable molecules which will require a post-chemical treatment to eliminate the resistant pollutants. Here it is cost-effective to start with a biological treatment and then use an AOP. The third group is composed of molecules with a toxic activity. In this case, an AOP pre-treatment could be preferred. Finally, the fourth group is made of solutions that contain molecules that are biologically transformed into dead-end intermediates compounds that cannot be further biologically degraded and can lead to strong inhibition of the biological activity. Thus, an additional AOP treatment should be performed in order to degrade these intermediate molecules (Scott and Ollis 1995).

This classification was completed by [START_REF] Comninellis | Advanced oxidation processes for water treatment: advances and trends for R&D[END_REF] who added the notion of TOC to classify the pollutants. They defined 4 groups, one with biodegradable compounds which correspond to the second group of Scott and Ollis and three groups with non-biodegradable compounds: (i) wastewaters with a high TOC (> 100 mg L -1 ) requiring a pre-treatment with an AOP before a biological treatment; (ii) wastewaters with a low TOC (< 100 mg L -1 ) but with a high toxicity; in this case a suitable AOP is necessary to degrade the toxic compounds but a biological treatment is not recommended as wastewaters have little metabolic value for the bacteria; (iii) wastewaters containing non-biodegradable molecules at low TOC and toxicity but with a characteristic which prevents their release, such as colored wastewaters. Here only an AOP is advised.

As it was seen with both classifications, most of the time an AOP is performed first followed or not by a biological treatment as illustrated by Mantzavinos and Psillakis as most of the already published works showed a rise in biodegradability and/or a decrease in toxicity with an AOP pre-treatment. But sometimes that is not the case [START_REF] Mantzavinos | Enhancement of biodegradability of industrial wastewaters by chemical oxidation pre-treatment[END_REF]Oller et al. 2011) for several reasons: (i) stable intermediates formed may be less degradable than the parent compound; (ii) AOPs are not specific, so the biodegradable molecules can be degraded before persistent ones; (iii) finding the good conditions for the pre-treatment is needed: if too short, the solution will not change enough to be treated by activated sludge and if too long the process will not be cost-effective; (iv) the compounds used during the AOP can be toxic for the biological treatment, for example, O3 or H2O2: if they are not eliminated the second treatment will not be efficient and (v) if a high concentration of halides is used, the AOP process could form toxic organic halogens which increase the toxicity of the solution [START_REF] Fang | Enhanced AOX accumulation and aquatic toxicity during 2,4,6-trichlorophenol degradation in a Co(II)/peroxymonosulfate/Cl-system[END_REF].

In short, even if some guidelines have been found to classify wastewaters, performing each treatment separately and studying the synergetic effect as well as measuring the toxicity and the biodegradability during each treatment are necessary in order to find the best combination with the highest efficiency and the lowest cost.

Effects of EF on effluents characteristics

Enhancement of biodegradability

Table 9 gives an overview of the recent studies in which biodegradability of pharmaceutical solutions is investigated, except for the combination of EF with a biological treatment, which will be more detailed in sub-section 4.3. According to this Table, among the twelve studies, only three studies were performed with both treatments: in two works, an EF treatment is followed by a biological degradation step and in the third one, the treatment is performed in the opposite order. In four cases, an EF treatment followed by a biological treatment was found quite inappropriate as only in one case the solution obtained was very toxic (L) [START_REF] Mousset | Treatment of synthetic soil washing solutions containing phenanthrene and cyclodextrin by electro-oxidation. Influence of anode materials on toxicity removal and biodegradability enhancement[END_REF], in two studies the final solution have a BOD5/COD lower than 0.3 after EF treatment (G,H) [START_REF] Elias | Effect of electro-Fenton application on azo dyes biodegradability[END_REF][START_REF] Fourcade | Electrochemical reduction prior to electro-Fenton oxidation of AZO dyes: impact of the pretreatment on biodegradability[END_REF]) and only in one study the biodegradability level was over the limit but the time of treatment was 24 h which made the process too expensive

Toxicity assessment

With biodegradability, toxicity is the second very important parameter for a biological treatment. A dozen studies reported the toxicity assessment of solutions during EF treatments.

Most of them used the Microtox ® test. This was applied to herbicides such as mesotrione and diuron, to azo dyes such as acid orange 7 or to pharmaceuticals such as sulfamethoxazole, sulfamethazine and sulfanilamide. For all those studies an increase of the toxicity was observed at the beginning of the EF treatment, this was due to the formation of intermediate compounds more toxic than the parent pollutant. This increase in solution toxicity was usually attributed to cyclic/aromatic by-products which are formed after a few minutes of electrolysis. As an example, one of these intermediates can be hydroquinone, obtained both in the degradation of phenol [START_REF] Pimentel | Phenol degradation by advanced electrochemical oxidation process electro-Fenton using a carbon felt cathode[END_REF]) and chlorophenol [START_REF] Yuan | Comparison treatment of various chlorophenols by electro-Fenton method: Relationship between chlorine content and degradation[END_REF]. Then after reaching a plateau in the percentage of inhibition, solution toxicity decreases due to the progressive destruction of aromatic/cyclic intermediates and formation of carboxylic acids (which are much less toxic) and to the progress in the mineralization of the solution. Depending on the starting pollutant, the complete removal of toxicity (0% of inhibition) can be achieved in 2 h as for the sulfamethazine or after 8.33 h for the mesotrione. Indeed, as for the efficiency of the EF process, the decrease of toxicity is also different when using, for example, different anode materials (mainly BDD or Pt) or different operating parameters such as current intensity or initial concentration of pollutant (Oturan et al. 2008c;[START_REF] Dirany | Study of the toxicity of sulfamethoxazole and its degradation products in water by a bioluminescence method during application of the electro-Fenton treatment[END_REF]El-Ghenymy et al. 2014;[START_REF] Murati | Application of the Electro-Fenton process to mesotrione aqueous solutions, kinetics, degradation, pathways, mineralization and evolution of toxicity[END_REF][START_REF] Barhoumi | Pyrite as a sustainable catalyst in electro-Fenton process for improving oxidation of sulfamethazine. Kinetics, mechanism and toxicity assessment[END_REF]Le et al. 2016b). In the study concerning the degradation of the herbicide diuron, the authors performed the toxicity assessment both on Vibrio Fischeri (Microtox ® ) and green algae and found a really different behavior for these two different tests. For the algae, toxicity decreases rapidly whereas for the bacteria the toxicity rises because of the intermediates formed and then decreases. These results highlight that the toxicity observed depends on the organism used in the test. To give an overall information on the toxicity and thus to be more reliable, tests need to be done on several different species or methods (Oturan et al. 2008c). Other toxicity assessments were done by using other species. Toxicity was also measured by the oxidative stress which is a pathophysiological mechanism coming from the imbalance of the production of reactive species and organism defenses that can lead from molecular damage to the cell death. In a study on the degradation of Yellow 9 performed by EF, the activity of catalase was measured in order to evaluate the oxidative stress.

Indeed, the enzyme catalase is involved in protecting the cell from oxidative stress. Thus a high concentration of catalase means that the cell is fighting against this process. It was found that only Yellow 9 induced toxicity as the catalase activity decreased with the time of electrolysis [START_REF] Kourdali | Degradation of direct yellow 9 by electro-Fenton: process study and optimization and, monitoring of treated water toxicity using catalase[END_REF]). Another study looked at toxicogenomics and evaluated in addition the type of stress. Three well-known molecules were studied: bisphenol A, triclosan and ibuprofen;

they were chosen because their degradation pathway is quite reliable. For bisphenol A, a large range of stress response was found whereas for its by-products the stress was concentrated on DNA and membrane. For triclosan, an oxidative stress was evaluated for the parent compound and for the intermediates; an anomalous production of proteins and an alteration of the cell membrane was assessed. Finally, for ibuprofen, the parent compound alters the cell transport function and induces a high membrane stress. The by-products, as far as they are concerned, promote membrane, protein and DNA stress [START_REF] Gou | A quantitative toxicogenomics assay reveals the evolution and nature of toxicity during the transformation of environmental pollutants[END_REF]. This type of study is really interesting to understand better how the toxicity operates and then it is easier to work with and to know when a bio-treatment could be useful.

In the five last years, only two studies were found dealing with both biodegradability and toxicity (H-L) [START_REF] Elias | Effect of electro-Fenton application on azo dyes biodegradability[END_REF][START_REF] Mousset | Treatment of synthetic soil washing solutions containing phenanthrene and cyclodextrin by electro-oxidation. Influence of anode materials on toxicity removal and biodegradability enhancement[END_REF]. In order to perform the combination of both EF and biological treatment more studies are in fact needed.

Combined process applied to pharmaceuticals removal

Anodic oxidation as pre-treatment for a post-biodegradation process

A few studies reported the experiment of a combined process using an AOP and a biological treatment applied to pharmaceutical wastewater. To the best of our knowledge, the first attempt of an hybrid process combining bio-treatment and an electrochemical process was done by [START_REF] Carvalho | Electrochemical degradation applied to the metabolites of Acid Orange 7 anaerobic biotreatment[END_REF]. Biotic degradation products formed during biological treatment of a synthetic solution of dye Acid Orange 7 was carried out by anodic oxidation using BDD anode obtaining a COD removal of 90% and a TOC removal of 30% after 8 h treatment. Then, some works on the combination between biodegradation and AO process were carried out [START_REF] Khoufi | Pilot scale hybrid process for olive mill wastewater treatment and reuse[END_REF]Ganzenko et al. 2014;Olvera-Vargas et al. 2015a;Olvera-Vargas et al. 2016b). During the combined process, AO was used as a pre-treatment step followed by an anaerobic digestion and an ultrafiltration to reach 88% of COD removal, 72.5% of color removal and 90% of total suspended solids removal with a residence time of 4 h in the EF reactor. The use of an aerobic treatment was also investigated [START_REF] Khoufi | Pilot scale hybrid process for olive mill wastewater treatment and reuse[END_REF]). [START_REF] Yahiaoui | Removal of tetracycline hydrochloride from water based on direct anodic oxidation (Pb/PbO2 electrode) coupled to activated sludge culture[END_REF] used an anodic oxidation process as AOP as a pre-treatment step for a biological treatment to degrade the antibiotic tetracycline. The optimal operating conditions were evaluated on a synthetic solution using a central composite design with an electrolytic cell equipped with Pb/PbO2 anode and Pt cathode. In order to keep enough organic carbon to perform a biological treatment, the solution was electrolyzed through the following conditions:

40 °C, 275 mA, and 100 mg L -1 initial concentration until the complete disappearance of tetracycline. Biodegradability was also assessed along the electrolysis and reached a value of 0.41 after 5 h of treatment. At this point, 36.7% of the initial DOC remained in the medium.

The aerobic treatment using activated sludge confirmed the beneficial effect of the biological treatment as the combined process reaching 76% DOC removal.

Coupling between EF and bio-treatment

The combination between EF and biological treatment called bio-electro- Microtox method as well as MCE and EC data in addition to the evaluation of the biodegradability (BOD5/COD) during EF treatment (Fig. 9). The biodegradability ratio of 0.4 was reached only after 1 h EF treatment at 500 mA constant current with 47% TOC removal of a solution of the beta-blocker metoprolol. A mineralization rate of 90% is attained after 6 d of biological treatment (Fig. 10). The shortness of the EF step allowed the reduction of a significant energy consumption (6 times) compared to EF alone. The evolution of solution For furosemide and ranitidine, some economic considerations were done using the French industrial price for electricity: 0.08 € (kWh) -1 . For 1 h electrolysis the energy cost was of 7.66 and 9.33 € (kg TOC) -1 for furosemide and ranitidine, respectively, whereas for 8 h electrolysis to achieve nearly total mineralization the cost attained 43.03 and 46.77€ (kg TOC) -1 for furosemide and ranitidine respectively with EF alone (Olvera-Vargas et al. 2016b). This clearly justifies the use of the combined process. For all the combined processes applied to pharmaceutical, the EF was used as a pre-treatment, except the report of Ganzenko et al. (2017) in which pre-and -post-EF scenarios have been performed, usually due to the toxicity of the compounds and the low biodegradability of the solution. In the latter work Ganzenko et al.

(2017), compared the degradation and mineralization of a pharmaceutical effluent spiked with caffeine and 5-fluorouracil with two different orders of treatment (EF-Bio or Bio-EF) (Table 10). The results demonstrated a quite similar mineralization (more than 90% of COD removal)

for both orders but using the EF as pre-treatment appeared to be cost saving. Indeed, the best condition for EF-Bio was 2 h at 200 mA whereas for Bio-EF it required 4 h at 500 mA (Ganzenko et al. 2017).

The combination of EF treatment with a biological treatment to treat pharmaceuticals was also carried out by Ferrag-Siagh et al. (2013) in which they showed the efficiency of combining electro-Fenton process with a biological treatment for the degradation of the antibiotic tetracycline. They evaluated different parameters to find the best conditions for the EF (see Table 10). The treatment of tetracycline by biological treatment attained a TOC removal of 10% which clearly showed the inadequacy of this process due to the low biodegradability rate (BOD5/COD) = 0.02) of the initial drug solution. To obtain a biodegradable solution, EF pretreatment was applied to reach 46% and 72% TOC removal and, 0.33 and 0.44 BOD5/COD ratio at 2 and 4 h treatment, respectively. Then the activated sludge treatment was performed during 25 d for the two electrolyzed solutions, obtaining an improvement of 23% and 14% in TOC removal for 2 h and 4 h EF pre-treated solutions, respectively (Ferrag-Siagh et al. 2013).

The same group also investigated the degradation of tylosin, sulfamethazine and trimethoprim using the combined process obtaining a biodegradable solution (BOD5/COD ratio ≥ 0.4) after EF pre-treatment step before biological treatment. The coupled process was then successfully applied to the treatment of a real effluent coming from a pharmaceutical industry (Ferrag-Siagh In all the experiments performed on synthetic solutions as well as on real wastewaters; the costeffectiveness and the efficiency of the coupling between EF and bio-treatment was proved.

However, it should be noticed that as mentioned in section 2.3.5, the electrical consumption may be significantly different at pilot and full scale. It is also very important to underline one of the major drawbacks; the solution pH which is adjusted around 3 in EF step should be raised around 7 for the bio-treatment .

Conclusions and perspectives

Wastewaters from hospitals and pharmaceutical industries contain relatively high concentrations of pollutants as they have not already been mixed with other effluents. As highlighted in different publications, the EF process is faster, more efficient and more costeffective on concentrated solutions of pharmaceuticals. Thus the best approach to treat these wastewater with EF seems to be the treatment of the wastewaters directly at the exit of hospitals or pharmaceutical production plants.

A major advantage of EF compared to the non-AOP treatments is that this process is able to destroy quite every type of pollutants thanks to the non-selectivity of the hydroxyl radicals and to reach almost complete mineralization of treated solutions.

One major problem concerning the EF is its relatively high treatment cost when using longer treatment times, required to reach high mineralization degree to destroy resistant intermediates formed during the process. Research was done to reduce this treatment cost or enhance the efficiency of the EF treatment. These studies were based on the improvements of geometrical parameters or flowing characteristics in order to improve mass transfer and to limit electrical resistance that are in general the limiting steps, on avoiding the addition of external O2 or on using micro-fluidic. Several reactor configurations were proposed and their advantages and drawbacks as well as the best configurations depend on the objectives, on the nature of the wastewater (organic pollutants, presence of solid particles, etc.). In order to reach the industrial scale with the EF process, more research is needed to assess the feasibility: (i) to operate in continuous mode, (ii) to treat real wastewaters, (iii) to decrease the cost associated to catalyst by testing recoverable solid catalyst catalysts at pilot scale. In terms of engineering, a better characterization of the hydrodynamic of the reactors at pilot scale is required to develop models coupling mass transfer phenomena, hydrodynamic behavior, electrochemical reactions, etc. in order to predict the efficiency of the reactors and helping to resolve a number of design aspects required for a mid-term pre-industrialization. In addition, more studies at pilot scale are required to get more information about the energetical consumption of each configuration. Only few studies gathered part of these parameters; this is why it is an important research focus.

EF process has demonstrated its efficiency but the treatment cost is still relatively high (at least 0.39 kWh (g TOC) -1 ) and a good way for saving energy in order to industrialize it, is its combination with a biological treatment. In the case of the treatment of pharmaceutical wastewaters, EF process as a pre-treatment followed by a biological treatment proved its accuracy. However, all the experiments were performed at laboratory scale and under batch conditions. Therefore, for a pre-industrialization of the combined treatment, experiments must be conducted at pilot scale and the feasibility of conducting continuous or semi-continuous reactors must be assessed. The extent of conditioning of the active biomass used for the biological treatment to the parent compound, to the reaction intermediates, or to the compound mixtures also needs to be addressed. In addition, more work is needed to determine the monitoring parameters to control the reactor's operation.

Thus to answer some of the previous points, a lab-scale pilot was built for this thesis and it was tested to treat different solutions of pharmaceuticals on its one or combined with a biological treatment. This will be developed in the next chapters.

Introduction

Nowadays, the elimination of pharmaceuticals from wastewaters has become a hot topic as these products are widely consumed all over the world. As an example, in western Europe around 300 mg of active substances are consumed every day per inhabitant (Khetan and Collins 2007;Bouissou-Schurtz et al. 2014;Luo et al. 2014;Vergili et al. 2019) and are hardly removed by wastewater treatment plants (WWTP). Indeed, between 60 and 90% of them remain in the environment. A study monitored around 50 pharmaceuticals after their treatment in WWTP in Europe and the United States; the average % of elimination was less than 50% [START_REF] Jones | Human pharmaceuticals in wastewater treatment processes[END_REF]Margot et al. 2015). Once used, pharmaceuticals are eliminated in whole or in part by the human or animal bodies and then enter into the wastewaters of hospitals or municipalities.

Wastewaters from pharmaceutical industries constitute another main source of drugs in natural water (Lindsey et al. 2001;[START_REF] Kabdasli | Assessment of pollution profile ans wastewater control alternatives of a pharmaceutical industry[END_REF]. As these molecules are highly active, they can be harmful toward the environment [START_REF] Boxall | The environmental side effects of medication[END_REF][START_REF] Küster | Pharmaceuticals in the environment : scientific evidence of risks and its regulation[END_REF]. Therefore, their elimination from wastewater using an efficient treatment process constitutes an important issue.

The diuretic hydrochlorothiazide (HCT) used for anti-hypertensive medications, is of particular interest as it was recently found both in hospital effluent (Mendoza et al. 2015) and in surface water (Bouissou-Schurtz et al. 2014). Indeed, this molecule was shown to have an average removal rate of 30% from typical WWTP (Margot et al. 2015). In addition, this molecule was reported to be hazardous toward the environment (Fernández et al. 2010) and therefore needs to be efficiently destroyed before being released into natural water body.

Nadolol (Nad) belongs to the b-blockers, it was chosen as it was detected in wastewaters since it is not completely removed by wastewater treatment plants (Gabet-Giraud et al. 2010).

Additionally, it is harmful toward the environment (Khetan and Collins 2007).

Tramadol (TMD) is a centrally acting analgesic used in case of moderate to severe pain [START_REF] Ghalwa | Electrochemical degradation of tramadol hydrochloride: novel use of potentiometric carbon paste electrodes as a tracer[END_REF][START_REF] Antonopoulou | Photocatalytic degradation and mineralization of tramadol pharmaceutical in aqueous TiO2 suspensions: evaluation of kinetics, mechanisms and ecotoxicity[END_REF]. This medicine is widely used, as an example in 2012 around 24 tons of TMD was prescribed in Germany (Lütke [START_REF] Eversloh | Electrochemical oxidation of tramadol in low-salinity reverse osmosis concentrates using boron-doped diamond anodes[END_REF]. It was not efficiently removed by WWTPs as it was detected in effluents from households and hospitals even after treatment and in surface waters [START_REF] Hummel | Simultaneous determination of psychoactive drugs and their metabolites in aqueous matrices by liquid chromatography mass spectrometry[END_REF]; Rúa-Gómez and Püttmann 2012; Rúa-Gómez and Püttmann 2013).

Ofloxacin (OFL) belongs to the antibiotic family which is particularly toxic to microorganisms. This molecule was found both in hospital effluents and surface water (Table 2) (Verlicchi et al. 2010). The presence of antibiotics in the environment leads to chronic toxicity and to the development of antibiotic resistant bacteria, this is why the substances have to be removed from wastewaters [START_REF] Pulicharla | Toxicity of chlortetracycline and its metal complexes to model microorganisms in wastewater sludge[END_REF][START_REF] Changotra | Natural soil mediated photo Fenton-like processes in treatment of pharmaceuticals : batch and continuous approach[END_REF][START_REF] Yang | Heterogeneous fenton-like degradation of o floxacin over sludge derived carbon as catalysts : mechanism and performance[END_REF]).

To eliminate persistent pharmaceuticals, the use of highly strong oxidants is thus compulsory.

Hydroxyl radicals ( • OH) seem to be a promising tool as being the second strongest oxidant after fluorine. They can degrade nearly any type of contaminants thanks to their very high oxidation potential (E°( • OH/H2O) = 2.8 V vs SHE in acidic medium) (Sirés and Brillas 2012). The advanced oxidation processes (AOPs) (Oturan and Aaron 2014) and particularly electrochemical AOPs (EAOPs) (Brillas et al. 2009;Panizza and Cerisola 2009;Sirés et al. 2014;[START_REF] Nidheesh | An overview on the removal of synthetic dyes from water by electrochemical advanced oxidation processes[END_REF] are interesting processes for removal of organic pollutants from water as they have shown their ability in high oxidation rate and quasi-complete mineralization yield [START_REF] Bellakhal | Olive Oil Mill Wastewater Treatment by the electro-Fenton Process[END_REF]Faouzi et al. 2006;Dirany et al. 2012;[START_REF] Özcan | Complete removal of the insecticide azinphos-methyl from water by the electro-Fenton method -A kinetic and mechanistic study[END_REF]Mousset et al. 2016). The electro-Fenton (EF) is one of most popular EAOPs being an environmentally friendly and effective process to destroy organic pollutants (Oturan et al. 2000;[START_REF] Ganiyu | Electrochemical mineralization of sulfamethoxazole over wide pH range using FeII FeIII LDH modified carbon felt cathode: degradation pathway, toxicity and reusability of the modified cathode[END_REF][START_REF] Monteil | Efficient removal of diuretic hydrochlorothiazide from water by electro-Fenton process using BDD anode: a kinetic and degradation pathway study[END_REF]. That is why this process was chosen to treat the pharmaceutical solutions.

This process is based on the generation of The process is electrocatalytic since the catalyst (Fe 2+ ) is continuously electro-generated at the cathode. Comparing with classical Fenton's process, the EF process presents several advantages (Nidheesh and Gandhimathi 2012; Oturan and Aaron 2014) such as i) the in situ production of H2O2 avoiding the reagent cost and risks related to its transport and storage, ii) avoiding the precipitation of iron ion (catalyst) under (Fe(OH)3 form because of its very low concentration and thus the formation of process sludge, iii) avoiding or decreasing significantly the rate of wasting reactions consuming • OH because of very low concentration of Fenton's reagent in the system.

When using an appropriate anode such as BDD, the process becomes even more efficient (Oturan et al. 2012;Sopaj et al. 2016) due to the formation of heterogeneous hydroxyl radicals M( • OH) on the anode surface from electrochemical oxidation of water (Eq. 22). The M( • OH)

thus formed are physically adsorbed on anode (M) in the case of non-active anodes like BDD, and are available for oxidation of organics (Panizza and Cerisola 2005;Rodrigo et al. 2010;[START_REF] Nidheesh | Environmental applications of borondoped diamond electrodes: 1. Applications in water and wastewater treatment[END_REF].

M + H2O ® M( • OH) + H + + 2 e - (22)
Therefore, this study presents, for the first time, a detailed investigation of oxidative degradation (including kinetics data) and mineralization of the four pharmaceuticals from different families by the EF process with a BDD anode. The comparison with a Pt anode was performed for the HCT. The electrical energy consumed for almost complete mineralization was also considered. Preliminary experiments were carried out to find optimal operating conditions. The oxidation reaction intermediates and carboxylic acids as well as inorganic ions were identified using GC-MS, LC-MS, HPLC and ion chromatography analyses, and their evolution during electrolysis was followed. Finally, based on this information, a plausible mineralization reaction pathway was proposed for three pharmaceuticals (HCT, TMD, OFL).

Material and methods

Chemicals

All chemicals used in this study were analytical grade and used as received without further purification. The pharmaceuticals HCT, TMD, Nad and OFL were supplied by Sigma Aldrich. Iron (II) sulfate heptahydrate (catalyst source, 99%) was purchased from Acros Organics. Sodium sulfate (Na2SO4) (supporting electrolyte), potassium sulfate (K2SO4), sodium perchlorate (NaClO4), potassium phosphate monobasic, iron(II) perchlorate were supplied by Sigma-Aldrich, Merck, and Acros. Short-chain carboxylic acids (oxalic acid, oxamic acid, acetic acid and maleic acid were obtained from Acros, Fluka and Alfa Aesar. The pollutant solutions were prepared with ultrapure water obtained from a Millipore Milli-Q Simplicity 185 system with resistivity >18 MΩ cm at 25 °C. The pH of solutions was adjusted using analytical grade sulfuric acid (Acros). Organic solvents and other chemicals used were HPLC or analytic grade from Sigma-Aldrich, Fluka and Merck.

Electrochemical equipment

All the experiments were performed in an undivided cylindrical glass cell of 7 cm diameter and with a capacity of 250 mL at room temperature. The 230 mL solution was vigorously stirred by a magnetic PTFE bar to have a good mass transport between the electrodes and compressed air was continuously bubbled into the cell through a silica frit. The air bubbling was started 5 min prior to electrolysis to ensure O2 saturation in the solution. The electrochemical cell was equipped with a 3D carbon felt (18.0 cm × 5.0 cm × 0.5 cm, from Mersen, France), placed on the inner wall of the cell covering the total internal perimeter as the cathode and a 25-cm 2 BDD film on niobium support as the anode (from CONDIAS GmbH, Germany). A platinum anode with a similar area is used for a few experiments. CyberScan pH 1500 pH meter from Eutech

Instruments was used to adjust the solution pH to 3.0 (well-known as the optimal value for electro-Fenton process [START_REF] Kremer | The Fenton Reaction. Dependence of the Rate on pH[END_REF]Brillas et al. 2009) using a 1 M H2SO4 or NaOH solution.

A constant current, ranging between 50 and 1000 mA, was supplied with a Hameg HM8040 triple power DC supply.

Analytical methods and procedures

The concentration decay of pharmaceutical degradation experuments was monitored by HPLC using a Merck Lachrom Liquid chromatography equipped with a reverse phase column Purospher RP-18, 5 µm, 25 cm × 4.6 mm (i.d.) column and coupled with a diode array detector (DAD). Samples of 20 µL were injected in the column at 40°C. For the HCT, the detector was set at λ = 270 nm and the mobile phase composed of methanol / water 25:75 (v/v) was eluted at a flow rate of 0.5 mL min -1 . For TMD, the mobile phase was a mixture of methanol / phosphate buffer 40:60 (v/v) at a flowrate of 0.6 mL min -1 . The wavelength of detection was 228 nm. The phosphate buffer was composed of 3.5 g of potassium phosphate monobasic and 17.8 g of dipotassium hydrogen phosphate in 1 L of ultra-pure water and the pH was adjusted to 6.0 using orthophosphoric acid. For the nadolol, the mobile phase was composed of an acetate buffer and methanol (63/37 v/v) at 0.4 mL min -1 . The buffer was made of 0.1 M of sodium acetate and the pH adjusted to 6 by acetic acid. The wavelength of detection was 277 nm. Finally, the OFL was eluted by a mobile phase made of acetonitrile and acetate buffer 20/80 (v/v) at 0.5 mL min -1 and detected at 292 nm. The buffer was composed of 0.05 mol L -1 of citric acid and of 1 mol L -1 of ammonium acetate.

The short-chain carboxylic acids were identified and quantified by the same HPLC equipped with a Biorad column (9 µm, 250 mm × 4.6 mm) and the DAD detector was set at 220 nm. The mobile phase was a solution of H2SO4 at a concentration of 4 mM with a flow rate of 0.6 mL min -1 . Carboxylic acids were identified by comparison of their retention time with that of standard solutions and their quantification was achieved thanks to calibration curves.

Nitrate, nitrite, chloride, sulfate and fluoride ions released to the treated solutions were analyzed using a Dionex ICS-1000 Basic Ion Chromatography (IC) system equipped with a DS56 conductivity detector containing a cell heated at 35 °C using an Ion Pac AS4A-SC, 25 cm× 4 mm (i.d.), anion-exchange column, linked to an IonPac AG4A-SC, 5 cm × 4 mm (i.d.), column guard. The eluent was a mixture of 1.7 mM sodium bicarbonate (NaHCO3) and 1.8 mM sodium carbonate (Na2CO3) at a flow rate of 1 mL min -1 . An ASRS-ULTRA II self-regenerating suppressor was used in order to improve the sensitivity of the detector. For ammonium ion detection, the same IC equipped with an IonPac CS12A, 25 cm × 4 mm (i.d.), cation-exchange column, linked to an IonPac CG12A, 5 cm × 4 mm (i.d.), column guard was used. Again, an electrochemical suppressor (CRS-ULTRA II) was employed, the. The eluent was a solution of 9.0 mM sulfuric acid at a flow rate of 1.0 mL min -1 . The volume of injections was 25 µL for the anions and the cations.

The mineralization degree of treated solutions was measured thanks to the total organic carbon (TOC) removal of the samples withdrawn from the treated solution at different electrolysis times. The TOC measurement was monitored by a Shimadzu VCSH TOC analyzer according to the 680 °C catalytic combustion method using platinum as catalyst. The injection volume was 50 µL. Reproducible TOC values with ± 2% accuracy were found using the non-purgeable organic carbon method. The TOC abatement was calculated by dividing the TOC at time t by the TOC at the initial time.

From these data, the mineralization current efficiency (MCE) value for each trial was then estimated from Eqs. 23.

!"#(%) =

) * + ,-./(0)1-./(2)3 4.67 * 90 : ; < 2 * 100

(23)
where TOC(0) -TOC(t) is the TOC decay (mg carbon L -1 ), m is the number of carbon atoms of the molecule studied, t is the time of electrolysis (h), I is the applied current (A), V is the volume of the solution (L), 4.32 Í 10 7 is a conversion factor to homogenize units and n is the number of electrons used to mineralize one molecule of pollutant into CO2, H2O and inorganic ions (NH4 + , SO4 2-, F -and Cl -) following the electrochemical combustion reaction as shown in Eqs.

24 a, b, c and d for HCT, TMD, Nad and OFL, respectively. The ion analysis showed that ammonium was formed dominantly (compared to nitrate), thus in the equations, were written considering ammonium as released nitrogen ion.

C7H8ClN3O4S2 + 18H2O ® 7CO2 + 32H + + 2SO4 2-+ 3NH4 + + Cl -+ 30e - (24a) C16H25NO2 + 30H2O ® 16CO2 + 81H + + NH4 + + 82e - (24b) C17H27NO4 + 30H2O → 17CO2 + NH4 + + 83H + + 84e - (24c) C18H20FN3O4 + 32H2O → 18CO2 + F -+ 3NH4 + + 72H + + 74e - (24d)
The energy consumed in kWh per g TOC removed was evaluated from Eq. 25.

#" ? Then the mixture was cooled to room temperature and 3 µL were injected into the GC-MS.

LC-MS analysis were performed for the Nad and the OFL using an Hypersil Gold column (100 mm × 2.1 mm) 1.9 µ, at 40 °C and with a liquid flowrate of 0.2 mL min -1 . After drying the sample, the OFL was diluted in 1 mL of methanol / water 50/50 (v/v) and eluted by 0.1% formic acid / methanol 80/20 (v/v). Concerning the MS conditions, the ESI + (electrospray ionization)

mode was used at 90 V. For the Nad, the evaporated solution was diluted in 1 mL of water and eluted by the same conditions. The ESI + mode was also used at 75 V.

When error bars are given on the graphs, they are calculated based on the standard deviation of the different experiments.

Results and Discussion

Influence of the operating parameters on the degradation kinetics

There are several parameters affecting the performance of the EF process during oxidation of organic pollutants. Among these parameters some are now well-known (Brillas et al. 2009;

Oturan and Oturan 2018): solution pH (= 3.0), nature of the catalyst (Fe 2+ ), nature of electrolyte (Na2SO4). The temperature has a very little effect. However, for a given anode/cathode couple, the effect of two parameters: concentration of the catalyst (Fe 2+ ) and the value of applied current, should be studied to have optimal operating values for an effective treatment. The nature of the anode is also very important to have an efficient degradation.

Effect of the concentration of Fe 2+

The concentration of ferrous irons, used as catalyst, is one of the key parameters for an effective degradation of organic pollutants because it is directly linked to the formation of • OH through Fenton's reaction (Eq. 19). To clarify the effect of catalyst concentration, experiments were done using 0.1 mM concentration of pharmaceuticals at 100 mA constant current. As shown in Figs. 12-15, rising the concentration of Fe + from 0.05 mM to 0.1 mM (from 0.05 mM to 0.2 mM for the case of TMD, Fig. 13) increases the amount of • OH formed and thus allows a quicker degradation of pharmaceuticals. However the concentrations above 0.1 mM Fe 2+ for HCT, OFL and Nad and 0.2 mM for TMD lead to a decrease in oxidation kinetics that can be explained by the increase of the rate of the wasting reaction given in Eq. 26. Indeed, at high concentration of Fe 2+ , this reaction can compete with oxidation of the pollutants for • OH leading to a decrease in degradation kinetics (Panizza and Cerisola 2001;Dominguez et al. 2018b).

•

OH + Fe 2+ ® Fe 3+ + OH - (26)
To evaluate quantitatively the effect of iron concentration on degradation rate of the pollutants, apparent rate constants of oxidation reaction were calculated assuming pseudo-first order reaction kinetics. Indeed, the experimental conditions being kept constant during the experiments, the formation rate of • OH/M( • OH) can be considered constant. On the other hand, the steady-state assumption can be applied to the concentration of hydroxyl radicals as they are very reactive species and cannot be accumulate in the medium (Sirés et al. 2007a;Brillas et al. 2009). Indeed the exponential decrease of the concentration in each case confirms this hypothesis. Accordingly, the value of apparent rate constant for oxidation of pharmaceuticals can be calculated from the slop of the straight lines obtained by plotting Ln (C/C0) = f(t) as

shown on the inset panels of the Figs. 12-15.

Effect of the applied current on the degradation kinetics

The applied current is one of the most important parameters in EAOPs in general and in EF in particular for the degradation of organic pollutants. completely oxidatively degraded in 15, 10, 12 and 10 min, respectively, at 500 mA constant current electrolysis. The apparent rate constants were also calculated for each current and pollutant. They were of 0.14, 0.19, 0.29 and 0.39 min -1 for 50, 100, 300 and 500 mA respectively for the HCT (Fig. 16) and of 0.70, 0.71 and 0.26 min -1 at 500 mA for TMD (Fig. 17), OFL (Fig. 18) and Nad (Fig. 19), respectively. For all the trend curves a good correlation coefficients were found proving the good accuracy of the pseudo first order reaction model.

Nevertheless, the use of too high currents can promote side reactions such as H2 evolution on the cathode or O2 evolution on the anode (Zhang et al. 2007;Oturan and Oturan 2018). This behavior is visible as the rate constants do not increase proportionally to the applied current because of this enhancement of rate of side reactions with current, causing thus a loss in process efficiency. The increase of the current also promotes the formation of gas bubbles which decreases the conductivity of the solution and thus the process efficiency. Moreover, using a higher current increases electrical energy consumption and consequently process cost.

Therefore, a good compromise has to be found between a suitable efficiency and a rational operating cost. To consider the effect of the ions present in a real wastewater on the degradation of the HCT, a synthetic solution was prepared by adding different ions. The concentration of ions used were at the level of a real wastewater coming from an industrial production sites of antibiotics [START_REF] Marcelino | Evaluation of aerobic and anaerobic biodegradability and toxicity assessment of real pharmaceutical wastewater from industrial production of antibiotics[END_REF]. Results are given in Fig. 20. There is a small decrease in the kinetic of degradation probably due to scavenging reactions with the added ions. 244 mg L -1 Na2CO3, 48 mg L -1 of Cl -, 4 mg L -1 of PO4 3-, 0.9 mg L -1 NO2 -and 1.5 mg L -1 of NO3 -.

Influence of the anode material on the degradation kinetics

The degradation of the Nad was also studied using a platinum anode. The Nad decay is represented in Fig. 21 for different current values. As for the BDD anode, the degradation efficiency is increasing with current until 500 mA. For 1000 mA, the kinetic curve is very similar to the one of 500 mA due to side and wasting reactions consuming the electrical energy and the hydroxyl radicals. The complete oxidation of Nad needs 40 min electrolysis at 300, 500

and 1000 mA which is much higher than for the degradation with the BDD being of 10, 12 and 20 min for 1000, 500 and 300 mA, respectively. This is due to the nature of anode material:

BDD promotes the formation of high amounts of adsorbed hydroxyl radicals (due to its high O2 evolution overpotential). Moreover, the BDD( • OH) is weakly sorbed on BDD surface enabling it to react easily with organics (Marselli et al. 2003) whereas the interaction is much stronger (chemisorption) in the case of Pt( • OH). constants obtained for the mineralization process. As for the BDD anode, the apparent kinetics constants for Pt anode are not increasing proportionally to the current for the same reasons. The degradation constants for both Pt and BDD are quite similar in early electrolysis times (t < 8 min). Indeed, BDD( • OH) degrade and mineralize simultaneously whereas Pt( • OH) mostly degrade which result in a very similar degradation. This behavior was found in other already published reports for high currents (Guinea et al. 2008;El-Ghenymy et al. 2014;Sopaj et al. 2016).

Determination of the absolute rate constant for oxidation of four pharmaceuticals by hydroxyl radicals

The absolute (second order) rate constant of the oxidation of the pharmaceuticals under study by hydroxyl radicals was determined using the competition kinetics method following the Eq. 27 (Oturan et al. 2000). Benzoic acid (BA) was selected as the standard competitor since the rate constant for its oxidation by . OH (kBA = 4.3 x 10 9 M -1 s -1 ) is well-known [START_REF] Buxton | Critical review of rate constants for reactions of hydrated electrons, hydrogen atoms and hydroxyl radicals (.OH/.O-) in aqueous solution[END_REF]).

ln ?

[K]0

[K] D = @P @BA Î ln ?

[PQ]0

[PQ] D 27) with kP and kBA the absolute rate constant of the pollutant and BA respectively, [R] 0, the initial concentration in pollutant and [R] the concentration of pollutant at t time. The concentrations are given in the same unity, mM usually.

To determine kP, the electrolysis of equimolar concentration of the pharmaceutical under study and BA is performed under electro-Fenton conditions at 50 mA applied current. The decay in the concentration of both compounds was carried out over short electrolysis time and at lower current to avoid the interference of intermediate products. The determination of kP was then done thanks to the slope of the regression line of Fig. 22 which was built following Eq. 27. The absolute rate constants were found to be: (4.37 ± 0.04) Î 10 9 , (5.59 ± 0.03) Í 10 9 , (5.63 ± 0.03)

Í 10 9 and (5.55 ± 0.05) Í 10 9 M -1 s -1 for HCT, TMD, Nad and OFL, respectively. These values are very close and highlight the similar reactivity of the hydroxyl radicals towards four pharmaceutical selected.

Mineralization of the pharmaceuticals

The mineralization of different pharmaceutical solutions was assessed by monitoring TOC measurements during treatment. The MCE and EC were then evaluated based on TOC removal data for different applied currents in order to find the better efficiency related to a lower cost.

The nature of the anode was also considered by comparing the use of a BDD and Pt electrode.

The measurements were carried out for currents ranging from 100 to 1000 mA.

Effect of the applied current on the mineralization of pharmaceutical solutions

As can be seen in Fig. 23, the TOC removal is faster at early treatment times and for high currents. Indeed the mineralization of 0.1 mM pharmaceuticals solutions lasts for 6 h for 300, 500 and 1000 mA, reaching TOC removal rates of 85%, 95%, 94% and 94% at 100 mA under same operating conditions, for HCT, TMD, Nad and OFL, respectively. Complete mineralization is attained at 500 and 1000 mA for all cases. However the increase in mineralization degree is not proportional to the applied current, particularly at high currents, as can be seen with the distance between the curves. The exponential decrease of TOC is due to the relatively easy degradation of the pharmaceuticals and their aromatic oxidation intermediates at early treatment times. Inversely, the mineralization efficiency becomes significantly slower at longer treatment times due to the formation of the aliphatic compounds such as carboxylic acids which are more resistant to hydroxyl radicals (Brillas et al. 2009;

Oturan and Aaron 2014) and the decrease in the amount of organic matter in the solution (Brillas et al. 2009;[START_REF] Yang | Heterogeneous fenton-like degradation of o floxacin over sludge derived carbon as catalysts : mechanism and performance[END_REF]). This phenomenon is also illustrated by the MCE values representing the part of the current which mineralizes the pollutant. The MCE values are decreasing highlighting less mineralization efficiency for longer treatment times. The decrease in mineralization efficiency with higher currents is mainly due to the raise in the rate of side reactions, decreasing the oxidative ability of hydroxyl radicals towards organic contaminants.

Among those reactions, the most occurring ones are the oxidation of BDD( .• OH) (Eq. 28) and the wasting of • OH by H2O2 (Eq. 29) (Brillas et al. 2009;Oturan and Aaron 2014).

2 M( • OH) ® 2 M + O2 + 2 H + + 2 e - ( 28
) H2O2 + • OH ® HO2 + H2O (29)
The maximum MCE values are obtained for 100 mA (as expected) for 1 h of treatment, reaching 10, 12, 15 and 17% for HCT, TMD, Nad and OFL, respectively. The mineralization rates are between 5 and 10% for 300 mA and below 5% for 500 and 1000 mA, for all the four pharmaceuticals showing an important loss of energy. This low efficiency can be increased by using higher concentrations, for example. This was demonstrated by Brillas' group who degraded different concentrations of propranolol and could reach nearly 100% of MCE with 616 mg L -1 (Isarain-Chávez et al. 2010b) at the beginning of the treatment.

The evolution of EC during the treatment of 0.1 mM of pollutants by EF is presented in Fig. 24. This parameter is also growing with the current and the treatment time, as expected, since the same current value is applied along electrolysis with lower MCE over longer electrolysis times. The results show that EC maximum values are around 16, 12, 12 and 7 kWh (g TOC) -1 for HCT, TMD, OFL and Nad, respectively, after 6 h electrolysis at 1000 mA. Considering the values of MCE and mineralization degree, the more efficient currents seem to be 300 mA and 500 mA for all the pharmaceuticals for a cost effective process. The mean EC values for those currents is around 2 kWh (g TOC) -1 .

Influence of the anode material on the mineralization efficiency

The TOC decay was measured for both BDD and Pt anode under different currents ranging from 100 mA to 1000 mA and results are depicted on the Fig. 25. For the BDD anode, the TOC decay is increasing with the current intensity, reaching a total mineralization of the solution in 6 hours for both 500 and 1000 mA. The very similar trend between these two curves is due to waste reactions degrading BDD( • OH) or H2O2 for example or to the generation of weaker oxidants such as S2O8 2-or O3 (Eq. 33 and 34) occurring at high currents (Panizza and Cerisola 2005;Guinea et al. 2008;Brillas et al. 2009;Oturan and Aaron 2014). Na2SO4 with and without additional ions using 0.1 mM Fe 2+ as catalyst. The additional compounds were 244 mg L -1 Na2CO3, 48 mg L -1 of Cl -, 4 mg L -1 of PO4 3-, 0.9 mg L -1 of NO2 -and 1.5 mg L -1 of NO3 -.

2 BDD( • OH) ® 2 BDD + O2 + 2 H + + 2 e - (30)

Formation and Evolution of short chain carboxylic acids and mineral ions

The cleavage of aromatic/cyclic moieties of organic compounds under oxidative action of hydroxyl radicals is known to form aliphatic compounds that are further oxidized to carboxylic acids (Brillas et al. 2009;Dirany et al. 2012). In order to well understand the degradation of the organics, formed short-chain carboxylic acids during mineralization experiments were identified and their evolution was followed by HPLC using an ion-exclusion column. The evolution of the carboxylic acids created during the degradation process is presented in Figs.

28-31. For the HCT, oxamic, oxalic, maleic and acetic acids were detected and followed until 4 h of treatment. Maleic acid was completely degraded in 90 min, likely being transformed to oxalic and acetic acids which are more recalcitrant (Oturan et al. 2008b). Oxamic acid could be formed from the degradation of compounds containing an amino group (Garcia-Segura and Brillas 2011). The higher accumulation of oxalic (0.009 mM) and oxamic acid (0.0045 mM) after 1 h treatment can be explained by their recalcitrant behavior towards • OH/BDD( • OH). For TMD, three carboxylic acids were identified: glyoxylic acid, oxalic acid and fumaric acid. The most important one was glyoxylic acid for which the concentration reached 0.07 mM after 1 h of treatment. The fumaric acid was accumulated in the solution only at trace level (around 10 -5 mM); it was probably oxidized quickly into glyoxylic and/or oxalic acids. After 6 h of electrolysis, all the identified carboxylic acids were not anymore detectable in accordance with the very low value of TOC (section 3.4.) in these conditions. During the Nad degradation, oxalic, oxamic, acetic, glyoxylic and formic acids were formed. Glyoxylic disappeared after 1.5 h of EF treatment, possibly transformed into formic acid (Oturan et al. 2008b). Oxalic and oxamic acids were again found to be the more resistant; they could be detected until 4 h of EF treatment. For OFL, acetic, oxalic and oxamic acids were formed during the oxidation process with a maximum concentration peak at about 1 h. Fumaric acid was detected in trace concentration. Then after 6 h they were all mineralized. During the mineralization of organics, the heteroatoms present in the initial molecules are oxidized to their higher (often to the highest) oxidizing state and released to the solution as inorganic ions. Thus the evolution of inorganic ions formed during the mineralization of the pharmaceutical solutions was also followed by IC and results are depicted in Figs. 32-35. The 3 N atoms of the HCT and OFL molecules (0.1 mM) and the N atom of TMD and Nad were mineralized into NH4 + and NO3 -. The presence of NO2 -was not detected. After 4 or 6 h electrolysis, the largest amount of NH4 + was measured. In order to confirm this behavior, 1 mM solutions of NH4 + and NO3 -were separately electrolyzed at different currents using K2SO4 as supporting electrolyte and 0.1 mM Fe 2+ as catalyst and the results are detailed in the Fig 36.

After 8 h treatment, the concentration of NH4 + was still equal to its initial value (1 mM) whereas NO3 -was totally transformed into NH4 + (Fig. 36). These results confirm the data obtained during the mineralization of HCT and highlight the electro-reduction of NO3 -into NH4 + taking the advantage of the large 3D surface of carbon-felt cathode. Similar results were already reported by Martin De Vidales et al. ( 2016) and [START_REF] Garcia-Segura | Electrocatalytic reduction of nitrate: fundamentals to full-scale water treatment applications[END_REF].

The mineralization of the 2 S atoms of the HCT molecule was done with complete release of SO4 2-ion (0.2 mM) into the solution after 2 h electrolysis confirming the rapid mineralization of entire S atoms. Concerning the Cl atom, the concentration of Cl -reached 0.086 mM after 30 min treatment before starting to decrease gradually due to its oxidation at BDD anode to the chlorine gas Cl2, which then reacts with water to form HClO (under EF condition of pH near 3). Chlorate and perchlorate can also be formed as the EF process is applied but those molecules could not be measured on the IC. This latter compound is an active chlorine specie that can contribute to the oxidation of HCT in the solution [START_REF] Randazzo | Comparative electrochemical treatments of two chlorinated aliphatic hydrocarbons. Time course of the main reaction by-products[END_REF][START_REF] Brito C Do | Understanding active chlorine species production using boron doped diamond films with lower and higher sp3/sp2 ratio[END_REF].

Finally, the F atom of the OFL was completely released into F -ion in 2 h of electro-Fenton treatment.

Proposition mineralization pathway

It was not possible to identify the intermediates for Nad with the LC-MS analysis. Therefore the mineralization pathways were suggested for HCT, TMD and OFL based on GC-MS, LC-MS, ion-exclusion HPLC, ion chromatography analyses and TOC removal data.

Mineralization of HCT by hydroxyl radicals

In order to propose a mineralization pathway, the oxidation intermediates and the end-products were investigated. In addition to the identified short-chain carboxylic acids which constitute the last stage of the mineralization process and mineral ions as end-products, aromatic/cyclic intermediates formed during oxidation of HCT were also analyzed using GC-MS. These compounds were identified thanks to their molecular mass and fragmentations behavior (Annex Table 3 and4). Based on these data, a plausible mineralization pathway was proposed (Fig. 37).

HCT (A) is first transformed into chlorothiazide which is possibly obtained by the double attack of hydroxyl radicals on the hydrogen linked to the nitrogen first and then on the hydrogen linked to the carbon connected to the nitrogen, as already reported by [START_REF] Borowska | Oxidation of cetirizine, fexofenadine and hydrochlorothiazide during ozonation: Kinetics and formation of transformation products[END_REF]. This chlorothiazide is then whether hydrolyzed in acidic medium to give the product B or a substitution of -NH2 group by -OH occurs to give product C. The intermediate D can be formed by simultaneous ipso-substitution of Cl (following an ipso attack of • OH on -Cl position [START_REF] Mousset | An unprecedented route of OH radical reactivity evidenced by an electrocatalytical process: Ipso-substitution with perhalogenocarbon compounds[END_REF])) and abstraction of H atom from -NH group linked to sulfonic acid group.

The intermediates B and C were also reported by [START_REF] Borowska | Oxidation of cetirizine, fexofenadine and hydrochlorothiazide during ozonation: Kinetics and formation of transformation products[END_REF]. The major pathway may be the formation of compound C as this peak is higher than that of B and D. The

+ H2O + NH4 + A -Cl - -SO 4 2- -SO 4 2- -SO 4 2- •OH -Cl - -Cl - -NH 4 + -NH 4 + -NH 4 + -NH 4 + B C D E F G H -NH 4 + J K L M N O P

3.5.2.. Mineralization pathway for oxidation of TMD by hydroxyl radicals

Gathering the results of sub-section 3.4.2. and the GC-MS results, it was possible to propose a degradation pathway for EF treatment of TMD (Fig. 38). Intermediates B to G were identified thanks to the fragmentation analysis of GC-MS spectrum, intermediates H, I and J were identified by ion-exclusion HPLC analysis. Two main routes can be established following paths Finally oxidative ring opening reactions lead to the formation of carboxylic acids H, I and J (Fig. 38).

Proposed mineralization pathway for oxidation of OFL by hydroxyl radicals

During the EF treatment of the OFL, four main routes of degradation could be identified thanks to HPLC and LC-MS analysis (Fig. 39). The structure modification happens at four different et al. 2013;[START_REF] Pi | Oxidation of ofloxacin by Oxone/Co2+: identification of reaction products and pathways[END_REF]. Finally the whole piperazinyl can be replaced by an hydroxyl group as in molecules B and F. Concerning the fluorine atom, it can be released in the solution as F - by an ipso attack of the hydroxyl radicals [START_REF] Michael | Light-induced catalytic transformation of ofloxacin by solar Fenton in various water matrices at a pilot plant: mineralization and characterization of major intermediate products[END_REF][START_REF] Mousset | An unprecedented route of OH radical reactivity evidenced by an electrocatalytical process: Ipso-substitution with perhalogenocarbon compounds[END_REF]. The quinolone can be transformed, by the loss of C3O, in a carboxylic group as in molecule E [START_REF] Michael | Light-induced catalytic transformation of ofloxacin by solar Fenton in various water matrices at a pilot plant: mineralization and characterization of major intermediate products[END_REF] or in an enone group as in molecule A and G [START_REF] Vasquez | Biodegradation potential of ofloxacin and its resulting transformation products during photolytic and photocatalytic treatment[END_REF][START_REF] Pi | Oxidation of ofloxacin by Oxone/Co2+: identification of reaction products and pathways[END_REF]Yahya et al. 2014). Finally, the cycle containing the oxygen and the azote can loss C2 (molecule C) or C3 (molecule D, G and I), this was also found for the ozonation treatment of the OFL [START_REF] Tay | Ozonation of ofloxacin in water: by-products, degradation pathway and ecotoxicity assessment[END_REF].

At the end of the degradation after oxidative ring opening reactions, three carboxylic acids could be identified (J, K, L).

CO 2 + H 2 O -NH 4 + •OH •OH J K L O HO O OH H 2 N O O OH O HO O OH -NH 4 + -NH 4 + -NH 4 + -F - -F - -F - A B C D E F G H I

Conclusions

The results and discussion done above show that the pharmaceutical solutions under study were efficiently treated reaching high mineralization rates by EF process under the best operating conditions: catalyst (Fe 2+ ) concentration of 0.1 mM except for TMD (0.2 mM), and currents between 300 and 500 mA in 0.05 M Na2SO4 as supporting electrolyte. The absolute rate constants for oxidation of HCT, TMD, Nad and OFL by hydroxyl radicals were also determined using the competition kinetics method and found to be (4.37 ± 0.04) Î 10 9 , (5.59 ± 0.03) Í 10 9 , (5.63 ± 0.03) Í 10 9 and (5.55 ± 0.05) Í 10 9 M -1 s -1 respectively. The complete mineralization of 0.1 mM of all the pollutant solutions was reached at 6 h treatment. Finally, this study ends by proposing a plausible mineralization pathway of the three pharmaceuticals (HCT, TMD and OFL) by hydroxyl radical generated in EF process. These results illustrate the non-selectivity of the hydroxyl radicals, being able to eliminate nearly any organic compounds.

The EF treatment is steel at scale up stage. Some pilots were built and successfully tested which suggest the possible use of this process at pre-industrial scale. This point will be discussed in next chapter. The idea would be to implement this system at the exit of industrial companies as the treatment is more efficient when the pollutant concentration is higher.

Introduction

Electrochemical advanced oxidation processes (EAOPs) such as anodic oxidation (AO) and electro-Fenton (EF) are promising technologies to eliminate persistent and harmful organic pollutants from the environment [START_REF] Iniesta | Electrochemical oxidation of phenol at borondoped diamond electrode[END_REF][START_REF] Hupert | Conductive diamond thin-films in electrochemistry[END_REF]Brillas et al. 2009).

EAOPs produce strong oxidants, such as hydroxyl radicals which are able to degrade in a nonspecific way nearly any type of pollutants as they are the strongest oxidants after fluorine (E 0 = 2.8 V vs SHE in acidic medium) (Eq. 8). EAOPs are able to achieve both degradation and mineralization of a large range of organic pollutants and effluents such as landfill leachate [START_REF] Oturan | Occurrence and removal of organic micropollutants in landfill leachates treated by electrochemical advanced oxidation processes[END_REF][START_REF] Zhou | Electrochemical oxidation of biological pretreated and membrane separated landfill leachate concentrates on boron doped diamond anode[END_REF], pesticides (Rodrigo et al. 2014b;Dominguez et al. 2018a), dyes [START_REF] Alcocer | Comparative study for degradation of industrial dyes by electrochemical advanced oxidation processes with BDD anode in a laboratory stirred tank reactor[END_REF][START_REF] Nidheesh | An overview on the removal of synthetic dyes from water by electrochemical advanced oxidation processes[END_REF] or pharmaceuticals (García-Montoya et al. 2015). The efficiency of these processes mainly depends on electrode materials, configuration of the electrochemical device, operating conditions, etc. (Monteil et al. 2018).

Boron-doped diamond (BDD) is a particularly interesting anode material as it has a strong ability for oxidation of organic compounds and a long lifetime (Panizza et al. 2001;Actis et al. 2008;[START_REF] Cañizares | Conductive-Diamond Electrochemical-Oxidation in the treatment of effluents from door-manufacturing factories[END_REF]Brillas et al. 2009;Wei et al. 2011). BDD has a high overvoltage of oxygen evolution reaction allowing a large window for generation at its surface weakly adsorbed hydroxyl radicals.

The influence of the different operating parameters on the efficiency of these processes is now well-known and understood thanks to studies performed at lab-scale in batch conditions. The current, the concentration of pollutant or the concentration of the catalyst (mainly Fe 2+ ion)

were studied in order to optimize the kinetics of degradation at the lower cost. Besides, several studies have identified the mineralization pathways of various pollutants (Sirés et al. 2010;Dirany et al. 2012;Loaiza-Ambuludi et al. 2013;Bocos et al. 2016b;Dominguez et al. 2018a;[START_REF] Ganiyu | Electrochemical mineralization of sulfamethoxazole over wide pH range using FeII FeIII LDH modified carbon felt cathode: degradation pathway, toxicity and reusability of the modified cathode[END_REF], and sometimes the toxicity of the resulting solution was assessed (Oturan et al. 2008c;Dirany et al. 2012). The challenge is now to scale up the process in order to move towards industrialization for removal of recalcitrant organic pollutants from water. Particularly, the development of continuous flow reactors is required in order to overcome the traditional limitations of conventional batch reactors for the treatment of large flow of effluents [START_REF] Liang | Electrochemical degradation of cyanobacterial toxin microcystin-LR using Ti/RuO2 electrodes in a continuous tubular reactor[END_REF][START_REF] Basha | Electrochemical degradation of textile dyeing industry ef fl uent in batch and flow reactor systems[END_REF][START_REF] Ling | Continuous multi-cell electrochemical reactor for pollutant oxidation[END_REF][START_REF] Pillai | Batch and continuous flow anodic oxidation of 2 , 4dinitrophenol : modeling , degradation pathway and toxicity[END_REF][START_REF] Rahmani | Degradation of azo dye C.I. acid red 18 using an eco-friendly and continuous electrochemical process[END_REF][START_REF] Pillai | Performance analysis of a continuous serpentine flow reactor for electrochemical oxidation of synthetic and real textile wastewater : energy consumption , mass transfer coef fi cient and economic analysis[END_REF][START_REF] Borzyszkowska | Electrochemical degradation of 5-FU using a flow reactor with BDD electrode : comparison of two electrochemical systems[END_REF].

A first step in scaling up is to move from lab-scale batch studies to the development of engineered pilot-scale continuous reactors. At this scale, only few studies have analyzed the influence of operating parameters such as current and catalyst concentration on the performance of the reactor (El-Ghenymy et al. 2013a;Olvera-vargas et al. 2015;Ren et al. 2016). Few studies have considered the influence of hydrodynamic conditions in the reactor, which becomes a key parameter for this kind of application. For instance, the link between the hydrodynamic and the degradation of phenol was examined by [START_REF] Polcaro | Characterization of a stirred tank electrochemical cell for water disinfection processes[END_REF] and [START_REF] Saravanathamizhan | Residence time distribution in continuous stirred tank electrochemical reactor[END_REF]. The reactor had a simple hydrodynamic configuration, which was represented by a continuous-stirred tank reactor (CSTR) [START_REF] Polcaro | Characterization of a stirred tank electrochemical cell for water disinfection processes[END_REF] or by a CSTR with a bypass, active and dead zones with exchange flow between the active and the dead zone [START_REF] Saravanathamizhan | Residence time distribution in continuous stirred tank electrochemical reactor[END_REF]. In another study, a continuous multi-cell reactor was implemented and the model of a plug flow reactor was found to describe accurately this system [START_REF] Ling | Continuous multi-cell electrochemical reactor for pollutant oxidation[END_REF]. However, there is a real need to go further in the analysis of the link between hydrodynamics and the efficiency of the continuous reactors since hydrodynamics governs the characteristic times of mixing and the retention time of pollutants in the reactor. The hydrodynamic behavior of pilot scale continuous reactors is in general more complex than lab scale reactors that follow CSTR behavior. Some behavioral deviations from ideal reactors may appear such as the presence of dead and recirculation zones or short circuit that affect the global efficiency of the process.

There is thus a need to perfectly know the influence of the operational parameters such as flow rate, reactor configuration or aeration on the hydrodynamic behavior of the reactor in order to optimize its efficiency.

In this context, the present work focuses on a new pilot-scale reactor operating in continuous mode. Hydrochlorothiazide (HCT), a diuretic, was chosen as a model pollutant as (i) it could be found in effluents from hospitals (Mendoza et al. 2015), (ii) it is hardly removed by wastewater treatment plants (Margot et al. 2015), (iii) its occurrence in surface water has been evidenced (Bouissou-Schurtz et al. 2014) and (iv) it is harmful toward the environment (Fernández et al. 2010). First, an empirical design of experiments (DOE) approach was used to identify the most critical operating parameters for the efficiency of the process such as current density, iron concentration, flow rate and the reactor configuration with the number of electrodes and the distance between them. Secondly, the influence of flow rate and reactor configuration on the efficiency of hydrodynamics was studied in depth and linked with the efficiency of the process. Finally, a hydrodynamic retention time distribution (RTD) model was developed and combined for the first time with a kinetic model in order to describe the mineralization of pollutants by EAOPs in such new pilot scale continuous reactor.

Materials and methods

Chemicals

The pharmaceutical HCT (C7H8ClN3O4S2, purity >99%) and sodium sulfate (Na2SO4), the supporting electrolyte, were purchased from Sigma Aldrich. HCT solutions were prepared with ultrapure water obtained from a Millipore Milli-Q Simplicity 185 system with resistivity >18

MΩ cm at 25 °C. The pH of the solutions was adjusted using analytical grade sulfuric acid or sodium hydroxide (Acros). Organic solvents and other chemicals used were HPLC or analytic grade from Sigma-Aldrich, Fluka and Merck.

where TOC0 is the the TOC at the beginning of the experiment and TOCf, the TOC at steady state. For both mineralization and degradation, the quasi steady state was reached after 3 residence times (Annex Fig. 1). The modelling of the kinetic was performed by the software Aquasim® [START_REF] Reichert | Computer program for the identification and simulation of aquatic system[END_REF].

Mineralization current efficiency (MCE) and energy consumption (EC)

The EC was determined to evaluate the operating cost of the process. The EC formula (Brillas et al. 2009) was modified to fit the continuous mode is given in Eq. 36.

#"(abℎ/efg") = E * < ,-./ ^ --./ `3 * h * 6i00

(36)

with E, the cell voltage (V), I the current (A) and D the flow rate in L s -1 .

The MCE was used to estimate the efficiency of the reactor in the treatment of the HCT solution.

The MCE usual formula (Brillas et al. 2009) was modified in order to be consistent with the continuous mode (Eq. 37).

!"#(%) =

) * ,-./ ^01-./ `3 j ; < h * 100 (37

)
where m is the number of carbon atoms of HCT (7), M is the molar mass of the carbon (12 000 mg mol -1 ) and n is the number of electrons used to mineralize one molecule of HCT following the Eq. 38 assuming that N atoms of HCT were mainly converted to NH4 + .

C7H8ClN3O4S2 + 18 H2O ® 7 CO2 + 32 H + + 2 SO4 2-+ 3 NH4 + + Cl -+ 30 e - (38)

Design of experiments (DOE)

To have an empirical overview of the influence of selected operating parameters on the degradation, the mineralization, the EC and the MCE during the EF and the AO treatments, the Doehlert experimental design was applied. This method requires the selection of input The response function (Y) is a second order polynomial function expressed by Eq. 40:

To identify the relative influence of different operating parameters on the degradation of HCT and mineralization of its solution, and therefore on the MCE and EC, a DOE method was applied. Three parameters were considered: i) the current density, ii) the flow rate and iii) the catalyst (Fe 2+ ) concentration. The 14-electrode configuration was used in this part. The thirteen different experiments and their results are detailed in Table 14. With regard to the oxidative degradation of HCT, a complete degradation was observed for several conditions and thus the model could not be applied. However, it was evidenced that the current was by far the most important parameter for the degradation efficiency of the reactor. This is in accordance with batch experiments [START_REF] Ellouze | Application of Doehlert design to the electro-Fenton treatment of bismarck brown Y[END_REF]). Additionally, other experiments have been performed

in order to compare two reactor configurations.

Mineralization, EC and MCE

Unlike the degradation, the mineralization fitted well with the polynomial model. The different coefficients estimated are given in Fig. 41a. A correlation coefficient of 0.972 and an adjusted one of 0.9 were obtained with these coefficients values thanks to Nemrod® software. In the range of operating parameters selected, from the coefficient values obtained it appears that the most important parameter was the flow rate (X2) followed by the current (|ä 7 | > |ä 9 | > |ä 6 |). The catalyst concentration appears to have much less influence on the mineralization than the two other operational parameters. This suggests that AO is the main process influencing mineralization. Although EF is generally more efficient than AO for degradation, the mineralization efficiency differs little (Guinea et al. 2010;Feng et al. 2014). This point will be discussed deeply in a next sub-section.

The positive sign of b1 and b3 means that the mineralization rate will increase by increasing the catalyst concentration and the current. This fact was expected as by increasing the current more hydroxyl radicals are formed and the increase in catalyst concentration (in a certain range)

promotes the Fenton reaction. The minus sign of the coefficient b2 means that to increase the mineralization rate, the flow rate has to be reduced; this is due to the longer treatment time inside the reactor. The important absolute value of coefficient b12 means that the current and the flow rate are connected and govern mainly the efficiency of the mineralization process. The minus sign indicates that those two parameters work in the opposite way, one brings a positive effect while the other plays a negative effect on the mineralization. This is consistent with the sign of the coefficient b1 and b2 which are opposite. Thus, for a better mineralization rate, the current should increase and the flow rate decrease. This behavior was verified by analyzing the results of run 4 to 8 (Annex Table 1). It can be observed that the higher is the flow rate, the lower is the mineralization. This result can be explained by the lower contact time of the solution with hydroxyl radicals. The effect of the current was also verified with run 9-11 (Annex Table 1) since an increase of the current from 10 to 22 mA cm -2 leads to an increase of the mineralization rate from 50% to 76%.

Thanks to the polynomial function Y, it is possible to predict the mineralization rates in the area defined by the upper and lower values of each parameter. Indeed, by replacing all the coded values: X1 (current), X2 (flow rate) and X3 (Fe 2+ concentration) in Eq. 46, the mineralization rate (Y) can be estimated. This can be illustrated by contour plots. Fig. 41b illustrates the flow rate versus the current at a given catalyst concentration at the center of its range (X3=0). For instance, the line blue 40 corresponds to the "iso-mineralization rate" of 40% representing all the conditions of flow rates and currents where the mineralization rate is of 40% at a defined Fe 2+ concentration of 0.13 mM. b a concentration has nearly no impact. The lower is the current, the higher is the MCE and the lower is the EC. This opposite influence of the high currents can be explained by scavenging reactions that occur at higher current intensity, for example water oxidation at the anode and H2 evolution at the cathode (Zhang et al. 2007). Finally, a compromise between a high degradation and an acceptable EC and MCE has to be chosen. In other terms, a good balance for the current and flow rate is required.

Finally, this process is really advantageous compared to more conventional processes because of its high mineralization power allowing high mineralization rate. For example, 97% of the initial TOC value could be removed under 47.6 mA cm -2 current, flow rate of 20.4 mL min -1 , 0.2 L min -1 of air flow rate and 50 mM of Na2SO4. This process is much more efficient than an activated sludge treatment were only 30% of HCT is in general eliminated in conventional wastewater treatment plants (Margot et al. 2015). It seems also to be more efficient than O3 treatment where only 80% of the solution is mineralized [START_REF] Rosal | Ozone-Based Technologies in water and wastewater treatment[END_REF]. However, even if this process is very efficient compared to the conventional methods, some energy is wasted in side reactions as can be seen with the MCE values of Fig. 42b. Thus a compromise should be reached between high efficiency and cost effective process. Therefore the combination of electro-Fenton with a biological treatment can constitute a promising option.

The goal of such a coupling being to transform the non-biodegradable compounds into biodegradable species in a first stage by electro-Fenton, and then apply a biodegradation step to reach high mineralization degrees.

Influence of the configuration on the mineralization rate and EC

To deeper understand the interactions between reactor configuration and operational parameters, two configurations were tested: the first with 14 electrodes and a gap of 4 cm compared to other studies. Adding a cathode on the floor or sides of the reactor directly linked to the air bubbling to promote H2O2 production could constitute a solution to this issue.

The role of the configuration is very important as it drastically influences the mineralization rate and the cost of the process. The comparison between the two configurations leads to the following statement: for both point of view, the 28-electrode configuration is better, with shorter distance between the electrodes and higher active surface area.

This part has demonstrated that the flow rate and the configuration play an important role on the EC and the mineralization rate. Thus, as they are closely associated to the path followed by the pollutants in the reactor, a deeper analysis of the hydrodynamics in the reactor was carried out.

Hydrodynamics

RTD curves

RTD curves were drawn in order to identify the variation of flow pattern in the reactor and to choose the most suitable model to characterize the pilot. Fig. 44 shows the RTD curves obtained for the system containing 28 electrodes at different flow rates. The equivalent chart for 14-electrode configuration is given in Annex Fig. 2. The parabolic shape of the curves shows that there is no by-pass or short-circuit in the reactor. One can observe that the higher is the flow rate, the thinner is the peak and the smaller is the area of the peak. The lower slope at the end of the curves highlights the presence of stagnant zones [START_REF] Haoran | The influence of porous structure and biofilm on the hydrodynamics of two types of trickle filters[END_REF]. For both 14and 28-electrode configurations, it can be observed similar shapes. Thus, the same model is applied for both configurations, as described below.

Fig. 44: RTD functions plotted against the time for different flow rates under 0.2 mL min -1 of air with 28electrode configuration.

The mean retention time (fp) can be estimated as the time when half of the area of the curve is reached. As the peaks are quite symmetric the mean retention times can be read on the curve at the top of each peak. Table 16 compares the theoretical retention time (τ) with the Ts for the two configurations at different flow rates. The τ is higher than Ts, which is a characteristic phenomenon observed when dead volumes are present in the reactor. Furthermore, fp for 28electrode configuration is longer than fp for 14-electrode configuration because of additional troublemakers in the 28-electrode system that lengthen the flow path inside the reactor (Annex Fig. 4).

- Because of the symmetric aspect of the curves and the presence of dead volumes, the model of continuously stirred tank reactors in series with dead volumes was selected and compared to the classical plug flow model. The CSTR model was also proposed as the vertical bubbling enable to realize an efficient mixing of the solution.

Model

In order to confirm this hypothesis, the experimental were compared to the CSTR in series model. The RTD (#(å)) for the CSTR model is given by the Eq. 47.

#(å) =

with m, the fraction of active volume, and N the number of continuously stirred tank reactors.

When N approaches 1, the reactor is similar to a CSTR and when N approaches infinity, the model is approaching a plug flow. Both m and N were estimated thanks to the Excel solver. In order to check the strength of the model, different initial parameter (N and m) were taken and the correlation coefficient was calculated by Eq. 11. All the correlation coefficients were over 0.9928 (Table 18) which confirms the suitability for the model.

This CSTR in series model was then compared to the plug flow model describe by Eq. 48.

Experimental parameters (N and m)

As flow rate is directly related to the RTD, the relation between this parameter and the number of CSTR as well as the fraction of active volume was studied in order to characterize the reactor for both configurations (Fig. 45). against the liquid flow rate, under 0.2 mL min -1 of air with 14-and 28-electrode configurations.

As shown in Fig. 45, for the 14-electrode system, the fraction of active volume is increasing with the liquid flow rate, being of 80% at 20.4 mL min -1 and of 92% at 128 mL min -1 . In the reactor the mixing and the shear rate are due to: (i) the flow rate in the reactor and the contact with the obstacles (electrodes) and (ii) the air flow of small bubbles from the air diffuser. Thus by increasing the liquid flow rates, the liquid velocity is increasing and the turbulence around the electrodes is also increasing leading to a better mixing and thus to lower dead zones and stagnant zones in the reactor at high flow rate. Higher liquid flow rate enhances the mixing between active and dead zones resulting in the activation of a portion of the volume of the stagnant zone [START_REF] Saravanathamizhan | Residence time distribution in continuous stirred tank electrochemical reactor[END_REF][START_REF] Djoudi | Flow modeling in electrochemical tubular reactor containing volumetric electrode : application to copper cementation[END_REF]. For the 28-electrode system, the fraction of active volume is nearly the same for all the flow rates and is higher than that obtained for the 14-electrode system. The difference between the two configurations can be explained by the larger number of electrodes which increases the turbulence in the reactor leading to a lower dead volume [START_REF] Djoudi | Flow modeling in electrochemical tubular reactor containing volumetric electrode : application to copper cementation[END_REF]. However, at the highest flow rate (127.5 mL min -1 ), the fraction of active volume is the same for both configurations due to the high velocity promoting a better mixing in the reactor.

The number of CSTR in series increases with the liquid flow rate for both configurations (Fig 45). At low flow rate, the reactor approaches a few continuously stirred tank reactors behavior and by increasing the flow rate it tends to a plug flow reactor [START_REF] Yuan | Liquid phase residence time distribution for a twophase countercurrent flow in a packed column with a novel internal[END_REF][START_REF] Furman | Tracer investigation of a packed column under variable flow[END_REF][START_REF] Saravanathamizhan | Residence time distribution in continuous stirred tank electrochemical reactor[END_REF][START_REF] Cao | Computational simulation of fluid dynamics in a tubular stirred reactor[END_REF][START_REF] Djoudi | Flow modeling in electrochemical tubular reactor containing volumetric electrode : application to copper cementation[END_REF]. The slope inflection at high flow rate highlights that the effect of the increase in the flow rate on the number of CSTR is lower in these conditions. The number of CSTR in series is higher for the configuration with 28-electrode configuration than for the configuration with 14 electrodes, whatever the liquid flow rate. This means that by increasing the number of electrodes, the hydrodynamic behavior tends more towards a plug flow reactor. This observation is logical since by increasing the number of electrodes, the number of "reactors perfectly mixed" is increased due to more complex water flow provided by the series of baffles (electrodes) under which the wastewater is forced to flow. However, this model has some limitations: i) the model used is a first stage analyze and does not reflect all the reactions occurring in the solution (see section 3.4.1), ii) the aeration may have a role to play in the hydrodynamics of the reactor but this aspect is not taken into account, iii) only the steady state is considered as the calculations could be simplified but an overall understanding of the process may be interesting.

Conclusion

In this study, a DOE was first carried out in order to find the more important parameters regarding the mineralization rate, the MCE and the EC for mineralization of HCT aqueous solution. It was shown that the flow rate was more impacting on the mineralization rate whereas the current was more important for the EC and the MCE. The lower was the flow rate, the higher was the mineralization rate due to the longer treatment time. Some deeper experiments were then carried out to evidence the effect of the reactor configuration on the mineralization efficiency and the operating cost. The 28-electrode configuration was found to be better than the 14-electrode configuration, the former promoting the mixing with more troublemakers and having a longer flow path.

To better understand the important role of the flow rate and the configuration, an hydrodynamic study was performed. RTD curves were obtained and compared to the CSTR in series with dead zones and to the plug flow models. These CSTRs suited better with the RTD curves thus, this model was chosen to model the reactor. With this model, the number of CSTR, the active volume and the mean retention time were studied for different flow rates and configurations.

The hydrodynamic study confirms the DOE: the 28-electrode configuration had a lower dead volume thanks to the better mixing and a lower EC. Low flow rates approached a perfectly mixed tank reactor with high mean residence time.

Finally, a mathematical model for the mineralization of HCT solution was established, combining a pseudo first order kinetic of mineralization with the hydrodynamic model. This model was compared to experimental points and a good IOA was calculated demonstrating the good suitability of this model. This is the first time that a model was developed combining both hydrodynamics and kinetic to describe an electrochemical pilot. This study enables to predict for this type of reactor the % of

Introduction

In our today's life pharmaceuticals are highly consumed especially to treat diseases or to prevent them. Once being used, they enter the environment mostly by human and animal excretion and discharge of polluted effluents coming from pharmaceutical industries, domestic and hospital effluents [START_REF] Duong | Occurrence , fate and antibiotic resistance of fluoroquinolone antibacterials in hospital wastewaters in Hanoi, Vietnam[END_REF]Kümmerer 2009b). Their occurrence in the environment is due to the lack of efficiency of the existing treatment in wastewater treatment plant (WWTP) mostly based on a biological elimination [START_REF] Stackelberg | Efficiency of conventional drinking-watertreatment processes in removal of pharmaceuticals and other organic compounds[END_REF].

Several treatment techniques are used to eliminate the organic micro-pollutants; some of them are based on the transfer of the pollutant from one phase to another one like physico-chemicals techniques such as adsorption [START_REF] Rivera-Utrilla | Pharmaceuticals as emerging contaminants and their removal from water. A review[END_REF]) and membrane technologies [START_REF] Göbel | Fate of sulfonamides, macrolides, and trimethoprim in different wastewater treatment technologies[END_REF]. Others are directly eliminating the pharmaceutical such as advanced oxidation processes (AOPs) [START_REF] Klavarioti | Removal of residual pharmaceuticals from aqueous systems by advanced oxidation processes[END_REF]Oturan and Aaron 2014). AOPs are based on the formation of strong oxidants such as hydroxyl radicals, the second strongest oxidant after fluorine, able to eliminate nearly any type of organic molecules (Brillas et al. 2009;Sirés and Brillas 2012). Among the AOPs, electrochemical advanced oxidation processes such as electro-Fenton (EF) or anodic oxidation (AO) are commonly used as they are i) very efficient to remove organic pollutants from water, ii) environmentally friendly and iii) easy to set up (Panizza and Cerisola 2001;Brillas et al. 2009;Sirés et al. 2014;Olvera-Vargas et al. 2018). For these treatments, the operating conditions and the nature of the anode are important parameters that significantly influence the efficiency of the treatment. Among anodes, boron doped diamond (BDD) and platine (Pt) are often used. BDD is used because of i) its high oxidation efficiency, ii) its high current efficiency and iii) its long lifetime (Actis et al. 2008;Brillas et al. 2009;Wei et al. 2011;Oturan et al. 2012;Sopaj et al. 2016). Comparatively, Pt is less efficient as its oxidation power is lower compared to BDD (Panizza and Cerisola 2005;Rodrigo et al. 2010).

BDD mineralizes quickly the organics without enabling the formation of a lot of intermediates whereas the Pt enables to accumulate intermediates such as carboxylic acids (CAs).

One of the most important drawback of EAOPs is the relatively high operating cost in the case of long treatment times. To overcome this disadvantage, the combination of an EAOP and a biological treatment to decrease the operating costs while keeping a good mineralization rate has been proposed (Ganzenko et al. 2017;Monteil et al. 2018). Two different treatment modes can then be employed, weather beginning by the biological treatment when the effluents contain high concentration of biodegradable products or starting with the EAOP when the solution is not biodegradable (Oller et al. 2011;Ganzenko et al. 2014;Olvera-Vargas et al. 2016a;Monteil et al. 2018). Some studies investigated the effect of such a combined treatment on the elimination of pharmaceuticals at lab scale and highlighted that this combination can be very effective (Ferrag-Siagh et al. 2013;Ferrag-Siagh et al. 2014;Mansour et al. 2014;Mansour et al. 2015a;Mansour et al. 2015b;Olvera-Vargas et al. 2016a;Olvera-Vargas et al. 2016b;Ganzenko et al. 2017;Roshini et al. 2017;[START_REF] Zaghdoudi | Direct and indirect electrochemical reduction prior to a biological treatment for dimetridazole removal[END_REF][START_REF] Aboudalle | Metronidazole removal by means of a combined system coupling an electro-Fenton process and a conventional biological treatment : by-products monitoring and performance enhancement[END_REF]).

However, a key step in the process control of this type of coupling is to better understand how the operating conditions applied on the EAOP will influence the efficiency of the biological treatment.

In this context, this chapter aims at trying to explain why some conditions are more interesting based on the nature of the anode (BDD or Pt) and on the impact of the parent pollutants and the by-products. The diuretic hydrochlorothiazide (HCT) was chosen as the model pollutant for this study as i) it is hardly eliminated by wastewater treatment plants (Margot et al. 2015), ii) it was found in surface water (Bouissou-Schurtz et al. 2014) and iii) it is toxic to the environment (Fernández et al. 2010). The EF and AO treatments of this compound with a pilot scale reactor were investigated in a previous chapter (Chapter 3). The first part of this chapter will focus on the interest of performing a combined process showing the inadequacy of a biological treatment alone and the possibility to decrease the operating costs. Some specific conditions are required to efficiently perform this combined treatment. The second part deals with identifying them.

Finally, the third part uses the specific conditions found previously to obtain a good mineralization rate.

Material and methods

Chemicals

All chemicals used in this study were analytical grade and used as received without further purification. The pharmaceutical HCT (C7H8ClN3O4S2, purity > 99%) was supplied by Sigma Aldrich. Iron (II) sulfate heptahydrate (catalyst source, 99%) was purchased from Acros Organics. Sodium sulfate (Na2SO4) (supporting electrolyte) was supplied from Sigma-Aldrich.

MgSO4 (Chimie-Plus Laboratoire), Cacl2 2H2O (Merck), FeCl3 (Sigma Aldrich), NH4Cl

(Merck), Na2HPO4 (Acros), K2HPO4 (VWR) and KH2PO4 (Acros) were used as nutrients or buffer for the biological treatment. Short-chain CAs (oxalic, oxamic, acetic and maleic acids were obtained from Acros, Fluka and Alfa Aesar. Solutions were prepared with osmosis water equipment purchased from Cloup which keeps conductivity below 20 µS cm -1 . The pH of solutions was adjusted using analytical grade sulfuric acid (Acros) and sodium hydroxide (Fluka). Organic solvents used were HPLC or analytic grade from Sigma-Aldrich, Fluka and Merck.

Electrochemical treatment

HCT solutions were prepared at different concentrations in 50 mM of sodium sulfate (as electrolyte) and the solution pH is adjusted to 3.0 by a CyberScan 1500 from Eutech Instruments and a 1 M H2SO4 solution. In the case of the AO oxidation the solutions were left as stated above for the EF treatment, 0.1 mM of ferrous iron was added to the solution, a usual concentration for the EF degradation.

Batch treatment

The different batch experiments were performed in an open undivided cylindrical cell of 250 mL. The electrochemical cell was composed of whether a BDD or Pt anode and a carbon felt cathode. The anodes had around 30 cm 2 surface area. BDD was purchased from CONDIAS GmbH in Germany. The cathode was a 3D structure measuring 18 * 5.0 * 0.5 cm 3 ; it was placed on the inner wall of the cell covering the total internal perimeter. It was purchased from Mersen (France). The solutions were stirred with a magnetic bar. The degradation and the mineralization of the synthetic effluent were performed under current-controlled conditions using a Hameg HM8040 triple power DC supply. Continuous saturation of O2 was ensured by bubbling air starting 10 min before starting the electrolysis.

Treatment with electrochemical pilot

The electrochemical pilot is made of plexiglass and has a capacity of 1.7 L (Fig. 50). The reactor is composed of 14 anodes and 14 cathodes separated by 2 cm. Anodes are made of BDD film on niobium support (from CONDIAS GmbH) of 30 cm 2 and cathodes are made of 3D carbon felt of 30 cm 2 (from Mersen). Two peristatic pumps FH100 drive the flow throughout the reactor. Aeration was set at 0.2 L min -1 provided from the whole surface of the bottom of the reactor and is adjusted by an air flowmeter. A constant current was delivered by a Microlab power supply to the electrodes mounted in parallel.

Fig. 50:

The pilot reactor of 1.7 L and the whole set up.

Biological treatment

Three inoculums were used for the biodegradation assays. The first one was taken from an activated sludge reactor from the WWTP from Jablines. The second inoculum was taken from the mixed liquor of an aerobic biofilter from the WWTP Saint-Thibault des Vignes (named biofilters). The last inoculum used was a sample of µ-organisms already acclimated to the pollutant and to the CAs from a previous treatment (named selected µ-organisms). The µorganisms were centrifuged, washed with tap water and then added to a new solution.

The initial volatile suspended solid (VSS) concentration was measured for every sample and a dilution factor was calculated to reach an initial concentration of VSS of 0.01 gVSS L -1 for each biodegradation test. This low value was chosen to avoid TOC measure disruption with soluble organic matter coming from biomass.

The biodegradation tests were performed in an erlenmeyer flask of 500 mL closed with cotton to enable gas exchange but prevent contamination. The biological reactor was operated as follows: pH adjustment with 1 M Na2SO4 and 1 M NaOH to keep the pH in the range 6.5-7.5;

mechanical stirring using a rotative plate at 122 RPM (revolution per minute) provided by Merck and room temperature (22 ± 3) °C. Different nutrients compositions were tested especially with different N concentrations (5 to 20 mg L -1 ) and no effect was observed between experiments, meaning that nutrients were not limiting. From these observations the following concentrations were applied: 18.75 mg L -1 of NH4Cl (corresponding to 5 mg L -1 N-NH4), 36.5 mg L -1 of Cacl2 2 H2O, 0.26 mg L -1 FeCl3 6 H2O and 11 mg L -1 of MgSO4. The phosphate buffer used was composed of 334 mg L -1 of Na2HPO4, 85 mg L -1 of KH2PO4 and 208 mg L -1 of K2HPO4 [START_REF] Rodier | Analyse de l'eau[END_REF].

Experiments with and without external addition of air were also performed and no difference was observed since oxygen was in excess, thus no external aeration was applied. The effect of the addition of sodium sulfite to eliminate the residual hydrogen peroxide was tested on the bacteria and no effect was found thus 5 mL of a 0.5 M sodium sulfite solution was added for all the biodegradation tests.

A bio-adsorption study was carried out to confirm that the degradation was due to a real elimination and not to the transfer of the pollutant to another phase. After 5 min, 1 h, 2 h and 6 h samples were taken of the HCT mixed with the bacteria and its concentration remained constant, meaning that adsorption was negligible.

carbonates. Reproducible values with ± 2% accuracy were detected. The injection volumes were of 50 µL.

Liquid chromatography (LC) -mass spectroscopy (MS) analysis

LC-MS analyses were performed to identify Aromatic intermediates (AIs) formed during the AO treatment of the HCT. The same Hypersil Gold column at 40°C and a liquid flowrate of 0.2 mL min -1 were used. After extracting from the treated solution with ethyl acetate, the sample was completely dried under vacuum. Then it was diluted in 1 mL of methanol / water 50/50 (v/v) and eluted by 0.1% formic acid / methanol 90/10 (v/v) solution. Concerning the mass conditions, the ESI (electrospray ionization) mode was used at 90 V.

Results and discussion

Interests of combining EAOPs and a biological treatment

To analyze the interest of combining AO or EF with a biological treatment, each treatment was first studied alone.

Inefficiency of a biological treatment alone in the case of persistent organics

An experiment was carried out in batch reactor to assess the ability of µ-organisms to eliminate the HCT without pre-treatment using an EAOP. Fig. 51 shows the concentration dynamics of HCT during one month in the biological reactor. The HCT concentration remained stable, highlighting the inability of the µ-organisms from WWTP to eliminate HCT despite the great microbial diversity of the inoculum. This result is in agreement with the low mean percentage of elimination of this molecule (30%) in wastewater treatment plants in Europe and United

States (Margot et al. 2015).

Fig. 51: Evolution of HCT concentration as a function of time for 0.1 mM HCT solution mixed with 0.1 gVSS L -1 of activated sludge.

AO to mineralize a HCT solution

The mineralization of 0.1 mM HCT solution was performed in the pilot-scale reactor by AO in the continuous mode. Results of the mineralization rate for different currents are given in Fig.

52

. As already studied in chapter 3, it was possible to observe that the mineralization rate at the pilot output depends on the operating conditions and can reach high mineralization rates.

Electro-chemical pre-treatment followed by a biological post-treatment

Biological treatment alone appeared inefficient whereas AO is really efficient but with an EC that increases with treatment time (chapter 2). In order to get a cost effective treatment, the feasibility of the combination of an EAOP and biological treatment was assessed. To evaluate the cost of the energy saved, the price (P) for the treatment of 1 m 3 of 0.1 mM HCT solution in the pilot reactor was estimated according to the Eq. 56. Whatever the experimental conditions, only a small part was mineralized by the biological treatment: respectively 9, 7, 7 and 2% for conditions A, B, C and D. These results highlight that the solutions are certainly composed by molecules very low biodegradable even by an inoculum with a large microbial diversity. This was confirmed by an analysis of the more concentrated intermediates during the biological treatment. Indeed, thanks to the uHPLC analysis, the peak area of these AIs was found to remain stable highlighting their non-biodegradability by the µorganisms. Then, by using LC-MS analysis, their structure were found and confirmed that these molecules were still containing aromatic groups difficultly degraded by µ-organisms (Table 20). Thus, the biodegradable fraction is certainly composed by smaller molecules and especially CAs, known to be biodegradable (Ganzenko et al. 2017). For condition D, the low biodegradation rate could also be linked to the very low carbon concentration remaining in the solution after the electrochemical treatment (5 mgC L -1 ), organic carbon certainly becoming limiting for the µ-organisms. In this part, it was demonstrated that combining an electrochemical treatment with a biological degradation could reduce the operating costs. However, the conditions applied in this part for the AO treatment cannot lead to an efficient biological degradation probably due to the presence of AIs. Thus, next section will focus on finding the best operating conditions to provide biodegradable compounds and eliminate the non-biodegradable species. Four CAs were formed during the degradation of HCT in batch reactor according to a previous study of our group [START_REF] Monteil | Efficient removal of diuretic hydrochlorothiazide from water by electro-Fenton process using BDD anode: a kinetic and degradation pathway study[END_REF] and chapter 2 of this work). Based on this observation, these 4 molecules were chosen as model of CAs and their relative concentrations were kept as already detected (Table 21). Three experiments were carried out using 250 mgC L -1 with 0.01 gVSS L -1 of inoculum. For experiments 1, 2 and 3 (Fig. 55), the inoculums were respectively activated sludge, biofilters and selected µ-organisms. For all the experiments, the CAs were sequentially degraded in the same order: first the maleic acid, then the acetic acid followed by the oxalic acid and finally the oxamic acid. The maleic acid is certainly transformed into smaller molecules thus, this molecule may not be mineralized first. Concerning the difference between the oxamic and the oxalic, it can be supposed that it is due to the amine function. The degradation rate was almost zero for the oxamic, oxalic and The degradation rate using the treatment with activated sludge biomass as inoculum was quicker than the mixed liquor from the biofilters but slower than the bacteria from the selected biofilters. Indeed, when bacteria are pre-selected, the concentration of active bacteria for these CAs is higher and the µ-organisms are already acclimated to the chemical conditions reducing significantly the lag phase resulting in a faster degradation kinetics. However, even with this pre-selection, the sequential degradation is still present with the maleic acid degraded first (0.5 day) and the oxamic acid the last.

The sequential degradation can be linked to the carbon catabolite repression which is an important regulatory mechanism for µ-organisms. It allows µ-organisms to preferentially use easily metabolizable carbon sources over relatively less favorable carbon source. For instance this phenomenon is illustrated by diauxic growth during which bacteria assimilate firstly energy-efficient and rapidly metabolizable sugars (glucose) for which they already have the enzymes then less-favored carbohydrates for which they need to synthesize the enzymes [START_REF] Loomis | Glucose-lactose diauxie in Escherichia coli[END_REF]. This sequential degradation of the CAs has not yet been studied whereas it can be very important in the case of a continuous treatment where only the most degradable acids would be eliminated if a continuously stirred tank reactor (CSTR ) reactor was used. To face this problem, a reactor with a plug flow behavior could be used; the most easily degradable CAs being degraded in the first layers of the reactors and the less degradable ones in the last layers of the reactor.

Inhibiting effect of the pollutant and of the AIs on the biological degradation

To clarify the inhibiting effect of the HCT and of its AIs, three experimental conditions were compared: (i) the solution of 250 mgC L -1 of CAs alone, (ii) the solution of 250 mgC L -1 of CAs mixed with a solution of 0.42 mM of HCT and, (iii) As can be seen in Fig. 56, the CAs are degraded in the same order as in sub-section 3.2.1: first the maleic acid, then the acetic, oxalic and finally oxamic acids were degraded. For all the CAs, the solution containing only the CAs is degraded first and as expected the curve are very similar to the experiment of sub-section 3.2.1. Then, the solution containing the CAs and the AI is mineralized. The degradation of the solution with the HCT and the CAs requires a longer treatment time which is composed of two parts: (i) an accommodation time and (ii) a degradation period. For instance, for the oxamic acid (Fig 56d ), the accommodation time is of 2, 3 and 4 days for the CAs alone, the CAs and the AI, and the CAs and the HCT, respectively, whereas the degradation period is of 0.5 and 1.5 days for the CAs alone and for the CAs and the HCT, respectively. The accommodation time and the degradation time are longer with the AIs and the HCT than the solution of CAs alone, showing that the presence of HCT and its AIs, at least temporarily, slows down the activity of the µorganisms. This behavior is also verified in Fig 56b and56c for oxalic and acetic acids, respectively. For the maleic acid (Fig 56a), it is harder to observe the difference between the curves as this acid is very quickly degraded; shorter interval of sampling should have been taken to identify it. In a continuous treatment, the acclimation time is not important as the µorganisms would already be selected for the different CAs; however the kinetic of degradation is very important in order to eliminate the CAs during the mean residence time. As the kinetics of degradation is longer with the AIs and the HCT in the solution, it is essential to eliminate them or at least to reduce significantly their concentration in order to have an efficient biotreatment on the biodegradable part of the solution (CAs). This model and the results obtained highlight that the lower is the concentration of AIs the quicker is the degradation rate after the lag phase. The development of these inhibition models could help to optimize the operational conditions of the electrochemical treatment, the possible dilution and help to predict the performance of the combination of treatment. However more experiments need to be performed to achieve this objective. For condition D5, for the acetic acid, the acclimation phase was similar to that of the CAs alone and the degradation kinetic was only 25% lower than for CAs. A maximum concentration of no or little inhibition could then be defined for the intermediates of Table 20 thanks to their peak area on the uHPLC.

Application of appropriate operating conditions to optimize the combined treatment

The main objective of this section is to use the lessons learned from the previous sections to select optimal conditions for the electrochemical treatment in order to optimize the efficiency of the combined treatment. The experiment were performed at lab scale first as the operating conditions are easier to handle and then they will be carried out in the pilot reactor. very powerful oxidants and efficient to eliminate organic compounds (Panizza and Cerisola 2009). This high efficiency has to be controlled for the combined process in order to only convert the mother pollutant to biodegradable intermediates and to limit as much possible its mineralization. Thus a very low current was applied (30 mA) to degrade 230 mL of a 0.2 mM HCT solution. The AIs and the CAs were followed over time and the objective was to stop the treatment when the concentration of CAs was maximal with the lowest concentration of AIs and with a reasonable amount of carbon in the solution. However, even at a very low current it was not possible to eliminate all the intermediates while keeping a high value of TOC. The reaction was thus stopped after 2.5 h when the HCT was completely degraded. At this time, the final TOC value was 11.2 mgC L -1 . The concentration of AIs were below the concentrations found in sub-section 3.2.3 for the condition D5, meaning that their concentration were not influencing too much the microbial activity. The TOC value was followed during the postbiological treatment and the results are depicted in Fig. 59. 28% of the initial TOC was mineralized during the AO treatment. The following biological treatment was slow during the three first days corresponding to an acclimation phase, then an exponential increase lasted from day 5 to day 10 and finally a stabilization phase was observed until the end of the experiment.

The residual carbon (5.4 mgC L -1 ) at the end of the biodegradation corresponds to the non-biodegradable part which was not converted into biodegradable carbon by the AO. 38% of initial solution were mineralized by the biological treatment enabling to globally mineralize 66%. This combined treatment was thus successful as the mineralization was mainly due to the biological treatment enabling to save energy. Thus these operating conditions to eliminate the HCT and the AIs and form CAs were successful and they were adapted to the pilot. To achieve higher mineralization rates, another step of electrochemical and biological treatments on the remaining solution could be considered. In view of a combination of continuous electrochemical and biological treatments, a recirculation loop at the inlet of the electrochemical treatment could be considered.

Fig. 59:

Combined treatment of 230 mL a 0.2 mM HCT solution at 30 mA and 2.5 h followed by a post biological treatment using activated sludge with 0.01 gVSS L -1 of µ-organisms. The squared point corresponds to the electrochemical treatment and the dot points to the biological mineralization.

Finding good operating conditions for pilot scale for an efficient combined treatment

In the same way than in the previous section, the objective was to avoid mineralization and to promote the degradation of HCT and of AIs by applying a very low current density. However as the pilot reactor works in the continuous mode, the flow rate constitutes also a key factor and its value has to be chosen carefully. As discussed in the previous chapter, a low flow rate encourages a high level of mineralization whereas a high flow rate leads to mean retention time lower than the degradation of the pollutant. Based on the results obtained in chapter 3, it was chosen to use an intermediate flow rate of 42.5 mL min -1 . Three experiments were carried out with three different current densities in order to reach different mineralization rates: 4.0, 8.6 and 16.2 mA cm -2 to mineralize 40, 50 and 60% of the solution, respectively. During these experiments the CAs, the main AIs, the mother pollutant and the TOC were followed. Table 22 gives the results obtained for the three runs. As expressed in this table, three different mineralization rates were obtained: 40, 47 and 60% for run 1, 2 and 3, respectively. In order to choose the best operating conditions, three characteristics were compared: (i) the total amount of CAs, (ii) the concentration of the more concentrated AI and (iii) the percentage of elimination of the HCT. Concerning the degradation of HCT, run 2 and 3 attained more than 99% of removal whereas run 1 reached only 91%. The total concentration of CAs was of 0.017, 0.018 and 0.031 mM of C for run 1, 2 and 3 respectively. Finally, only run 3 attained the concentrations of AIs below the maximum value defined in sub-section 3.2.3 meaning that in run 1 and 2, the biodegradation could be affected by the presence of these AIs. Based on those criteria, the operating conditions of run 3 were selected as the best ones.

Finding good operating conditions with Platine electrode at batch scale

To find the optimal operating conditions for a combined treatment a suitable anode is required.

As it was demonstrated previously, the BDD anode is very efficient to mineralize a solution but its very high mineralization power makes the process uncontrollable. Thus, to choose a more suitable anode the process has to be well understood. In the case of active metal anode, the adsorbed hydroxyl radicals formed by oxidation of water (Eq. 58) can lead to the formation of high state oxide by oxygen transfer reaction (Eq. 59). This high state oxide performs a selective oxidation of organic compounds contrarily to the non-active anode (such as BDD) where the adsorbed hydroxyl radicals are able to non-selectively mineralized the solution which results in complete combustion of organics to CO2. Thus an active anode is more suitable for a selective oxidation of the pollutant. Pt was chosen as it is a common active electrode.

Conclusions

This chapter demonstrates the clear interest of a combined treatment between an EAOP and biological treatment to significantly improve the performance of biological treatment and to reduce the cost of the electrochemical treatment alone. The biological treatment alone was not able to degrade the solution of HCT. It was also demonstrated that the operating cost could be divided by 3 while using the pilot in the case of a pre-treatment of 1 m 3 HCT solution followed by a biodegradation step. for the acetic acid. The results obtained showed that for the dilution by 5, the concentration of AIs and HCT was not disrupting significantly the biodegradation. Thus, to perform an efficient combined treatment, it was evidenced that some specific conditions for the electrochemical treatment should be applied with the aim of i) degrading the pollutant HCT while reaching a low mineralization degree, ii) reducing the AI concentration and iii) promoting the formation of the biodegradable compounds like CAs.

Finally the combined treatment at lab scale was tested in the conditions described previously.

As BDD anode was used a low current was applied to minimize the mineralization and promote the formation of CAs. The combined treatment was successful as 66% of the solution was mineralized including 38% by the biotreatment. The same strategy was then applied to the pilot with three conditions tested, the best condition was a current of 6.8 A corresponding to a current density of 16.2 mA cm -2 , a flow rate of 42.5 mL min -1 leading to 60% mineralization before the biological treatment.

In order to favor the formation of CAs and a low mineralization rate of the solution, a pretreatment with electro-Fenton using a Pt electrode as an active anode was performed. With this anode, a mineralization degree of 85% was attained, including 50% of TOC removal by biotreatment which implies a drastic reduction in the operating costs. To conclude, combining an electrochemical treatment with a biodegradation seems to be a very efficient way to reduce the operating costs. However, some specific conditions should be applied and the use of an active anode could give even better results. Research should be done on this last aspect.

This thesis presents a way of treating pharmaceuticals from wastewaters. First electrochemical advanced oxidation processes experiments were performed at lab scale to understand the different operating parameters then the process was scale-up and experiments were carried out in a pilot reactor. Then based on these results, a model was designed to describe the mineralization rate in the reactor. Finally to reduce the operating costs, the electrochemical process was combined with a biological treatment.

EF treatment of the different molecules

The oxidative degradation and mineralization of four pharmaceuticals, namely hydrochlorothiazide (HCT), tramadol (TMD), Nadolol (Nad) and ofloxacin (OFL), by electro-Fenton (EF) were deeply studied in this work. The impact of two operating parameters (the catalyst concentration and the current) was considered for the four pollutants. For all of them increasing the Fe 2+ concentration from 0.05 to 0.1 mM (or 0.2 mM for TMD) increased the kinetic of oxidation of 0.1 min -1 as an average among the four pollutants. However a higher concentration slowed down the oxidation due to side reactions. A similar behavior was identified for the current. The absolute rate constant for the oxidation of the different molecules was determined using the competition kinetics method and found to be of (4.37 ± 0.04) Î 10 9 M -1 s -1 , (5.59 ± 0.03) Í 10 9 M -1 s -1 , (5.63 ± 0.03) Í 10 9 M -1 s -1 and (5.55 ± 0.05) Í 10 9 M -1 s - 1 for HCT, TMD, Nad and OFL respectively. The mineralization of these molecules was also successfully performed as after 6 h of EF all the solutions were totally mineralized. Thus the EF treatment was able to successfully eliminate the HCT, Nad, TMD and OFL with similar kinetic of degradation and totally mineralized the pharmaceutical solutions highlighting the non-selectivity of the hydroxyl radicals and their ability to treat pharmaceuticals.

which favors this second configuration. Two configurations were tested but this pilot gives the opportunity to try closer distance between electrodes. Thus by reducing the gap between the electrodes, the electrical resistance could be reduced and could make the pilot even more cost effective.

This pilot was proved to be very efficient, able to reach nearly total mineralization rate, thus to make it even more efficient an hydrodynamic study was carried out, aiming at identifying the role of the flow rate in the mineralization process and in the flowing.

Hydrodynamic study of the pilot

It is one of the first time that the impact of the hydrodynamic have been considered in an electrochemical reactor performing EF and AO. The first step in this part was to model the flowing by a Retention Time Distribution (RTD) study. Whatever the experimental conditions the model of several perfectly mixed tank reactor in series with a dead volume mimicked the experimental RTD curves. Two operational parameters were analyzed with this model: the configuration of the electrodes and the flow rate. For all the conditions tested, the dead volume was less than 20 % which shows that the mixing inside the reactor was good and it was even better for the 28-electrode configuration. Concerning the flow rate, the lower it was the more the reactor tends to a low number of perfectly mixed tank reactors the higher is the mean retention time which enables a longer contact time between the organics and the hydroxyl radicals.

Model of the mineralization

A new model was developed by combining the hydrodynamic and the pseudo first order kinetic for the mineralization of the solution. The model is able to describe the mineralization for different positions in the reactor at steady state. Thanks to this model it is possible to optimize an "electrochemical plug flow" reactor depending on the treatment objectives: having the highest mineralization rate, performing a combined treatment, reaching a setpoint value, … To improve this study, an overall characterization of the role of the aeration should be performed related to the mineralization process and the hydrodynamic study as it plays a very important role in the production of H2O2 and on the mixing. Estimating the value of constant K (by considering the gas-liquid transfer, the mass transfer, the troublemakers impact, the flow rate, etc.) will be a very important point to consider for further research. Indeed, it will enable to save energy by having Japp close to Jlim which maximizes the MCE. Jlim can also be reached by

Combined treatment

The chapter 4 demonstrated the clear opportunities of a combined treatment for two main reasons: first, because of the incapacity of the µ-organisms to degrade the HCT, used as a model pollutant, and secondly, because of the drastic reduction in operating costs provided by this combined treatment. Indeed, to treat 1 m 3 solution of pollutant with the use of the pilot followed by a biodegradation, the operating costs could be divided by three.

To perform an efficient treatment, it was evidenced that some specific conditions for the electrochemical treatment should be applied: the aim was to degrade the pharmaceutical pollutant, reduce the aromatic intermediates concentration and promote the formation of the biodegradable compounds. The presence in the solution of the active substance and the aromatic intermediates (AIs) disrupt the µ-organisms' activity by lengthening the lag phase and by decreasing the kinetic of degradation, promoting a sequential degradation. For the lag phase a continuous treatment could solve the problem as the µ-organisms would be selected and will thus be able to metabolize the biodegradable compounds, mainly the carboxylic acids (CAs).

For the kinetic of degradation there could be a problem as only the more biodegradable 2-((aminomethyl)amino)-3,5,6trihydroxycyclohexa-2,5-diene-1,4-dione 73, 60, 57, 55, 71, 85, 129, 69, 157, 115,171, 200 G 2-((aminomethyl)amino)-5hydroxycyclohexa-2,5-diene-1,4dione 59, 71, 57, 55, 73, 77, 105, 112, 133 168 H N-(2,4,5-trihydroxy-3,6dioxocyclohexa-1,4-dien-1yl)formamide 71,73,55,153,57,69,60,199,170,99 I

Fig. 15 :Fig. 16 :Fig. 17 :Fig. 18 :Fig. 19 :

 1516171819 Fig. 15: Effect of catalyst (Fe 2+ ) concentration on the normalized Nad concentration decay kinetics at 100 mA constant current. [Nad]0 = 0.1 mM, [Na2SO4]0 = 0.05 M. a: y = 0.22x, R 2 = 0.9943, b: y = 0.18x, R 2 = 0.9945, c: y = 0.16x, R 2 = 0.9978 and d: y = 0.15x, R 2 = 0.9996 134 Fig. 16: Effect of current on degradation kinetics of 0.1 mM HCT in 0.05 M of Na2SO4 solution containing 0.1 mM Fe 2+ (catalyst); a: y = 0.39x, R 2 = 0.994, b: y = 0.29x, R 2 = 0.996, c: y = 0.19x, R 2 = 0.999 and d: y = 0.14x, R 2 =0.996 135 Fig. 17: Effect of current on degradation kinetics of 0.1 mM TMD in 0.05 M of Na2SO4 solution containing 0.2 mM Fe 2+ (catalyst); a: y = 0.7x, R 2 = 0.995, b: y = 0.59x, R 2 = 0.995, c: y = 0.42x, R 2 = 0.998 and d: y = 0.21x, R 2 =0.991 136 Fig. 18: Effect of current on degradation kinetics of 0.1 mM OFL in 0.05 M of Na2SO4 solution containing 0.1 mM Fe 2+ (catalyst); a: y = 0.71x, R 2 = 0.985, b: y = 0.49x, R 2 = 0.956, c: y = 0.32x, R 2 = 0.986 and d: y = 0.25x, R 2 =0.990 136 Fig. 19: Effect of current on degradation kinetics of 0.1 mM Nad in 0.05 M of Na2SO4 solution containing 0.1 mM Fe 2+ (catalyst); a: y = 0.26x, R 2 = 0.994, b: y = 0.19x, R 2 = 0.986, c: y = 0.14x, R 2 = 0.998 and d: y = 0.11x, R 2 =0.997 137

Fig. 21 :Fig. 22 :Fig. 23 :Fig. 24 :Fig. 25 :Fig. 26 :

 212223242526 Fig. 21: Effect of current on decay of normalized concentration of Nad by electro-Fenton process with anode. I (mA): 100 (l), 300 (◆), 500 (■) and 1000 (▲); [Na2SO4] = 50 mM; [Fe 2+ ] = 0.1 mM 139 Fig. 22: Determination of the absolute second order rate constant for oxidation of pharmaceuticals HCT (a), Nad (b), TMD (c) and OFL (d) by hydroxyl radicals following Eq. 27. Experimental conditions: [HCT] = [Nab] = [TMD] = [OFL] = [BA] = 0.1 mM, [Fe 2+ ] = 0.1 mM (0.2 in the case of TMD), [Na2SO4] = 0.05 M, room temperature 142 Fig. 23: Evolution of TOC decay and MCE (insert panel) as a function of time and current during treatment of 0.1 mM of HCT (corresponding to 8 mg L -1 initial TOC) (a), 0.1 mM of TMD (corresponding to 19.2 mg L -1 initial TOC) (b), 0.1 mM of Nad (corresponding to 20.1 mg L -1 initial TOC) (c) and 0.1 mM of OFL (corresponding to 22 mg L -1 initial TOC) (d) by electro-Fenton. Experimental conditions: [HCT] = [Nab] = [TMD] = [OFL] = [BA] = 0.1 mM, [Fe 2+ ] = 0.1 mM (0.2 in the case of TMD), [Na2SO4] = 0.05 M, room temperature 145 Fig. 24: Energy consumption (EC) for mineralization of pharmaceutical solutions: HCT (a), TMD (b), OFL (c) and Nad (d), calculated from Eq. 25, following TOC removal data of the Fig. 23, as a function of electrolysis time for different currents 147 Fig. 25: Effect of current on normalized TOC removal of a 0.1 mM solution of Nad in 50 mM of Na2SO4 in EF process with BDD and Pt anodes in presence of 0.1 mM Fe 2+ as catalyst 148 Fig. 26: Evolution of TOC decay as a function of time and applied current during treatment of a mixture of four pharmaceuticals (HCT + TMD + Nad + OFL) containing 0.025 mM of each by EF process of [Na2SO4] = 50 mM of [Fe 2+ ] = 0.1 mM 149

Fig. 29 :Fig. 30 :Fig. 31 :Fig. 32 :Fig. 33 :Fig. 34 :Fig. 35 :Fig. 36 :Fig 38 :Fig 39 :Fig. 40 :Fig. 41 :Fig. 42 :Fig. 43 :

 2930313233343536383940414243 Fig. 29: Time-course of carboxylic acids during the electrolysis of 0.1 mM TMD at 500 mA and room temperature. [Fe 2+ ] = 0.1 mM, [Na2SO4]0 = 0.05 M Fig. 30: Time-course of carboxylic acids during the electrolysis of 0.1 mM Nad at 500 mA and room temperature. [Fe 2+ ] = 0.1 mM, [Na2SO4]0 = 0.05 M, room temperature Fig. 31: Time-course of carboxylic acids formed during the electrolysis of 0.1 mM OFL at 500 mA constant current. [Fe 2+ ] = 0.1 mM, [Na2SO4]0 = 0.05 M, room temperature Fig. 32: Evolution of inorganic ions during the electrolysis of 0.1 mM HCT at 500 mA under conditions of Fig. 16. In the case of NH4 + and SO4 2-, K2SO4 and NaClO4 were used as supporting electrolyte respectively Fig. 33: Evolution of inorganic ions during the electrolysis of 0.1 mM TMD at 500 mA under conditions of Fig. 17. In the case of NH4 + , K2SO4 was used as supporting electrolyte Fig. 34: Evolution of inorganic ions during the electrolysis of 0.1 mM Nad at 500 mA under conditions of Fig. 19. K2SO4 and NaClO4 were used as supporting electrolyte, for NH4 + and SO4 2-, respectively Fig. 35: Evolution of inorganic ions during the electrolysis of 0.1 mM OFL at 500 mA under conditions of Fig. 18. In the case of NH4 + , K2SO4 was used as supporting electrolyte Fig. 36: Evolution of NH4 + concentration during EF treatment of 1 mM NO3 -solution at different currents using BDD (▲) or Pt anodes (l) Fig 37: A plausible mineralization pathway for HCT by hydroxyl radicals based on the intermediate products identified using HPLC, IC and GC-MC analysis results Fig 38: Degradation pathway for oxidation of TMT during electro-Fenton at 500 mA. [Na2SO4] = 50 mM, [Fe 2+ ] = 0.2 Fig 39: Degradation pathway during the electrolysis at 500 mA of OFL, 50 mM of Na2SO4 as the supporting electrolyte at 0.1 mM iron (II) Fig. 40: Schematic representation of the lab scale pilot in the configuration with 14 electrodes done with the software Solidworks, 1: inlet, 2: outlet, 3: BDD anode, 4: carbon felt cathode and 5 air inlet Fig. 41: Coefficients of the polynomial function (a) and contour plots versus the liquid flow rate (mL min - 1 ) and the current density (mA cm -2 ) obtained from the Doelhert matrix for the mineralization rate of HCT

Fig. 45 :Fig. 46 :Fig. 47 :

 454647 Fig. 45: Number of continuously stirred reactors in series (N) and fraction of active volume (m) plotted against the liquid flow rate, under 0.2 mL min -1 of air with 14-and 28-electrode configurations Fig. 46: Number of continuously stirred reactors in series (N) plotted against the liquid flow rate, under 0.2 L min -1 and 1 L min -1 of air with 14 electrodes Fig. 47: Fraction of active volume plotted against the liquid flow rate, under 0.2 L min -1 and 1 L min -1 of air with 14 electrodes

  Péchaud Y,Oturan N, Oturan MA (2018). A review on efficiency and cost effectiveness of electro-and bio-electro-Fenton processes: application to the treatment of pharmaceutical pollutants in water. Chemical Engineering Journal (in press) doi:10.1016/j.cej.2018.07.179.

Fig. 2 :

 2 Fig. 2: Scheme of the trickle bed reactor. 1: Cell body, 2: Gasket, 3: Ti/PbO2 anode, 4: Gasket ring, 5: Nylon diaphragm, 6: Cathode frame for loading graphite chips and 7: Nickel plate. Reprinted with permission from ref. (Lei et al. 2013). Copyright 2013, Elsevier

Fig. 6 :

 6 Fig. 6: Scheme of the reactor with two parallel rotating cathodes with 1: speed controller, 2: motor, 3: electrolytic cell, 4: rotating graphite felt disk cathode, 5: Pt anode and 6: carbon brush. Reprinted with permission from ref. (Yu et al. 2014). Copyright 2018, American Chemical Society.

(

  0.43 mM of Acid Orange 7 and 0.5 mM of FeSO4, 0.035 M of Na2SO4 (only MR) at a pH of 3 in 500 min: COD abatement was of 75 % (MF) and of 30 % (MR) under 20 and 100 A/m 2 respectively for MF and MR.

  85 mg L -1 of tricyclazole(TCZ), with 0.05 M of Na2SO4, 1.0 mM of Fe 3+ , pH of 3, membrane flux of 103 L m -2 h -1 and current density of 10 Am -2 have a TCZ degradation of around 80% in 30 min, a production of H2O2 of 1586 and 1494 mg m -2 h -1 and a . OH concentration of 85.2 and 88.3 µM after 20 min with and without aeration 0.1 mM of paracetamol, 0.2 mM of Fe 2+ , 0.05 M of Na2SO4 at pH 3, at 18 ºC, pressure of 2 bars and pumped at 3.0 L min -1 was mineralized at a rate of 44% under 100 mA in 8 hours.

  cathodes are directly connected to the power supply which provides a lower operating cost because of lower cell potential. In bipolar connection, only both border electrodes are connected to the current supplier which results in a higher cell voltage but this configuration provides a better depletion of organic pollutants (Martínez-Huitle et al. 2015). The parameter mixing is directly related to the mass transport and mass transfer phenomena. It can become a real problem during the scale-up and can decrease significantly the efficiency of the process. Martínez-Huitle et al. (2015) highlighted this fact, identifying a real lack of mass transport in big-sized mixed tank reactor compared to flow reactor. For example, dos Santos et al. (dos Santos et al. 2014) compared COD removal at a mixed cell and at a flow-pass cell, at 25 °C,

Fig. 8 :

 8 Fig. 8: Apparent kinetic constants for different pharmaceuticals treated by EF process or by a biotreatment

  : the average between the maximum and minimum value for the energy consumption. D: MCE between x% and y%: x and y are the two extremum values during the EF treatment *: best value for the parameter (1), (2), (3): letters to identify different conditions of treatment

  (B) (LedezmaEstrada et al. 2012). Meanwhile for those four experiments (B,G,H,L) the biodegradability was improved. It is only by performing the biological treatment that it will be possible to confirm if this treatment is actually useless. For the others experiments the efficiency of combining both processes should work and for some of them it can be really interesting in order to reduce the operating costs as the EF treatment allows to have a biodegradable solution in a short treatment time (C,D,E,F)[START_REF] Mansour | Biodegradability improvement of sulfamethazine solutions by means of an electro-Fenton process[END_REF][START_REF] Chu | Electrochemical degradation of m-cresol using porous carbon-nanotube-containing cathode and Ti/SnO2-Sb2O5-IrO2 anode: Kinetics, byproducts and biodegradability[END_REF][START_REF] Annabi | Degradation of enoxacin antibiotic by the electro-Fenton process: optimization, biodegradability improvement and degradation mechanism[END_REF][START_REF] Tian | Efficient degradation of rhodamine B using modified graphite felt gas diffusion electrode by electro-Fenton process[END_REF]).

  [START_REF] Coledam | Electrochemical mineralization of norfloxacin using distinct boron-doped diamond anodes in a filter-press reactor, with investigations of toxicity and oxidation by-products[END_REF] studied the degradation of the antibiotic norfloxacin and measured the evolution of toxicity using the bacteria E. Coli showing that the toxicity remained very high during the first two hours of EF treatment before starting to decrease. By identifying reaction intermediates, they attributed the toxicity of the solution to fluoroquinolone, a cyclic compound that remained in the medium during two h before being degraded. This test was also performed on Daphnia magna, a freshwater crustacean, for the degradation of the azo dye orange II. The same scheme appeared, a rise in toxicity after 5 min of treatment and then a decrease of this toxicity until no inhibition after one hour of treatment[START_REF] Lin | Electro-fenton removal of orange II in a divided cell: reaction mechanism, degradation pathway and toxicity evolution[END_REF]. Among all those studies, one assessed the cytotoxicity during the degradation of the herbicide sulfentrazone on sheep corneal epithelial cells and no toxicity was found for both the parent compound and the byproducts[START_REF] Lima Ac De | Electroanalytical studies of sulfentrazone in protic medium, its degradation by the electro-Fenton process, and toxicity assessment using ss-DNA[END_REF]).

Fig. 11 :

 11 Fig. 11: % of inhibition of bioluminescence of V. fischeri bacteria after 15 min for the EF treatment of 0.22 L of 0.1 mM of metoprolol solution in 0.05 M Na2SO4 and 0.1 mM Fe 2+ at pH 3 using BDD anode at 300 mA and room temperature. Reprinted with permission from ref. (Olvera-Vargas et al. 2016a). Copyright Elsevier, 2016.

  et al. 2014; Mansour et al. 2014; Mansour et al. 2015a; Mansour et al. 2015b).The efficiency of the combination of both treatment is demonstrated in terms of TOC or COD removal but hardly ever the biodegradability or the toxicity of the solution resulting is taken into consideration. This point is of major interest as this treatment aims at delivering a clean water in the environment. One study performed toxicity tests on the treated solution(Roshini et al. 2017). Phytotoxicity tests on Vigna radiada seeds and microbial toxicity tests on Pseudomonas aeruginosa and Bacillus sp. were performed in order to demonstrate the efficiency of their combined process for the degradation of textile wastewater. Both tests showed a decrease in toxicity. No such test could be found concerning the pharmaceutical effluent witch is upsetting.

o

  Efficient removal of diuretic hydrochlorothiazide from water by electro-Fenton process using BDD anode: a kinetic and degradation pathway study o Electro-Fenton treatment of the widely used analgesic tramadol using BDD anode: a kinetic, energetic and degradation pathway study o Complete study of the mineralization and degradation of the b-blocker Nadolol by electro-Fenton, comparison of different anodes

  analyses were done by using a Thermo Scientific devise equipped with a TRACE 1300 gas chromatography coupled to an ISQ single quadrupole mass spectrophotometer operating in electron ionization mode at 70 eV. A 1 mM HCT and TMD solutions were electrolyzed during 10 min and 8 min respectively and then the 230 mL aqueous solutions were extracted two times with ethyl acetate and dichloromethane. Then the organic phase was dried over MgSO4, filtered and concentrated to dryness with a rotary evaporator under vacuum. A few drops of ethyl acetate were added into the flask to dilute the residual solid phase and 3 µL of this solution were then injected into the GC-MS analyser through a TG-5Ms 0.25 mm, 30 m 0.25 mm (i.d.), column. The temperature of the injector and detector were 200 and 250°C, respectively. Helium was used as carrier gas at a flow rate of 1.5 mL min -1 . The temperature ramp was as following for the HCT: 40°C (held during 2 min), 10 °C min -1 to 280°C and 4 min at 280 °C and for the TMD: 100 °C during 1 min, 20 °C min -1 to 280 °C and kept at this temperature for 15 min. The mass spectra were identified by using Xcalibur datalibrary software for the aliphatic molecules and by analyzing the different mass spectrum for the aromatic compounds. For the aliphatic molecules, a derivatization procedure was used. 100 µL of the solution containing the solid phase diluted in ethyl acetate was added to 100 µL of BSTFA (N,O-bis-(trimethylsilyl)trifluoroacetamide), the final solution was put at 80°C for 1 h in an incubator.

Fig. 12 :Fig. 13 :

 1213 Fig. 12: Effect of catalyst (Fe 2+ ) concentration on the normalized HCT concentration decay kinetics at 100 mA constant current. [HCT]0 = 0.1 mM, [Na2SO4]0 = 0.05 M. a: y = 0.23x, R 2 = 0.998, b: y = 0.15x, R 2 = 0.993, c: y = 0.14x, R 2 = 0.999 and d: y = 0.13x, R 2 = 0.993.

Fig. 14 :Fig. 15 :

 1415 Fig. 14: Effect of catalyst (Fe 2+ ) concentration on the normalized OFL concentration decay kinetics at 100 mA constant current. [OFL]0 = 0.1 mM, [Na2SO4]0 = 0.05 M. a: y = 0.32x, R 2 = 0.9863, b: y = 0.29x, R 2 = 0.9853, c: y = 0.24x, R 2 = 0.9838 and d: y = 0.20x, R 2 = 0.9884.

   illustrate the results obtained for currents ranged from 50 to 500 mA. For all the currents, the mass transfer is limiting the reaction. These results highlight that high current values promote the oxidative degradation kinetics of pharmaceutical under study due to the increase in the formation rate of the Fenton's reagent through the enhancement of electrochemical reactions (Eqs. 20-22) and therefore the amount of • OH/M( • OH) in the bulk solution (Eq. 19) and on the anode surface (Eq. 22). The production of high amount of • OH/M( • OH) leads to the high degradation rate of the organics[START_REF] Ren | Simultaneous sulfadiazines degradation and disinfection from municipal secondary effluent by a flow-through electro-Fenton process with graphene-modified cathode[END_REF]. The results obtained in this study highlight the high efficiency of the EF treatment of the pharmaceuticals HCT, TMD, Nad and OFL solutions as these solutions were

Fig. 16 :Fig. 17 :Fig. 18 :Fig. 19 :

 16171819 Fig. 16: Effect of current on degradation kinetics of 0.1 mM HCT in 0.05 M of Na2SO4 solution containing 0.1 mM Fe 2+ (catalyst); a: y = 0.39x, R 2 = 0.994, b: y = 0.29x, R 2 = 0.996, c: y = 0.19x, R 2 = 0.999 and d: y = 0.14x, R 2 =0.996.

Fig. 20 :

 20 Fig. 20: Kinetic of degradation during the elimination of 0.1 mM HCT solution under 500 mA, 50 mM of Na2SO4 and 0.1 mM Fe 2+ with and without additional inorganic ions. The additional compounds were

Fig. 21 :

 21 Fig. 21: Effect of current on decay of normalized concentration of Nad by electro-Fenton process with Pt anode. I (mA): 100 (l), 300 (◆), 500 (■) and 1000 (▲); [Na2SO4] = 50 mM; [Fe 2+ ] = 0.1 mM.

Fig. 22 :

 22 Fig. 22: Determination of the absolute second order rate constant for oxidation of pharmaceuticals HCT (a), Nad (b), TMD (c) and OFL (d) by hydroxyl radicals following Eq. 27. Experimental conditions: [HCT] = [Nab] = [TMD] = [OFL] = [BA] = 0.1 mM, [Fe 2+ ] = 0.1 mM (0.2 in the case of TMD), [Na2SO4] = 0.05 M, room temperature.

Fig. 23 :Fig. 24 :

 2324 Fig. 23: Evolution of TOC decay and MCE (insert panel) as a function of time and current during treatment of 0.1 mM of HCT (corresponding to 8 mg L -1 initial TOC) (a), 0.1 mM of TMD (corresponding to 19.2 mg L -1 initial TOC) (b), 0.1 mM of Nad (corresponding to 20.1 mg L -1 initial TOC) (c) and 0.1 mM of OFL (corresponding to 22 mg L -1 initial TOC) (d) by electro-Fenton. Experimental conditions: [HCT] = [Nab] = [TMD] = [OFL] = 0.1 mM, [Fe 2+ ] = 0.1 mM (0.2 mM in the case of TMD), [Na2SO4] = 0.05 M, room temperature.

Fig. 26 :

 26 Fig. 26: Evolution of TOC decay as a function of time and applied current during treatment of a mixture of four pharmaceuticals (HCT + TMD + Nad + OFL) containing 0.025 mM of each by EF process of [Na2SO4] = 50 mM of [Fe 2+ ] = 0.1 mM.

Fig. 27 :

 27 Fig. 27: TOC evolution during the mineralization of 0.1 mM HCT solution at 500 mA in 50 mM of

Fig. 28 :

 28 Fig. 28: Time-course of carboxylic acids during the electrolysis of 0.1 mM HCT at 500 mA constant current. [Fe 2+ ] = 0.1 mM, [Na2SO4]0 = 0.05 M, room temperature.

Fig. 29 :Fig. 30 :

 2930 Fig. 29: Time-course of carboxylic acids during the electrolysis of 0.1 mM TMD at 500 mA and room temperature. [Fe 2+ ] = 0.2 mM, [Na2SO4]0 = 0.05 M.

Fig. 31 :

 31 Fig. 31: Time-course of carboxylic acids formed during the electrolysis of 0.1 mM OFL at 500 mA constant current. [Fe 2+ ] = 0.1 mM, [Na2SO4]0 = 0.05 M, room temperature.

Fig. 32 :Fig. 33 :Fig. 34 :

 323334 Fig. 32: Evolution of inorganic ions during the electrolysis of 0.1 mM HCT at 500 mA under conditions of Fig. 16. In the case of NH4 + and SO4 2-, K2SO4 and NaClO4 were used as supporting electrolyte respectively.

Fig. 34 :Fig. 36 :

 3436 Fig. 34: Evolution of inorganic ions during the electrolysis of 0.1 mM OFL at 500 mA under conditions of Fig. 18. In the case of NH4 + , K2SO4 was used as supporting electrolyte.

  intermediate E can be formed from B, C or D following consecutive attacks of • OH leading to the opening of N-N bond and desulfonation. Subsequent hydroxylation of E leads to the formation of compounds F, G and H. The quinone (G and H) are probably formed through previous formation of hydroquinones (that are not identified in this study probably because of their quick oxidation to quinone forms). Finally the compound I is formed from oxidation of the lateral group of G. These final polyhydroxylated compounds are not stable and undertake oxidative ring opening reactions to form carboxylic acids (compound J to P). The oxidation of these short-chain carboxylic acids by • OH/M( • OH) constitutes the last stage of the mineralization, i.e., transformation to CO2, H2O and ammonium ion. The results obtained for the inorganic ions released are in accordance with the degradation pathway as ammonium ions are released after chlorine and sulfate ions.

Fig. 37 :

 37 Fig.37: A plausible mineralization pathway for HCT by hydroxyl radicals based on the intermediate products identified using HPLC, IC and GC-MC analysis results.

  from A to G or from A to F. In the first pathway, an O-demethylation associated with Ndemethylations can transform TMD (A) into molecule D. These reactions were also reported during the degradation of TMD by ferrate and ozone[START_REF] Zimmermann | Kinetic and mechanistic investigations of the oxidation of tramadol by ferrate and ozone[END_REF]) and by electrochemical oxidation (Lütke[START_REF] Eversloh | Electrochemical oxidation of tramadol in low-salinity reverse osmosis concentrates using boron-doped diamond anodes[END_REF]. The attack of hydroxyl radicals can also produce a dehydration of TMD producing intermediate B following an O-demethylation. This molecule is then degraded in product E by two N-demethylations. The further attack of hydroxyl radicals on the tertiary carbon of molecules E and D leads to the formation of molecule G by releasing ammonium ions[START_REF] Giannakis | Solar photo-Fenton and UV / H 2 O 2 processes against the antidepressant Venlafaxine in urban wastewaters and human urine . Intermediates formation and biodegradability assessment[END_REF][START_REF] Lambropoulou | Degradation of venlafaxine using TiO 2 / UV process : kinetic studies, RSM optimization, identification of transformation products and toxicity evaluation[END_REF]. The second route starts with the N-demethylation of TMD and the hydroxylation of the aromatic cycle giving molecule C which is then followed by an O-demethylation to forms intermediate F.

  groups of the molecule; piperazinyl substituent, fluorine atom, quinolone moiety and to the cycle containing oxygen and azote atoms. Different degradation pathways of the OFL occurring during its oxidation under action of OH are depicted in Fig. 39. For the piperazinyl group, two types of degradation can occur: an S abstraction of H atom leading to the elimination of the methyl group to form intermediate E or a T abstraction where an hydrogen from the cycle is eliminated enabling the formation of molecules C, D and G. These types of degradation were already found during the degradation of levofloxacin by ozonation (Witte et al. 2009), during the photocatalytic treatment of ofloxacin (Vasquez and Hapeshi 2013) and during the degradation of ciprofloxacin by ozonation (Dewitte et al. 2008). Further oxidation by hydroxyl radicals on the piperazinyl substituent can lead to the formation of molecules H and I (Hapeshi

Fig 39 :

 39 Fig 39: Degradation pathway during the electrolysis at 500 mA of OFL, 50 mM of Na2SO4 as the supporting electrolyte at 0.1 mM iron (II).

Fig. 41 :

 41 Fig. 41: Coefficients of the polynomial function (a) and (b) contour plots versus the liquid flow rate (mL min -1 ) and the current density (mA cm -2 ) obtained from the Doelhert matrix for the mineralization rate of HCT.

Fig. 42

 42 Fig. 42 illustrates the contour plots for MCE and EC. In Fig. 42a and Fig. 42b, one isocurve

Fig. 42 :Flow

 42 Fig. 42: Contour plots of the EC in kWh (gTOC) -1 (a) and MCE in % (b) versus the liquid flow rate (mL min -1 ) and the current density (mA cm -2 ).

Fig. 45 :

 45 Fig. 45: Number of continuously stirred reactors in series (N) and fraction of active volume (m) plotted

Fig. 49 :

 49 Fig. 49: Experimental TOC of the solution (Exp) and TOC calculated by the model (Mod) plotted against the number of the CSTR, under 0.2 L min -1 of air with 28 electrodes, 7 A for different flow rates. The circles represent the model, the cross with 3 branches and two branches represent the experimental points at 20.4 and 42.5 mL min -1ş respectively. The + and _ symbols represent flow rates of 85.0 and 127.5 mL min -1 , respectively.

Fig. 52 :

 52 Fig. 52: % of mineralization at steady state during the AO treatment of a 0.1 mM HCT solution in 50 mM Na2SO4, 0.2 L min -1 of air and 42.5 mL min -1 for different currents and with a mean residence time of 36 min in the pilot reactor of Fig. 50.

Fig. 54 :

 54 Fig. 54: Evolution of mineralization of 0.1 mM HCT solution during the combined treatment of 0.01 gVSS L -1 of activated sludge.

Table 20 :

 20 Structure of AIs identified by LC-MS during the treatment of a concentrated solution of HCT (0.42 mM) by AO. Name Structure 6-hydroxy-4H-benzo[e][1,2,4]thiadiazine-7the CAs formed during the AO be mineralized by biological treatment?

Fig. 55 :

 55 Fig. 55: CAs degraded by a) activated sludge inoculum b) biofilters inoculum c) the µ-organismspreviously selected from biofilters. Oxamic acid (¢), oxalic acid (), acetic acid (▲), maleic acid (¿).

  their lag phase.

  the solution of 250 mgC L -1 of CAs with a pre-treated solution containing AIs and a residual HCT concentration of 0.015 mM. Fig. 56 depicts the degradation of maleic acid (Fig 56a), acetic acid (Fig 56b), oxalic acid (Fig 56c), and oxamic acid (Fig 56d).

Fig. 58 :

 58 Fig. 58: Influence of the dilution of the pre-treated solution on the degradation kinetic of the acetic acid.

  good operating conditions at lab scale for an efficient combined treatmentThe pilot reactor is composed of BDD electrodes that are very efficient to eliminate pollutants as they are non-active anodes(Martínez-Huitle et al. 2015;[START_REF] Nidheesh | Environmental applications of borondoped diamond electrodes: 1. Applications in water and wastewater treatment[END_REF]. They enable to form adsorbed hydroxyl radicals M( • OH) on the surface of the anode material. M( • OH) are

Fig. 60 :

 60 Fig. 60: Combined treatment of 230 mL of a 0.2 mM HCT solution. Biodegradation of electro-Fenton pretreated solution at 100 mA during 1 h using Pt anode under following operating conditions: : [Na2SO4] = 50 mM, [Fe 2+ ] = 0.1 mM, pH 3 and room temperature. Biological treatment was performed using activated sludge and 0.01 gVSS L -1 of µ-organisms. The squared point corresponds to the electrochemical treatment and the circle points to the biological mineralization.

  Two experiments were carried out to understand to disrupting effect of the AIs and the HCT during the degradation of four biodegradable CAs: oxalic, oxamic, maleic and acetic acids. The presence in the solution of the HCT and the AI inhibit the µ-organisms' activity by increasing the acclimation time and decreasing the degradation kinetic of CAs under study. A different lag phase and degradation kinetic could be identified for these CAs. The maleic acid was by the acetic and oxalic acids. The oxamic acid required the longest treatment time. Different dilutions of the AIs and HCT solution were realized to measure their inhibiting effect on the mineralization of CAs. This inhibition can be modelled by a Monod inhibition model highlighting that the inhibition depends on the concentration of HCT and AIs

  putting the adequate distance between the electrodes which will optimize the EC and to keep Japp close to Jlim during the entire treatment different currents could be applied to fit with the concentration decrease. The mathematical model will have to consider the aeration but also other chemical reactions such as mediated oxidations to perfectly model the pilot and pave the way for an industrialization of the process. In addition, to evaluate possible disrupting effects, experiments on real wastewaters are needed especially the effect of the water hardness should be tested.This pilot is very reliable as no clogging nor leak were seen during the two years of experimentation. It is particularly interesting to treat effluents with high or moderate concentrations of pollutants and small particles as it increases the mass transfer and the conductivity of the solution and thus the MCE (LedezmaEstrada et al. 2012). For effluents with low concentration of pollutants the use of reactive membranes or micro-reactors(Mousset et al. 2019b) should be used to reduce the mass transfer limitation but required a solution without particles to avoid fouling or clogging. Nevertheless, the use of our pilot is also possible with a pre-concentration step by means of active carbon[START_REF] El Kateb | Electrochemical advanced oxidation processes using novel electrode materials for mineralization and biodegradability enhancement of nanofiltration concentrate of landfill leachates[END_REF], for example, or by improving the mass transfer with the optimization of the mixing or by reducing the loss of energy by reducing the gap between the electrodes. Thus this pilot seems to be a promising solution to treat wastewater effluents. To sum-up: * The scale up of the EF and the AO was shown to be very successful by demonstrating a very high efficiency. * The importance of the flow rate and the configuration were clearly identified thanks to the hydrodynamic study and the DOE. * A model was designed enabling to optimize the operating conditions to reach a given mineralization rate.

Study on the efficiency of a new pilot scale continuous reactor for wastewater treatment by electrochemical advanced oxidation processes: influence of operating conditions and focus on hydrodynamics

	CHAPTER 3. 1. Introduction
	2. Materials and methods
	2.1. Chemicals
	2.2. Electrochemical equipment
	2.3. HCT degradation and mineralization
	1. Introduction 2.4. Mineralization current efficiency (MCE) and energy consumption
	2. Material and methods (EC)
	2.1. Chemicals 2.5. Design of experiments (DOE)
	2.2. Electrochemical equipment 2.6. Retention time distribution (RTD)
	2.3. Analytical methods and procedures 2.7. Mathematical tools
	3. Results and Discussion 3. Results and discussion
	3.1. Influence of the operating parameters on the degradation kinetics 3.1. Influence of operating parameters and configuration on the reactor
	efficiencies 3.1.1. Effect of the concentration of Fe 2+ 3.1.1. Mineralization, EC and MCE
		3.1.2. Effect of the applied current on the degradation kinetics 3.1.2. Influence of the configuration on the mineralization rate and
		EC
	3.2. Influence of the anode material on the degradation kinetics 3.2. Hydrodynamics
	3.3. Determination of the absolute rate constant for oxidation of four 3.2.1. RTD curves
	pharmaceuticals by hydroxyl radicals 3.2.2. Model
	3.4. Mineralization of the pharmaceuticals 3.2.3. Experimental parameters (N and m)
	3.4.1. Effect of the applied current on the mineralization of 3.3. Mineralization and modelling
		pharmaceutical solutions 3.3.1. Kinetic model
		3.4.2. Influence of the anode material on the mineralization 3.3.2. Comparison between the experimental values and the
		efficiency model
	4. Conclusion	3.4.3. Mineralization in conditions closer to real wastewaters
	References
		3.4.4. Formation and Evolution of short chain carboxylic acids
		and mineral ions
	3.5. Proposition mineralization pathway 3.5.1. Mineralization of HCT by hydroxyl radicals CHAPTER 4.
		3.5.2.. Mineralization pathway for oxidation of TMD by hydroxyl
		radicals
		3.5.3. Proposed mineralization pathway for oxidation of OFL by
		hydroxyl radicals
	4. Conclusions
	References

Electrochemical advanced oxidation processes combined with a biological treatment for wastewater treatment: a deep understanding on the influence of operating conditions and global efficiency

	2.2.2. Treatment with electrochemical pilot 2.3. Biological treatment LIST OF TABLES
	2.4. Analytical methods
	2.4.1. HPLC analysis
	2.4.2. TOC analysis
	2.4.3. Liquid chromatography (LC) -mass spectroscopy (MS)
	analysis
	3. Results and discussion
	3.1. Interests of combining EAOPs and a biological treatment
	3.1.1. Inefficiency of a biological treatment alone in the case of
	persistent organics
	3.1.2. AO to mineralize a HCT solution
	3.1.3. Electro-chemical pre-treatment followed by a biological
	post-treatment
	3.2. Evaluation of the biodegradability of specific intermediates formed
	during the AO treatment of HCT
	3.2.1. Can the CAs formed during the AO be mineralized by
	biological treatment?
	3.2.3. Inhibiting effect of the pollutant and of the AIs on the
	biological degradation
	3.3. Application of appropriate operating conditions to optimize the
	combined treatment
	3.3.1. Experiments with BDD anodes
	3.3.1.1. Finding good operating conditions at lab scale for
	an efficient combined treatment
	3.3.1.2. Finding good operating conditions for pilot scale for
	an efficient combined treatment
	3.3.2. Finding good operating conditions with Platine electrode at
	batch scale
	4. Conclusions
	References
	CONCLUSIONS AND PERPECTIVES
	SCIENTIFIC COMMUNICATIONS
	1. Introduction
	2. Material and methods 2.1. Chemicals ANNEXES
	2.2. Electrochemical treatment
	2.2.1. Batch treatment

Table 1 :

 1 Definition of the different toxicities

Table 2 :

 2 Occurrence and typical removal of common pharmaceuticals in WWTP

Table 3 :

 3 Pharmaceuticals selected with their molecule structure

Table 4 :

 4 New cell configurations and the conditions of treatment

Table 5 :

 5 EF applications by continuous reactors

Table 6 :

 6 EF applied to β-blockers

Table 7 :

 7 EF applied to treatment of anti-inflammatories and analgesics

Table 8 :

 8 EF applied to antibiotics and diuretics

Table 9 :

 9 Biodegradability studies with EF and biological treatments

Table 10 :

 10 Combined process applied to pharmaceuticals

Table 11 :

 11 Apparent rate constants (kapp) for Nad decay with BDD and Pt anode calculated based on the first 8 mins

Table 12 :

 12 Experimental variables and Doehlert matrix

Table 13 :

 13 Extremum values for each parameter

Table 14 :

 14 Mineralization and degradation rate, MCE and EC obtained for different experiments

Table 15 :

 15 Influence of the configuration on the mineralization rate at steady state. 0.1 mM of HCT, 50 mM of Na2SO4 and 0.2 L.min -1 of air were used for all the experiments

Table 16 :

 16 Residence time (τ) and mean residence time (Ts) for different flow rates with 14-and 28-electrode configurations

Table 17 :

 17 Mean residence time Ts, Pe number and the correlation coefficient for different flow rates with 28 electrodes

Table 18 :

 18 Correlation coefficients for the two configurations at different flow rates for the CSTR in series model

Table 19 :

 19 IOA for different flow rates

Table 20 :

 20 Structure of AIs identified by LC-MS during the treatment of a concentrated solution of HCT (0.42 mM) by AO

Table 21 :

 21 Structure and concentration of CAs used in the study in biodegradation experiments

Table 22 :

 22 Results for the different experiments

Table 1 :

 1 Definition of the different toxicities

	Abbrev.	Complete name	Definition
			The concentration of a pollutant of the surrounding medium
	LC50	Half Lethal Concentration	that will kill half of the sample population of a specific test
			species in a specified period of exposure
			The concentration of a drug at which 50% of its maximum
	EC50	Half Efficient Concentration
			response is observed
	NOEC No Observed Effect Concentration	The highest concentration in which no effect is observed
		Low Observed Effect
	LOEC

Concentration

The lowest concentration for which an effect is observed Finally, another parameter called PNEC (Predicted No Effect Concentration) was introduced by the European guidelines (

Table 4 :

 4 New cell configurations and the conditions of treatment

	Name of the reactor	Volume	Distance between electrodes	Number of electrodes	Nature of the electrodes	Type of out streaming	Stirring	Active aeration	Efficiency	References
	Trickle bed reactor Fig. 2	Active dimensions of the electrode: 42 cm 3		2 cathodes 1 anode	-Cathode frame for loading graphite chips -Ni cathode plate -Ti/PbO2 anode	Recirculation flow or continuous mode	By the flow and aeration	Air flow rate of 0.1 m3 h -1	Solution of X-3B 123 mg L -1 mineralized at 87% in 3 h at 4,5 V (10 mA cm -2 )	(Lei et al. 2013)
	Plate electrodes Fig. 1	5L (reactor)	1.5 cm	3 cathodes 2 anodes	-Cathodes: Ti coated with stainless steel -Anodes: Ti coated with IrO2/RuO2	Batch mode	Yes with a stirrer	No	At 58 mM of H2O2, 1.07 mM of Fe 2+ , 22.5 respectively after 1 h mM of aniline initial concentration the percentage of removal of aniline were 87.2 and 77.2 % for plate and rod electrodes	(Su et al. 2013)
	Flow through single system Fig. 4	150 mL (reactor)	0.8 cm	1 anode 1 cathode	Cathode: modified graphite felt Anode: perforated DSA	Continuous	By the flow and aeration	Pumped air at 50 mL min -1

Table 5 :

 5 EF applications by continuous reactors

	EF system reactor configuration	Pollutants	Solution	Current / Voltage	Residence time	TOC removal	Pollutant degradation rate	COD removal	Energy consumption	Ref
	Cylindrical plexiglas reactor, two
	electrodes placed at the center of the reactor separated by 10 mm Anode: graphite Cathode: modified graphite electrode using multi-walled carbon	C.I. Acid Red 14 (a) C.I. Acid Blue 92 (b)	0.05 mM Fe 3+, flow rate: 0.33 L h -1 *, pH: 3*, effective volume 1 L, aeration, 10 mg L -1 dye, NaCl: 1 g L -1	0.14 A			91.22 % (a) 93.45% (b)	86.78% in 60 min	0.024 kWh (gCOD) -1 at 1 h	(Pajootan et al. 2014)
	nanotubes with cationic surfactant
	System composed of 3 pairs of
	anode/cathode disc electrodes of 62
	mm diameter
	Solution flows through the different porous electrodes Anodes: Carbon felt activated by ethanol treatment by ethanol then by mixed valence Cathodes: Carbon felt treated first	Diclofenac	Effective volume 49 mL, addition of O2 flow rate: 10 23°C, tap water, ~0.625 mg L -1 diclofenac, neutral pH, no	2 V* for each pair of cathode/anode	14.6 s	35%	85%			(Plakas et al. 2016)
	iron FeCl2 4 H20 and FeCl3 6 H2O
	(0.4 mg Fe m -2 of cathode)
	Cell with two electrodes of 11 cm 2 separated by 6 cm Anode: BBD Cathode: graphite sheet	Imadacloprid	Air bubbling at 1 L min -1 , = 2* working volume 0.15 L, 100 mg L -1 4.27g* Fe alginate gel beads (0.05 M Fe 3+ ), pH	5 V	4 h		90%			(Iglesias et al. 2014)

*: Best value for the parameter concerned a: Vvm, volume of air under standard conditions per volume of liquid per minute

Table 6 :

 6 EF applied to β-blockers

	Products	Electrodes Anode,	Pollutant initial	Solution	Current	Pollutant	Degradation	Pseudo first	MCE,	Ref
		Cathode, surface and O2 feeding	concentration		Voltage	removing	efficiency	order kinetic constant (pollutant	Energy consumption
								degradation)
	Atenolol	BDD/GDE (1)	158 mg L -1	Batch, 0.05 M	50 mA (1)	Total	in 360 min:	0.39 10 -3 s -1 (1)	Between: 18	(Isarain-
		Pt/GDE -Pt/Carbon		NA2SO4, pH = 3,	50 -12 mA (2)	removal	77% (1)	2.03 10 -3 s -1 (2)	and 50% D (1)	Chávez
		felt (2)		Vsol = 100 mL,		in: > 60	81% (2)	1.56 10 -3 s -1 (3)	22 and 35% D	et al.
		BDD/GDE-		35°C, magnetic bar		min (1)	90% (3)		(2)	2010a)
		BDD/Carbon felt		for stirring (800	50 -12 mA (3)	25 min (2)
		(3) 3 cm 2 for each		rpm), 0.5 mM Fe 2+		30 min (3)			18 and 50% D (3)
		electrode
		GDE fed with air at
		20 mL/min
	Atenolol	Platinized titanium	45 mg L -1	Batch, 0.05* M	0.3* A	-	in 240 min,	1.58 10 -4 s -1	After 4 h:	(El-
		(3*15 cm 2 )		NA2SO4, pH = 3*,			COD	(for COD	22.33% D	Hanafi
		Graphite* (6 cm 2 )		Vsol = 250 mL,			removal:	removal)	0.194	et al.
		With bubbling		20°C, magnetic bar			87%		kWh (kgCOD) -1	2014)
		compressed air at		for stirring (400
		1L/min		rpm), 5* mM Fe 2+
	Atenolol	Pt mesh (4.5 cm 2 )	40 mg L -1	Batch, 0.05 M	300 mA	Total	in 360 min	9.5 10 -3 s -1	Between:	(Sirés
		Carbon felt (14*5 cm 2 ) With bubbling		NA2SO4, pH = 3, Vsol = 220 mL, room temperature,		removal in 6 min	97%		3 and 15% D 1.08 kWh (gTOC) -1	et al. 2010)
		compressed air at		magnetic bar for
	Atenolol	1L min -1	40 mg L -1	stirring, 0.2 mM		-	in 480 min	-	0.42 kWh
	Metoprolol		52 mg L -1	Fe 3+			100%		(gTOC) -1
	Propranolol		52 mg L -1

Table 7 :

 7 EF applied to treatment of anti-inflammatories and analgesics

										Pseudo first
	Products	Electrodes Anode, Cathode, surface and O2 feeding	Pollutant initial concentration	Solution		Current Voltage	Pollutant degradation	TOC removal	order kinetic constant (pollutant	MCE, Energy consumption	Ref
										degradation)
	Diclofenac	BDD onto Nb (64 cm 2 ) Stainless steel (64 cm 2 ) with compressed air bubbled separated by 2 cm	50 mg L -1 100 mg L -1	Batch recirculation, 0.05 M Na2SO4, pH=3, Vsol = 4L, room temperature, 0.5 mM Fe 2+, flow rate 2 L min -1	1.56 mA cm -2 6.25 mA cm -2 1.56 mA cm -2 6.25 mA cm -2	-	in 180 min: 62% 84% 69% 79%	-	0.27 kWh (gTOC) -1 0.78 kWh (gTOC) -1 0.41 kWh (gTOC) -1 0.68 kWh (gTOC) -1	(Garcí a-Monto ya et al. 2015)
		DSA-Cl2 20 cm 2		Batch recirculation,
		GDE containing		0.1 M H2SO4 0.1 M
	Dipyrone	CoPc and PTFE (20 cm 2 ) fed with O2 pseudo reference electrode	50 mg L -1	K2SO4, pH = 3, Vsol = 1.5 L, 20°C, 1 mM Fe 2+, flow rate 50 L h -1	-2.1 V	90 % removal in 30 min	in 90 min 62.8 %	1.2 10 -3 s -1	0.049 kWh (gTOC) -1	(Barro s et al. 2014a)
		Pt//Ag/AgCl/KCl
		DSA-Cl2 20 cm 2		Batch,	0.1	M
		GDE containing		H2SO4	0.1	M
	Dipyrone	CoPc and PTFE (20 cm 2 ) fed with O2 pseudo reference electrode	50 mg L -1	K2SO4, pH = 3, Vsol = 400 mL, 1 mM Fe 2+	-0.9* V	96 % removal in 90 min	in 90 min 54.4%	6.7 10 -3 s -1	0.27 kWh (gTOC) -1	(Barro s et al. 2014b)
		Pt//Ag/AgCl/KCl
	Ibuprofen	BDD onto Si (3cm 2 ) Carbon-PTFE (3cm 2 ) fed with 20mL min -1 of O2	41 mg L -1	Batch, 0.05 M NA2SO4, pH = 3, Vsol = 100 mL, 25°C, magnetic bar for stirring, 0.5 mM	33.3 mA cm -2	Total removal of Ibuprofen 30 min	81 % (DOC removal) in 6h	4.9 10 -4 s -1	Between 5 and 8% D	(Skou mal et al. 2009)
				Fe 2+

*: best value for the parameter (average): the average between the maximum and minimum value for the energy consumption. D: MCE between x% and y%: x and y are the two extremum values during the EF treatment

Table 8 :

 8 EF applied to antibiotics and diuretics

	Products	Electrodes Anode,	Pollutant	Solution	Current	Pollutant	TOC removal	Pseudo first	MCE, Energy	Ref
		Cathode, surface	initial		Voltage	degradation		order kinetic	consumption
		and O2 feeding	concentration					constant
								(pollutant
								degradation)
	Amoxicillin	BDD (4.5cm 2 )	48 mg L -1	Batch, 0.05 M NA2SO4,	60 mA	Total removal	In 10 h: 80 %	2.7 10 -3 s -1	Between 10 and	(Panizza et
		Carbon felt (77 cm 2 )		pH = 3, Vsol = 250 mL,		in 20 min			30 % D	al. 2014)
		With compressed air		magnetic bar for stirring,
				0.2 mM Fe 2+	300 mA	10 min	> 90 %	1.0 10 -3 s -1	3 and 11 % D
	Amoxicillin	Ti4O7 (4 * 6 cm 2 )	20 mg L -1	Batch, Vsol = 230 mL,	120 mA	Total removal	In 6h 90 %		Between 10 and	(Oturan et
		Carbon-felt (14 * 5 cm 2 )		0.05 M NA2SO4, pH = 3, 23 °C, magnetic bar for		in 5 min		10.3 10 -3 s -1	0.2 and 0.7 25 % D	al. 2017)
		Bubbling of		stirring, 0.1 mM Fe 2+					kWh gTOC -1
		compressed air at 1
		L min -1
	Cefalexin	RuO2/Ti mesh 54	200 mg L -1 *	Batch, 0.05 M NA2SO4,	6.66* mA cm -	Total removal	In 8 h 60%	-	Between 7.5	(Ledezma
		cm 2		pH = 3*, Vsol = 500	2	in 4.5 h			and 22% D	Estrada et
		Activated carbon felt		mL, room temperature,						al. 2012)
		(54 cm 2 ) with pure		magnetic bar for
		oxygen at 0.1L min -1		stirring, 1 mM* Fe 2+
	Chlorophene	BDD onto Si (3 cm 2 )	84 mg L -1	Batch, 0.05 M NA2SO4,	60 mA		In 6 h 67 % (1)	-	Between 7 and
		Carbon-PTFE (3 cm 2 ) fed with 20mL		20°C, magnetic bar for pH = 3, Vsol = 200 mL,					27% (1) D
		min -1 of O2 (1) BDD		stirring, 4 mM Fe 3+ (1)
		onto Si (3 cm 2 )		or 0.2 mM Fe 3+ (2)	300 mA	Total removal	(1) 25% (2) 100%
		Carbon felt (70 cm 2 )				in 90 min (1)
		bubbling compressed
		air 1 L min -1 (2)

Table 9 :

 9 Biodegradability studies with EF and biological treatments

	Raw Solution	Biological treatment	Treatment EF	Order of treatment	Efficiency of EF	Efficiency of biological treatment	Effect of both treatment	Evolution of toxicity	Evolution of biodegradability	Ref
	(A): Lurgi coal gasification wastewater collected after a treatment plant with COD: 140-190 mg L -1 , BOD5/COD: 0.05-0.09, TOC: 90-125 mg L -1	Sludge deserved activated carbon from sewage	Undivided 1 L cell with Ti/SnO2 as anode and Active carbon fiber as cathode (20 cm 2 ) for 1h at 10 mA cm -2 , raw pH	B-EF-B	TOC removal (1h) 42%, (1h): 50 % COD removal	COD 8 hours removal: 45.82%, 46.28% in removal: Color	COD 84% removal: removal: 82.72%, Color TOC 95.26%, removal:	-	0.45 after 1h	(Hou et al. 2015)
									At 2h
									BOD5/COD: 0.05
	(B): Synthetic solution of cefelaxin at 200 mg L -1 (COD= 325 mg L -1 ), solution non-biodegradable	-	Undivided cell with as anode: RuO2/Ti (54 cm 2 ) as cathode: activated carbon felt (54 cm 2 ), 500 mL of Na2SO4 at 0.05 M, pH: 3*, 1 mM* Fe 2+ , 6.66* 1 , 8h* mA cm -2 , O2 at 0,1 L min -	EF	COD removal (8h): 72%, Total removal of Cefelaxin (4 h)	-	-	-	and BOD14/COD: 0.08 at 4.5 h: BOD5/COD: 0.71 0.43 at 24 h: and BOD14/COD: BOD5/COD: 0.1 and BOD14/COD: 0.2 at 8 h: BOD5/COD: 0.26	(Ledezma Estrada et al. 2012)
									and BOD14/COD:
									1.0

Table 10 :

 10 Combined process applied to pharmaceuticals

	Raw Solution	Biological treatment	Treatment EF	Order of	Efficiency of EF	Efficiency	Effect of both	Evolution of	Refs
				treatment	alone	of	treatment	biodegradability
						biological
						treatment
						alone
	100 mg/L tetracycline BOD5/COD: 0.02, COD: 128 mgO2/L	Activated sludge at 25°C, during 21 to 25 days in 500 mL with a cellulose cap to ensure the oxygenation by batch	Undivided Cell 1 L with a carbon felt cathode (194 cm 2 ) and a cylindrical platinum anode (5 cm * 1 cm) under agitation and O2 flow, 800 mL of solution containing 0.1* mmol Fe 2+ pH: 3*, 0.05* mol/L Na2SO4, at 300* mA current	EF-B	COD removal at 2 h: 66% 4 h: 86% 6 h: 93% TOC removal at 2 h: 46% 4 h: 72%	TOC removal after 25 days: 10%	TOC removal after 2 h of EF: 69%, after 4 h of EF: 86%	BOD5/COD at 2h: 0.33, at 4h 0.44, at 6h 0.56	(Ferrag-Siagh et al. 2013)
			Divided Cell of 1 L with a carbon felt
	100 mg/L Tylosin BOD5/COD: 0, COD: 136 mgO2/L AOS: 0.52	Activated sludge under aerobic conditions at 25°C, during 25 days in 500 mL by batch	cathode (194 cm 2 ) and a cylindrical platinum anode (5 cm * 1 cm) under agitation and O2 flow, 800 mL of solution containing 0.1* mM Fe 2+ pH: 3*, 0.05* mol/L Na2SO4, at 300* mA current	EF-B	TOC removal after 2h: 45%, 4h: 62%, 6h: 88%	TOC removal 0%	TOC removal 88% after 2 h of EF: 77%, after 4 h of EF:	BOD5/COD at 2h: 4h: 1.95 0.3, at 4h 0.5 AOS at 2h: 1.87	(Ferrag-Siagh et al. 2014)
	0.2 mM of Sulfamethazine	Fresh activated sludge at 25°C under continuous stirring by batch	Undivided Cell 1 L with a carbon felt cathode (208 cm 2 ) and a cylindrical platinum anode (5 cm * 2 cm) under agitation and O2 flow rate of 450 mL/min, the electrolyte containing: 0.5 mM Fe 2+ , pH: 3, 0.05 M Na2SO4, at 500 mA and 18°C	EF-B	TOC removal after 6h: 6.5% (with 0.1 mM pollutant 0.1 mM mA) Fe 2+ and at 200	TOC removal 0%	TOC removal 93.9% after 4 h: after 0.5 h of EF: 61.4%, after 1 h of EF: 78.8%	BOD5/COD at 0.5h: <0.4, at 1h 0.5	(Mansour et al. 2014)
	Wastewater from a pharmaceutical industry from Tunisia, 0.2 g/L Sulfamethazine, TOC: 188.7 mg/L, BOD5/COD: 0.17	Fresh activated sludge at 25°C under continuous stirring by batch	Undivided Cell 1 L with a carbon felt mA and 18°C* cathode (208 cm 2 ) and a cylindrical platinum anode (5 cm * 2 cm) under agitation and O2 flow rate of 450 mL/min, the electrolyte containing: 0.5* mM Fe 2+ , pH: 3, 0.05 M Na2SO4, at 500*	EF-B	TOC removal after 100 min: 7.5%	-	TOC removal after 100 min EF: 81.4%	BOD5/COD after 100 min: 0.35	(Mansour et al. 2015b)

  Table 11 gives the value of the apparent rate

	1
	0.8
	0.6
	C/C 0
	0.4
	0.2
	0
	0	10	20	30
		Time (min) Without additional ions
		With additional ions

Table 11 :

 11 Apparent rate constants (kapp) for Nad decay with BDD and Pt anode calculated based on the first 8 mins.

	Anode Current (A) kapp (min -1 ) Correlation coefficient
		50	0.11	0.995
		100	0.15	0.998
	BDD	300	0.19	0.990
		500	0.26	0.995
		1000	0.3	0.988
		100	0.15	0.994
		300	0.19	0.994
	Pt
		500	0.26	0.962
		1000	0.28	0.980

Table 12 :

 12 Experimental variables and Doehlert matrix

		Experimental variables	Doehlert Matrix
		U1	U2	U3	X1	X2	X3
	1 33.3	95.2	0.13	1	0	0
	2	3.3	95.2	0.13	-1	0	0
	3 25.8 170.0 0.13 0.5 0.866	0
	4 10.8	20.4	0.13 -0.5 -0.866	0
	5 25.8	20.4	0.13 0.5 -0.866	0
	6 10.8 170.0 0.13 -0.5 0.866	0
	7 25.8 120.7 0.25 0.5 0.287 0.816
	8 10.8	69.7	0	-0.5 -0.287 -0.816
	9 25.8	69.7	0	0.5 -0.287 -0.816
	10 18.4 144.5	0	0	0.577 -0.816
	11 10.8 120.7 0.25 -0.5 0.287 0.816
	12 18.4	45.9	0.25	0 -0.577 0.816
	13 18.4	95.2	0.13	0	0	0

Table 13 :

 13 Extremum values for each parameter

	Variables	Maximum Ui	Minimum Ui	Maximum Xi Minimum Xi
	Current (A) 1	33.3	3.3	1	-1
	Flow rate (mL min -1 ) 2	170	20.4	0.866	-0.866
	Fe 2+
	concentration	0.25	0	0.816	-0.816
	(mM) 3

Table 14 :

 14 Mineralization and degradation rate, MCE and EC obtained for different

			experiments
	Experiment	Mineralization rate (%)	Degradation rate (%)	MCE (%)	EC (kWh (g TOC) -1 )
	1	25.5	100	2.06	6.13
	2	5.2	23	4.19	1.01
	3	9.7	67	1.79	6.08
	4	45.1	97	2.54	2.86
	5	82	100	2.33	4.56
	6	7.3	30	2.58	2.71
	7	17.1	100	1.85	6.05
	8	13.9	54	2.85	2.66
	9	27.8	100	1.97	5.98
	10	7.4	68	1.55	6.16
	11	11.7	56	2.84	2.59
	12	37.8	100	2.20	4.30
	13	14.8	91.5	1.70	5.40

Table 16 :

 16 Residence time (τ) and mean residence time (Ts) for different flow rates with 14-and 28-electrode configurations.

	Q (mL min -1 )	20.4	42.5	85.0	127.5	170.0
	τ (min)	83	40	20	13	10
	fp (28-elect) (min)	78	38	19	12	9
	fp (14-elect) (min)	66	34	18	11	-

Table 18 :

 18 Correlation coefficients for the two configurations at different flow rates for the CSTR in series model.

	Flow rate (mL.min -1 )	20.4	42.5	85	127	170
	14 electrodes	0.9928	0.9988	0.9995	0.9995	-
	28 electrodes	0.9995	0.9994	0.9994	0.9992	0.9988

Table 21 :

 21 Structure and concentration of CAs used in the study in biodegradation experiments.

	Name	Molecule	Concentration (mM)
			O
	Oxalic Acid	HO	OH	6.5
		O
		O
	Oxamic Acid	H 2 N	OH	3.35
			O
		HO	O
			O
	Maleic Acid			0.075
			OH
		O
	Acetic Acid			0.075
			OH

Table 22 :

 22 Results for the different experiments

	Run	1	2	3
	% of mineralization	40	47	60
	Oxalic acid (mM)	0.0037	0.0050	0.011
	Oxamic acid (mM)	0.0015	0.0026	0.0037
	Maleic acid (mM)	0.0012	0.00057	0.00025
	Acetic acid (mM)	0.0010	0.00045	0.00031
	Total CAs (mMC)	0.017	0.018	0.031
	Intermediates below the level	No	No	Yes
	% of HCT eliminated	91	> 99	> 99

Annexes Annex Table 1 :

 Annex1 Influence of t the configuration on the mineralization rate at steady state. 0.1 mM of HCT, 50 mM of Na2SO4 and 0.2 L.min -1 of air were used for all the experiments.

	Annex
	Run	Flow rate (mL min -1 )	Current density (A)	Configuration	Mineralization (%)
			20.4	16.6		14	61
			28.9	16.6		14	48
			42.5	16.6		14	36
			85.0	16.6		14	14
			170.0	16.6		14	9
			42.5	10.0		28	50
	10		42.5	16.2		28	60
	11		42.5	22.4		28	76
	12		42.5	16.2		14	37
			Annex Table 2: K for different flow rates
	Q (mL min -1 ) 20.4	42.5	85.0	127.5
	K (m s -1 )	2.52*10 -5	1.62*10 -5	1.44*10 -5	1.01*10 -5

Table 3 :

 3 Compounds found using the following conditions: 1 mM hydrochlorothiazide, 1 mM Fe, 10 min of electrolysis, extraction with dichloromethane and ethyl acetate.

	Name	Molecule	Fragmentation	Identification letter
	N-(5-chloro-2,4-		73, 313,75,
	disulfamoylphenyl)formamide		132, 55, 69,	B
			189, 206,285
	6-chloro-4H-
	benzo[e][1,2,4]thiadiazine-7-sulfonic acid 1,1-dioxide		294, 199, 296, 201, 279 93, 141, 213,	C
	6-hydroxy-4H-
	benzo[e][1,2,4]thiadiazine-7-		125, 77, 201,
	sulfonamide 1,1-dioxide		277, 157, 175	D
	5-hydrazineylbenzene-1,2,4-triol		77, 141, 170	E
	N-(2,4,5-		109, 151, 69,
	trihydroxyphenyl)formamide		85, 57, 67	F
			169,

Table 4 :

 4 Aliphatic compound found using the following conditions: 1 mM hydrochlorothiazide, 1 mM Fe, 10 min of electrolysis, extraction with dichloromethane and ethyl acetate, derivatization (BSTFA).

	Name	Molecule (without sillilation)		Fragmentation	Probability (%)	Retention (min) time	Letter identification
		O
	Butyric acid				75, 145, 73, 146, 117	72	4.55	L
	HO
		O
	2-hydroxyacetic acid		OH		147, 73, 66, 148, 77	67	6.13	M
		HO
	3,4-dihydroxybutanoic acid	O	OH	OH	73, 147, 233, 189, 231	69	9.55	J
	HO

V* 6 h (a) 12.5 h (b)81% (a) 87% (b)[START_REF] Iglesias | Electro-Fenton decolourisation of dyes in an airlift continuous reactor using iron alginate beads[END_REF]

s -1 aiming at a TOC removal of 63% in 90 min.

and 12% (1) 4 and 22% (2) D (Sirés et al. 2007a)

I•OH•OH •OH •OH •OH •OH •OH •OH •OH •OH

•OH •OH •OH •OH •OH •OH •OH •OH •OH -NH 4 +

For the Pt anode, the increase of the current also results in higher mineralization rate until 300 mA. Higher currents increase the efficiency of side reactions that hinder the generation or favor the consumption of • OH/ Pt( • OH) (Eqs. 11 and 28). This decrease in efficiency happens at lower currents than for the BDD because the homogeneous • OH reacts with the iron (Eq. 26) at the same kinetic rate than homogeneous • OH with carboxylic acids and Pt( • OH) are less efficient than BDD( • OH) in the oxidation process. In other studies (Sopaj et al. 2016), this result was also obtained and they performed the anodic oxidation treatment to compare the role of • OH/ M( • OH). The difference of efficiency is explained by the predominant role of M( • OH) in the mineralization process compared to homogeneous • OH especially for long treatment times (6h) which enables the pollutant to reach the anode surface/diffusion layer. Thus as it was previously explained for the degradation, as BDD( • OH) are more formed and slightly adsorbed onto the anode surface their efficiency is much higher than Pt( • OH).

CHAPTER 3

Study on the efficiency of a new pilot scale continuous reactor for wastewater treatment by electrochemical advanced oxidation processes: influence of operating conditions and focus on hydrodynamics

Electrochemical equipment

The electrochemical reactor made of plexiglass was equipped with two different configurations.

In the first configuration, the reactor is composed of 7 anodes and 7 cathodes separated by 4 cm each other (Fig. 40), in the second one it is composed of 14 anodes and 14 cathodes separated by 2 cm each other. Anodes are made of BDD film on niobium support (from CONDIAS GmbH) of 30 cm 2 and cathodes are made of 3D carbon felt (from Mersen, France) of 30 cm 2 . Two peristatic pumps FH100 were used to control the flow throughout the reactor at flow rates of 20.4, 42.5, 85.0, 127.5 and 170.0 mL min -1 . Air supply was provided by a fine bubble diffuser (Fig. 40) and controlled by an air flowmeter. To assess the influence of the airflow, two airflows were tested (0.2 and 0.6 L min -1 ) without changing the other operating conditions. The results were very similar for oxidative degradation of HCT and for the mineralization of its solution. Thus, it was decided to work with airflow of 0.2 L min -1 in all experiments. The solutions to be degraded contain 0.1 mM of HCT, 50 mM of Na2SO4 and the pH is adjusted to 3 with 1 M of sulfuric acid. A constant current was delivered by a Microlab power supply to the electrodes mounted in parallel.

parameters of interest. The liquid flow rate was chosen as it is an important parameter for continuous process, since it governs the hydraulic retention time in the reactor. The current and the catalyst (Fe 2+ ) concentration were added as they are key parameters in electrochemical processes. The Doehlert matrix for three parameters consists of 13 experiments. For each parameter (Ui), different values were tested according to the Doehlert matrix (Table 12)

between a maximum and a minimum level chosen according to lab scale studies. for α1, α2 and α3 respectively. Y = bo + b1X1 + b2X2 + b3X3 + b11X11 2 + b22X22 2 + b33X33 2 + b12X1X2 + b13X1X3 + b23X2X3 (40) with bi and bii the coefficients of the model. The software Nemrod® was used to estimate the different coefficients.

Retention time distribution (RTD)

RTD experiments were performed by the tracer method [START_REF] Viva | The influence of modular structure on the hydrodynamics of catalytic structured packings for reactive separation processes[END_REF]. When the reactor was entirely filled and the flow stable, a salt tracer (6 mL of 233 g L -1 of KCl) was injected at the entrance of the reactor by a syringe of 10 mL. The conductivity of the tracer leaving the reactor was measured. After each acquisition the reactor was emptied and ultrapure water was used to clean it.

The theoretical residence time (τ) is defined thanks to the volume of the reactor and the flow rate (Eq. 41) and represents the theoretical mean time to flow through the reactor.

with τ in min, V the volume of the reactor (mL) and Q the flow rate (mL min -1 ).

To model the RTD functions a dimensionless time is convenient; it is defined by the time divided by the theoretical retention time (Eq. 42).

with t, the time (min) and τ the theoretical residence time (min).

The residence time distribution function (E) gives the age distribution frequency of fluid elements leaving the reactor (Eq. 43).

with θ, the dimensionless time, C(θ), the tracer concentration at the exit of the reactor at time θ and C0, the tracer concentration obtained if all the injected mass was diluted into the volume of the reactor (1.7 L). C0 (concentration of the tracer) is equal to 1.1 * 10 -2 M. Thanks to the shape of the RTD curves the reactor modelling can be identified. The mean residence time is then obtained thanks to the following equation (Eq. 44).

fp =

∫ r E(r)sr t ∫ E(r)sr t ^

(44)

Mathematical tools

To compare the model results with experimental measurements the index of agreement (IOA) (Eq. 45) was used (Mousset et al. 2016;Trellu et al. 2016b).

where K is the number of observed values, É m is a numerically simulated value, [ m is the corresponding experimentally observed value, É j is the average of the numerically simulated values.

The correlation coefficient was calculated by Eq. 46.

) * (ÉU -ÉÖ)] / (á(∑ [(ÑU -ÑÖ K mÜ9

)^2] * á∑ [(ÉU -ÉÖ K mÜ9

)^2] ) (46)

with R, the correlation coefficient, xi, an experimental point, xm, the mean of the experimental points, yi, a theoretical point, ym, the mean of the theoretical points and P the number of points.

Results and discussion

Influence of operating parameters and configuration on the reactor efficiencies

between electrodes and the second with 28 electrodes and a distance of 2 cm between electrodes. The mineralization of the solution was tested for both configurations with different currents or with similar currents but different current densities. For the same total applied current, the mineralization is higher for the 28-electrode system (by comparing Run 1 and 2 of the Table 15). This can be explained by the low extent of side reactions at low current densities, thus the applied current is less wasted in the system allowing better mineralization rate. For the same current per cell, the mineralization is much higher with this 28 electrodes system (around 30% of increases by comparing Run 1 and 3). This confirms that the most important parameter is the contact time between the solution and the electrodes which was already highlighted by the importance of the flow rate.

The operating cost (EC) was also considered for both configurations. Fig. 43 depicts the results obtained. As can be seen from this figure, the 28-electrode configuration is more cost effective as for the same current density per branch the EC is always higher for the 14-electrode configuration.

with Pe the Peclet number representing the proportion of convective flow compared to diffusive flow. The plug flow model is known to be well representing a system when the Pe number is over 100. Both Ts and Pe were estimated thanks to Excel solver with the RTD curves. The different parameters were calculated for the 28-electrode configuration. fp, Pe and the correlation coefficient are given in Table 17 for all the flow rates. The Pe numbers grow with higher flow rates, which mean that the flow is being more convective for higher flow rates. Meanwhile the Pe numbers are below 100 meaning that the plug flow model is not relevant, even more as the more accurate flow rates are the lower ones.

The correlation coefficient for the CSTR model (Table 18) are closer to 1 than for the plug flow model (Table 17) which is also in favor of the mixed tank reactors. Thus the CSTR model in series with dead volumes will be used to study the reactor (Annex Fig. 3).

Another parameter, the air flow rate, was tested since it is an important characteristic of the hydrodynamics of the pilot. Two flow rates were chosen: 0.2 and 1 L min -1 and the number of CSTR in series and the fraction of active volume were considered (Fig. 46). The increase of the aeration rate decreases the number of CSTR being of 4, 6, 9 and 11 reactors for 20.4, 42.5, 85 and 127.5 mL min -1 , respectively. The lower number of reactors is due to a better mixing produced by the aeration which makes the system tend toward the CSTR model. Concerning the fraction of active volume (Fig. 47), no real difference was observed between 0.2 L min -1 and 1 L min -1 . Higher air flow rates were tested (5 and 8 L min -1 ) but the introduction of high air flow lead to high gas retention that decreased the liquid volume in the reactor. This volume could not be estimated properly due to the high gas retention in the reactor and therefore these results were not considered. This parameter still need to be optimized since it could significantly affect the production of H2O2. A compromise should be found: a too low air flow rate will lead to a poor efficiency of the electro-Fenton whereas a too high flow rate could lead to high gas retention which is likely to increase the electrical resistance of the solution. account to simplify the model, and iii) the temperature and the volume of the reactor are considered constant, as observed experimentally (Mousset et al. 2019a).

Two different behaviors can be observed depending on whether the reaction is limited by mass transport or charge transfer reaction. The limiting current density (Jlim) has to be calculated according to Eq. 49 and compared to the applied current density (Japp) in order to know in which regime the reactor is operating (Eq. 50).

where n is the number of electrons to mineralize the pollutant (Eq. 38), F is the Faraday constant equal to 96500 C mol -1 , K is an overall mass transfer coefficient (m s -1 ), C is the concentration of carbon (mmol L -1 ), I is the current (A), A is the anode surface (m 2 ) and J the current density in A m -2 .

When Japp < Jlim, the system is controlled by the current and when Jlim > Japp then the system is operating under mass transfer control. The expression of the anodic oxidation rate rao is given by the Eqs. 51 and 52 (Panizza et al. 2001;[START_REF] Michaud | Comportement anodique du diamant synthetique dopé au bore[END_REF][START_REF] Lan | An experimental and modelling study of the electrochemical oxidation of pharmaceuticals using a boron-doped diamond anode[END_REF]Mousset et al. 2019a).

If Japp < Jlim then, rao = KAC 0 /V (51)

where C 0 is the initial concentration of carbon (0.7 mM), V is the volume of the reactors (m 3 ) and rao in mmol L -1 s -1 . For the pilot Japp is equal to 166 A m -2 (0.5/0.003) for each CSTR. As K is gathering all the transfer system it is strongly linked to the flow rate. Thus values of K were estimated using the software Aquasim® (Annex Table 2) for all the different flow rates. These values are in the same order of magnitude (10 -5 ) which is a common value. For all the flow rates, Japp > Jlim, so it can be assumed that the pilot is operating under mass transport control (Eq. 52). Mass balance can be written assuming that only an AO is occurring (Eq. 53). The model was drawn and the value of the concentration for each CSTR was calculated in the steady state. The model was then challenged to the experimental points and the results are given in Fig. 49. The experimental values and the model are following the same trend. In order to validate the model, the IOA was calculated. The different values are given in Table 19.

All the IOA values are higher than 0.984. Thus, the model is able to describe the degradation for different positions in the reactor at steady state. By the use of a numerical model, more complex kinetic models could be implemented. Thanks to this model (3.4.1.) it is possible to optimize an "electrochemical plug flow" reactor having the higher mineralization rate with the smaller mean retention time. As an example more than 90 % of the solution is mineralized with the 20.4 mL min -1 flow rate thus using a lower flow rate (mean residence time) will consume more energy for nearly the same result. For a different approach, when only a certain part of the solution needs to be mineralized (for a combined treatment for instance) then thanks to this model, the flow rate and hydrodynamic can be estimated to minimize the energy consumption.

The operating cost can also be decreased by tempting to apply a different Japp in all the "perfectly mixed reactors" close to Jlim to maximize the MCE. This is possible by controlling the current (Japp) or to fix aeration, flow rates, configurations to adjust it to Jlim. To a larger extent this study can enable to design reactors and optimize their performances.

mineralization for the more accurate mean residence time and the best electrode configurations.

Several applications of this model can be offered: having the best % of mineralization but also in order to perform a combined treatment, to find the best conditions to reach a given % of mineralization (length of the reactor, configuration, flow rate, current) at the lowest cost.

Eventually, to a larger extent, the mathematical model can be implemented in order to consider the mediated oxidation, the role of the aeration, etc. to perfectly model the pilot and control the process paving the way for the scale-up of the treatment.

CHAPTER 4

Electrochemical advanced oxidation processes combined with a biological treatment for wastewater treatment: a deep understanding on the influence of operating conditions and global efficiency.

Analytical methods

Before each analytical measure the solution was filtered through a 0.2 µm filter made of PTFE (polytetrafluoroethylene). When error bars are given on the graphs, they are calculated based on the standard deviation of the different experiments.

HPLC analysis

HCT and its aromatic degradation intermediates were followed and measured by using an uHPLC equipped with a Hypersil Gold column (100 mm * 2.1 mm), 1.9 µ and coupled with a diode array detector (DAD). The temperature of the column was set at 40°C. The injection volumes were of 5 µL. The mobile phase was composed of water/methanol (95:5) both containing 1% of acetic acid at 0.08 mL min -1 . The wavelength of detection was set at 270 nm.

Oxamic, maleic, acetic and oxalic acids were identified and quantified by a Merck Lachrom Liquid chromatograph equipped with a Bio-Rad Aminex 5 µm, 25 cm × 4.6 mm (i.d.) column and coupled with a DAD. The mobile phase was H2SO4 at 4 mM at a flowrate of 0.6 mL min - 1 . The detector was set at 220 nm. A calibration curve was performed for each detected acid.

Experiments were carried out using a total concentration of CAs of 250 mg C L -1 with 6.5, 3.35, 0.075 and 0.075 mM of oxalic, oxamic, maleic and acetic acid respectively (concentrations are proportional to the one of chapter 2).

TOC analysis

The mineralization of the solution was followed by measuring the total organic carbon (TOC) of treated solutions using a Shimadzu VCSH TOC analyzer. 2 drops of 1 M H2SO4 were added to each sample of 5 mL to acidify the solution in order to eliminate mineral carbon such as it is then necessary to mineralize the residual solution with a biological treatment to reach the same global mineralization efficiency.

Fig. 53:

Cost for the AO treatment described in Fig. 52, using 6, 7 and 8 V to mineralize 50, 60 and 70% of the 0.1 mM HCT solution, respectively. Some combined treatments were then performed in order to assess the efficiency of such system by varying the operational conditions. For most of them the biomass was not able to mineralize the electrochemically pre-treated solution. However, for one operational condition some information could be extracted. In this study, a solution of 0.84 mM of HCT (62 mgC L -1 ) was treated by AO in the pilot at a flow rate of 20.4 mL min -1 (with mean residence time of 82 min)

under 24 mA cm -2 constant current density and effluents were collected after 15 (A), 30 (B), 45

(C) and 120 (D) min. A biological batch experiment was then performed with 0.01 gVSS L -1 of activated sludge mixed with 250 mL of the treated solutions. After the electrochemical pretreatment, HCT was totally degraded for all conditions except for the condition A where 27%

of HCT was still remaining in the solution. The results of the combined treatment concerning the total mineralization rate are given in Fig. 54. As an example, oxalic acid and formic acid could not be mineralized by the EF treatment with Pt anode of 2.5 L of a ranitidine solution at 112.6 mg L -1 concentration (Olvera-Vargas et al. 2015c). Based on this example, an EF treatment was applied in order to form the maximum amount of CAs. Again, the concentration of CAs, the AIs and the mineralization rate were followed during the process. After 1 h, all the intermediates were degraded with mineralization rate of only 35% and the total CA concentration of 6.3 mgC L -1 in the solution. The CAs concentration was 3 times higher compared to the AO using BDD anode highlighting the better suitability of the use of a Pt electrode. Fig. 60 describes the results obtained for the combined treatment. In the same way as the experiments with BDD anode, an acclimation phase was observed but it was much shorter. Then, an exponential phase was observed until day 9 followed by a stabilization of the mineralization rate. The biological treatment mineralized 50% of the initial TOC reaching an overall mineralization degree of nearly 85% with a residual TOC of 2.3 mgC L -1 . So, to oxidize selectively the pollutants and to perform an efficient combined treatment, the use of Pt anode seems more appropriate. However, Pt being highly expensive, (even than BDD), a relatively chipper anode should be selected for a scale up at industrial level.

In this frame, DSA or sub-stoichiometric TiO2 anodes can then be proposed for an industrial scale as they are relatively cheaper active anodes. At short term, the combined treatment should be tested with DSA or sub-stoichiometric TiO2 anodes to perform the EF or AO in the pilot reactor before combining with biodegradation.

Scale-up of the experiment

Once the EF at lab scale was well understood and the best operating parameters found in batch, a pilot was designed and built to assess the efficiency of the process at larger scale and in the continuous mode to be closer to real conditions. Different choices were made during construction of the lab scale pilot presented in this thesis work: (i) the configuration and the size of the electrodes, (ii) the aeration supply and (iii) the manner to plug the electrodes weather in series or in parallel.

(i)

We have decided to design a "plug flow reactor" by putting a stack of electrodes weather than bigger sized electrodes as the use of bigger electrodes produces a nonuniform repartition of the current on the electrode surface which reduces the yield.

A reasonable number of electrode-pairs was chosen (7 or 14 anodes) in order to study different parameters influencing the treatment efficiency and energy consumption.

The electrodes were placed alternately (cathode-anode-cathode...) in order to make the flow run between the electrodes to increase the contact time between the molecules and the electrodes.

(ii) The air supply had two objectives: to provide O2 gas to the cathode to form H2O2 and to increase the mixing between two electrodes. A rubber layer with small sized holes was put on the bottom of the reactor and compressed air was blown under the pilot.

This system is used in biological wastewater treatment plants and provides small sized gas bubbles (d < 2 mm) which favor the oxygen mass transfer to the cathode and the mixing.

(iii) Finally the way of plugging the electrodes was considered. When BDD electrodes are used, a too high current can damage them. Using a parallel plugging enables to have a higher amount in each brunch of the circuit which is not possible in series as the current is not divided. Thus the electrodes were plugged in parallel.

The study of the pilot was carried out in three parts, first a DOE (design of experiments) was performed in order to analyze the performance of the reactor and define the more impacting parameters. Then, to understand the flowing in the reactor, an hydrodynamic study was performed. Finally, a model coupling hydrodynamics and mineralization kinetics was developed using the results of the two previous points.

DOE

The DOE had two objectives: first identifying the more impacting operational parameter and then to have an overview of the ability of the reactor to degrade and mineralize the selected pharmaceuticals. The current density, the flow rate and the catalyst (Fe 2+ ) concentration were chosen based on the lab scale results (chapter 2) and on the literature. For the mineralization rate, the flow rate was the more impacting parameter with a coefficient 2.5 and 10 times higher than the coefficient for the current and the catalyst concentration respectively. The Fe 2+ concentration appeared to have nearly any effect. This latter can be explained by the size of the cathode: at batch scale, the cathode covers all the inner surface of the electrochemical cell (Scathode/Vsolution = 0.9) (chapter 2) whereas in the pilot, this ratio is significantly lower (Scathode/Vsolution = 0.25) which implies less production of H2O2 and thus less EF efficiency. A solution could be to add a carbon felt layer on the floor and on the sides of the pilot to produce higher amounts of hydrogen peroxide. Nevertheless there could be a worst repartition of current with this addition of cathode thus, this will need to be tested.

To evaluate the operating cost, the DOE was also performed for the EC. The current was found to be divided by two when the configuration was changed from 14 electrodes to 28 electrodes compounds could be degraded during their residence time in the biological reactor. In order to face this problem a reactor working as a plug flow could be used. Indeed, in the first sections the more biodegradable compounds would be degraded and then in further sections the less biodegradable compounds would be eliminated as being the only carbon sources for the µorganisms.

Finally these operating conditions were tested with the combined treatment at lab scale in batch as the operating parameters are easier to control in order to degrade the pollutant and the AIs and form a maximum of biodegradable compounds (CAs,..). As BDD anodes constituted the pilot, a BDD anode was selected to perform the combined treatment at lab scale in order to find good operating conditions. As BDD anode is a very powerful anode, a low current can be applied to minimize the mineralization to allow the solution to have enough carbon for the µorganisms. The combined treatment was successful as 66% of the solution was mineralized and 38% by the biotreatment.

The same strategy was then applied to the pilot with three conditions tested. The best one was able to form CAs at higher concentrations and degrade more AIs. During this experiment a current of 6.8 A was applied and a flow rate of 42.5 mL min -1 leading to 60% of mineralization on the electrochemical pre-treatment was used. Only three conditions could be tested because of the lack of time, but a deeper analysis would be very interesting such as realizing a DOE using at least the flow rate and the current as parameters and by measuring the amount of CAs, AI, mother pollutant and the TOC value.

In order to form more CAs and to mineralize less the solution, a treatment with a Pt electrode was performed as it is active less powerful anode. With the Pt electrode, 85% of the solution was mineralized including 50% by the biotreatment which implies a drastic reduction in the operating costs. However, Pt electrode is very expensive thus in the aim of industrialization a cheaper active anode should be used and therefore DSA or sub-stoichiometric TiO2 anodes can be proposed. To predict the efficiency of the combined treatment it would be interesting to developpe kinetic models such as the Monod one used to describe the degradation of the acetic acid.

Summary of the perspectives

The perspectives of this thesis work would mostly be concentrated on the decrease of the costs which are at the moment the real problem of the method. The different axes are:

-The use of concentrated solutions in order to have better MCE% and to decrease the wasting of energy. Two options were proposed whether to concentrate diluted solution on activated carbon, for example, or to perform the electrochemical treatment directly at the exit of a pharmaceutical production unit or directly on an industrial organic wastewater.

-The use of cheaper active anodes, such as DSA or sub-stoichiometric TiO2 anodes, at pilot scale and with the combined treatment.

-To deeper the hydrodynamic study by considering the role of aeration, the calculation of the global transfer K to better model the pilot and optimize its performance.

-To identify the effect of gathering the electrodes to reach Jlim and to evaluate the impact of the iron with higher cathode surface.

-To complexify the mathematical model to consider other chemical reactions and the aeration.

-To developpe mathematical models for the biological reactions of the post-treatment.

Scientific communications:

Publications:

o A review: