Boron doped diamond electrodes

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Evaluation of Electrochemical Reactors as a New Way to Environmental Protection, 2014: 59-78 ISBN: 978-81-308-0549-8 Editors: Juan M. Peralta-Hernández, Manuel A. Rodrigo-Rodrigo and

Carlos A. Martínez-Huitle

4. BDD electrochemical reactors

Ignasi Sirés and Enric Brillas Laboratori d’Electroquímica dels Materials i del Medi Ambient

Departament de Química Física, Facultat de Química, Universitat de Barcelona c/ Martí i Franquès 1-11, 08028 Barcelona, Spain

Abstract. The extraordinary performance of BDD electrodes at

laboratory scale for a wide range of applications, particularly the

environmental protection, has fostered in recent years their

integration in electrochemical reactors as high power, large

O2-evolution overpotential anodes. A large variety of both,

commercial and purpose-made systems, have been used for the

decontamination and disinfection of synthetic and real aqueous

solutions. This chapter pays a look to the fundamentals of direct and

mediated electro-oxidation with BDD, and then describes the

influence of key operation parameters on the degradation rate,

current efficiency and energy consumption. Special focus is put on

the remediation of waters containing organic pollutants, since the

vast majority of studies in the literature refer to the degradation of

industrial chemicals, pesticides, pharmaceuticals and dyes.

Introduction

Over the last 15 years, the use of boron-doped diamond (BDD) thin-film

electrodes for the electrochemical treatment (primordially via

electrochemical oxidation (EO)) of organic contaminants contained in waters

Correspondence/Reprint request: Prof. Enric Brillas, Departament de Química Física, Facultat de Química

Universitat de Barcelona, c/ Martí i Franquès 1-11, 08028 Barcelona, Spain. E-mail: [email protected]

Ignasi Sirés & Enric Brillas 60

has received increasing attention. BDD is synthesized by chemical reduction

of CH4(g) with H2(g) in the presence of B2O3(g) at near 825 ºC and deposited onto

a conductive substrate such as Si, Ti, Nb or a carbonaceous material [1-3].

Thin films thus obtained, usually having between 1 and 10 m thickness,

possess technologically relevant characteristics including an inert surface with

low adsorption properties, remarkable corrosion stability even in strongly

acidic media and extremely large O2-evolution overpotential. Thanks to these

properties, they are excellent anodic materials for EO, with much higher

oxidation ability than other common materials like Pt and PbO2 [2,3]. This

chapter is devoted to describe the main applications of BDD electrochemical

reactors. Firstly, the fundamentals of the oxidation process with BDD anodes

are explained. The treatment of polluted waters by different kinds of either

bench-scale or pilot plants is further discussed, giving details of the most

typical parameters for assessing the current efficiency and energy

consumption. Finally, the effects of relevant operating variables on the

performance of EO with a BDD anode are examined.

1. Fundamentals

EO is the most popular electrochemical method for removing organic

pollutants from wastewaters [2]. This technique has been recently applied to

degrade industrial chemicals, pesticides, pharmaceuticals and dyes from

aqueous solutions. It is based on the oxidation of pollutants and their

reaction intermediates in an electrolytic cell by:

(i) Direct anodic oxidation (or direct electron transfer to the anode), which

yields very poor decontamination.

(ii) Chemical reaction with species that are electrogenerated from water

discharge at the anode M, such as physically adsorbed active oxygen

(physisorbed hydroxyl radical M( OH)) or chemisorbed active oxygen

(oxygen in the lattice of a metal oxide anode (MO)). These oxidizing

species can yield total or partial decontamination, respectively.

Two main approaches have been proposed for the abatement of

wastewater pollution by EO based on the indirect or mediated oxidation with

different heterogeneous species formed from water discharge [2,4,5]:

(i) The electrochemical conversion method, in which organics are

selectively converted into biodegradable compounds, usually carboxylic

acids, by chemisorbed active oxygen.

(ii) The electrochemical combustion (or electrochemical incineration)

method, in which organics become totally mineralized, i.e., oxidized to

CO2, water and inorganic ions, by physisorbed OH. This radical is the

BDD electrochemical reactors 61

second strongest oxidant known after fluorine, with a high standard

reduction potential (Eº = 2.80 V/SHE) that ensures its fast reaction with

most organics giving dehydrogenated or hydroxylated derivatives up to

their mineralization.

In both cases, high potential differences are required between the

electrodes of the electrolytic cell to simultaneously oxidize pollutants and

water, thus maintaining the anode activity. In contrast, low potential

differences hinder the O2 evolution and frequently cause the loss of the

anode surface activity due to the adsorption of by-products involved in the

direct anodic oxidation of the initial organics and, consequently, EO cannot

actually be utilized for wastewater treatment. The nature of the anode

material has a strong influence on both the selectivity and efficiency of the

EO process. This was explained by Comninellis [4] through a

comprehensive model for organics destruction in acidic medium including

the competition with the O2-evolution reaction (OER). Its predictions fit

quite well with results obtained with BDD anodes, which present the largest

O2-overpotential known [6].

According to the Comninellis’ model, the anodes can be classified

according to two extreme cases: the so-called active and non-active anodes.

Typical examples are Pt, IrO2 and RuO2 for the former and PbO2, SnO2 and

BDD for the latter. The proposed model assumes that the initial reaction in

both kind of anodes (denoted as M) is the oxidation of water molecules by

reaction (1) to give the physisorbed M( OH):

M + H2O M( OH) + H+ + e (1)

Both the electrochemical and chemical reactivity of M( OH) depend on

the electrode material. The surface of active anodes interacts strongly with

OH and, as a result, a higher oxide or superoxide MO can be formed from

reaction (2). This occurs when higher oxidation states are available for a

metal oxide anode above the standard potential for O2 evolution (Eº = 1.23

V/SHE).

M( OH) MO + H+ + e (2)

The redox couple MO/M acts as a mediator in organics oxidation via

reaction (3), which competes with the side OER via chemical decomposition

of the higher oxide species from reaction (4).

MO + R M + RO (3)

MO M + ½ O2(g) (4)


In contrast, the surface of a non-active anode interacts so weakly with

OH that organics react with M( OH) to yield fully-oxidized reaction

products such as CO2 via the overall reaction (5):

a M( OH) + R M + m CO2 + n H2O + x H+ + y e (5)

where R is an organic compound with m carbon atoms and without any

heteroatom, which needs a = (2m + n) oxygen atoms to be totally

mineralized to CO2. The oxidative reaction (3) with chemisorbed MO is

much more selective than the mineralization reaction (5) with physisorbed

hydroxyl radical. The latter reaction also competes with waste reactions of

M( OH) such as its direct oxidation to O2 by reaction (6) or its indirect

consumption through dimerization to hydrogen peroxide by reaction (7):

M( OH) M + ½ O2 + H+ + e (6)

2 M( OH) 2 M + H2O2 (7)

According to the aforementioned model, it is expected that a non-active

anode acts as an inert substrate and as a sink for the removal of electrons,

only allowing outer-sphere reactions and water oxidation. Thus, hydroxyl

radical produced from water discharge by reaction (1) is subsequently

involved in the oxidation process of organics. Furthermore, the

electrochemical activity (related to the overvoltage for O2 evolution) and

chemical reactivity (related to the rate of organics oxidation) of physisorbed

M( OH) are closely related to the strength of the M- OH interaction. In

general, the weaker the interaction, the faster the chemical reaction of

organics with M( OH). Since BDD thin-films are the best non-active

electrodes verifying this behavior, they have been proposed as the most

suitable anodes for EO [6].

On the other hand, other weaker oxidants such as ozone can be produced

from water oxidation at the anode by reaction (8) with Eº = 1.51 V/SHE [5]:

3 H2O O3 + 6 H+ + 6 e (8)

Although in EO, reactive oxygen species (ROS) such as M( OH), H2O2

and O3 are formed by reaction (1), (7) and (8), respectively, pollutants are

primordially oxidized by the strongest oxidant M( OH). In fact, this radical

has such a short lifetime that only acts while direct current is supplied to the

anode.


When a BDD anode is utilized, weaker oxidizing species like

peroxodisulfate, peroxodicarbonate and peroxodiphosphate ions can be

competitively produced along with ROS from the anodic oxidation of

sulfate, bicarbonate and phosphate ions present in the electrolyte as follows

[1]:

2 HSO4 S2O82

+ 2 H+ + 2 e (9)

2 HCO3 C2O62

+ 2 H+ + 2 e (10)

2 PO43

P2O84

+ 2 e (11)

A different behavior is found when solutions to be treated by EO contain

chloride ions. Active chlorine species (mainly Cl2, HClO and/or ClO )

formed from Cl oxidation at the anode can effectively degrade some

pollutants in competition with ROS. It is known [2,7] that the electrolysis of

Cl aqueous solutions involves its direct oxidation on the anode to yield

soluble chlorine by reaction (12):

2 Cl Cl2(aq) + 2 e (12)

If the local concentration of dissolved chlorine exceeds its solubility,

supersaturation drives the formation of chlorine gas bubbles. When

electrogenerated chlorine diffuses away from the anode, it can react with

chloride ion to form trichloride ion by reaction (13) or it can be

disproportionated through hydrolysis to yield hypochlorous acid and Cl ion

by reaction (14).

Cl2(aq) + Cl Cl3 (13)

Cl2(aq) + H2O HClO + Cl + H+ (14)

In the solution bulk, hypochlorous acid is in equilibrium with

hypochlorite ion (pKa = 7.55) according to reaction (15):

HClO ClO + H+ (15)

Usually, Cl3 ion is formed in very low concentration up to pH near 4,

whereas the predominant species is Cl2(aq) up to pH close to 3, HClO in the

pH range 3-8 and ClO for pH > 8.0. The mediated oxidation of organics

with these species is then expected to be faster in acidic than in alkaline

media, due to the higher standard potential of Cl2(aq) (Eº = 1.36 V/SHE) and

HClO (Eº = 1.49 V/SHE) compared to that of ClO ion (Eº = 0.89 V/SHE).


The concentration of electrogenerated ClO can be limited by its anodic

oxidation to chlorite ion from reaction (16) and consecutive oxidation to

chlorate and perchlorate ions from reactions (17) and (18), respectively

[2,7]:

ClO + H2O ClO2 + 2 H+ + 2 e (16)

ClO2 + H2O ClO3 + 2 H+ + 2 e (17)

ClO3 + H2O ClO4 + 2 H+ + 2 e (18)

The loss of ClO ion is also possible from waste reactions (19)-(21) in

the solution bulk as well as by reduction to Cl ion on the cathode of an

undivided cell from reaction (22):

2 HClO + ClO ClO3 + 2 Cl + 2 H+ (19)

2 ClO 2 Cl + O2 (20)

ClO + H2 Cl + H2O (21)

ClO + H2O + 2 e Cl + 2 OH (22)

The rate of electrode reactions for generation of M( OH) and other ROS

as well as for Cl2(aq), ClO2 , ClO3 and ClO4 production at the BDD anode

depends mainly on current density, Cl concentration, flow rate and

temperature. In contrast, the rate of homogeneous chemical reactions is a

function of the diffusion rate of organic pollutants through the solution,

which is influenced by the flow rate, the concentration of such pollutants,

temperature and pH. In sections below, the role of these variables during the

degradation of organics using BDD electrochemical reactors will be

discussed.

2. Treatment of organics using BDD electrochemical reactors

A large variety of electrochemical plants and reactors has been tested for

the removal of organic pollutants from waters by EO with a BDD anode.

Figure 1 exemplifies some bench-scale systems containing different

electrochemical reactors operating in batch or in continuous mode [8-14].

Most reactors in that figure are undivided flow cells equipped with planar

(rectangular or circular) and monopolar electrodes in parallel-plate

configuration, like reactors (a), (b) and (d). The bipolar trickle tower reactor

shown in Figure 1c contained Raschig rings coated with BDD, thus giving


Figure 1. Sketches of electrochemical plants and cells used in EO. Bench-scale flow

plants with a: (a) One-compartment flow reactor with a turbulence promoter (C)

operating either in batch (A) or in continuous (B) mode (adapted from ref. [8,9]), (b)

one-compartment flow reactor with parallel-plate electrodes working in batch mode

(adapted from ref. [10]) and (c) bipolar trickle tower reactor in batch mode (adapted

from ref. [11,12]). (d) FM01-LC filter-press flow reactor (adapted from ref. [13,14]).

rise to a large total surface of this material. Scialdone et al. [15,16] have

utilized microreactors with a micro-gap between electrodes that are even able

to destroy pollutants in pure water without the addition of any electrolyte.

The high oxidation ability of BDD( OH) formed from water discharge by

reaction (1) on the BDD surface in the above reactors ensures the large

effectiveness of EO, although slower reactions with other ROS (H2O2 and O3)

and weaker electrogenerated oxidants (peroxodisulfate, peroxodicarbonate or

peroxodiphosphate) are also feasible. This behavior has been corroborated for

synthetic solutions of several industrial chemicals, pesticides, dyes and


pharmaceuticals [12,13,15-35], as well as for urban and industrial wastewaters

[9-11,36-44]. The decontamination process of organic pollutants in the

effluents is monitored from the abatement of their chemical oxygen demand

(COD) and/or total organic carbon (TOC). From COD data, for example, the

percentage of COD removal is calculated as follows:

(23)

where COD is the experimental COD decay (in mg L-1

) at electrolysis

time t and COD0 is the corresponding initial value before treatment. A

similar equation can be defined for the percentage of TOC removal.

Table 1 collects the percentage of COD or TOC removal for selected EO

experiments. As can be seen, large percentages of organic matter removal

were found in most cases operating in batch mode during prolonged

electrolysis time. In contrast, a poor mineralization (only 38%) was obtained

for 50 mg L-1

ketoprofen under continuous conditions working at a high

current density of 235 mA cm-2

and a very low liquid flow rate of 1.42 mL

min-1

[28]. The presence of Cl ion in the effluent leads to the additional

production of active chlorine species (Cl2/HClO/ClO ) from reactions (12)-

(15), which can also destroy organics in competence with BDD( OH), other

ROS and weaker oxidants produced from other anions. Although the

degradation is usually enhanced in the presence of Cl ion, several works

have demonstrated the generation of hazardous by-products such as

organochlorides, chloramines, trihalomethanes and inorganic ions like ClO3

and ClO4 [9,10,21,27], even in drinking water [45]. Determination of all

these species during the EO processes with a BDD anode is then necessary

to confirm whether their viability is possible, especially when actual urban

and industrial wastewaters with high Cl contents are treated. Several coupled and integrated processes have been proposed and

applied to the treatment of wastewaters in order to enhance the degradation

power of EO with a BDD anode [46-58]. Coupled processes involve solar

photoelectro-Fenton [46,49,51,53], EO/Microwave [48], EO with reverse

osmosis/nanofiltration [52], EO with generated H2O2/UV-C photolysis [58]

and EO/Microfiltration [57]. Examples of electrochemical plants used for the

two first techniques are schematized in Figure 2. Table 2 shows that ≥ 93%

mineralization was achieved for different synthetic solutions under treatment

with the above coupled processes, which were more effective than the

corresponding EO ones. For example, the use of EO/Microwave for the

degradation of 100 mg L-1

of the pesticide 2,4-D (2,4-dichlorophenoxyacetic

acid) under the conditions given in Table 2 yielded 98% COD removal after

240 min of electrolysis in batch, whereas the comparative EO process only


Table 1. Percentage of COD or TOC removal for selected degradations of organic

pollutants in synthetic and real waters by EO with a BDD anode in an undivided flow

plant.

Pollutant Solution Electrolytic system % COD

removal

Ref.

o-Cresol

p-Cresol

2,4-D a

Ketoprofen

Mecoprop

4-Nitro-

phenol

Phenol

Petro- chemical

wastewater

Textile

wastewater

1 L of 2 mM pollutant, 1 M

H2SO4, room temperature

3 L of 50 mg L-1 pollutant,

0.05 M Na2SO4, pH 3.0,

room temperature

50 mg L-1 pollutant, 0.50 M Na2SO4, pH 3.99, 25 ºC

1.8 L of 178 mg L-1

pollutant, 0.05 M Na2SO4, pH 3.0, 40 ºC

0.5 L of 2300 mg L-1 COD

of pollutant, 5 g L-1

Na2SO4, pH 2, 25 ºC

4.5 L of 30 mM pollutant, 0.1 M Na2SO4, pH 7, 30 ºC

1.5 L of wastewater with

250 mg L-1 COD, 0.61 mS

cm-1, 47.8 mg L-1 Cl , pH

6.9, 25 ºC

1 L of wastewater with 566

mg L-1 COD, 5.74 mS cm-1,

350 mg L-1 Cl , pH 12.5, room temperature

FM01-LC reactor with 64

cm2 electrodes, 180 min in batch, 10 mA cm-2,

Reynolds number 42,631

Flow reactor with 64 cm2

electrodes, 160 min in batch,

31 mA cm-2, flow rate 10 L min-1

Flow reactor with 12.5 cm2

electrodes, in continuous,

235 mA cm-2, flow rate 1.42 mL min-1

Flow reactor with 20 cm2 electrodes, 25 h in batch,

50 mA cm-2, flow rate

230 L h-1


electrodes, 7 h in batch, 30 mA cm-2, flow

rate 150 L h-1


electrodes, 8 h in batch,

20 mA cm-2, flow rate 4 L min-1

Flow reactor with 63.5 cm2 electrodes, 120 min in batch,

5 mA cm-2, flow rate 160 L

h-1

Bipolar trickle tower reactor,

352 cm2 total electrode area, 8 h in batch, 1 mA cm-2,

flow rate 36.3 mL min-1

84

90

70 b

38

100 b

100

97

90 100 b

91

[13]

[34]

[28]

[19]

[18]

[30]

[10]

[11]

a 2,4-Dichlorophenoxyacetic acid. b Percentage of TOC removal.


Figure 2. Schemes of (a) a 2.5 L bench-scale flow plant and (b) one-compartment

filter-press electrochemical reactor with a BDD anode and a gas (O2 or air) diffusion

cathode, both of 20 cm2 area, used for solar photoelectro-Fenton (SPEF) treatment.

Adapted from ref. [46]. (c) Experimental setup for EO degradation with microwave

activation of the BDD anode at frequency of 2450 MHz and output power of 127.5 W

using a one-compartment cylindrical reactor. Adapted from ref. [48].

reduced the COD by 67% [48]. This increase in mineralization degree was

associated to the enhancement of the mass transport of organics toward the

BDD anode, which was induced by the high frequency of the microwave

radiation applied. On the other hand, interesting integrated pilot plants have

been used by the Ortiz group for the decontamination of landfill leachate by

Fenton followed by EO [47] and municipal wastewaters by ultrafiltration/

reverse osmosis followed by EO [55]. For the former integrated method, for

example, 97% COD removal was found starting from a landfill leachate with

4430 mg L-1

COD, as shown in Table 2.


Table 2. Percentage of COD or TOC removal for selected degradations of organic

pollutants in synthetic and real waters by coupled or integrated processes, which

include EO with a BDD anode, using undivided bench-scale or pilot plant.

3. Current efficiency and energy consumption

To assess the effectiveness and viability of the mineralization of organic

pollutants in waters by EO alone or coupled using BDD electrochemical

reactors, the current efficiency and energy consumption are determined from

the COD and/or TOC abatement of the effluent.

Operating in batch mode at constant current, COD data are used to

calculate the instantaneous current efficiency (ICE, in %) as follows [59]:


(24)

where F is the Faraday constant (96,485 C mol-1

), Vs is the solution volume

(in L), COD is the experimental COD decay (in mg L-1

) at the time

interval t (in s) and I is the current (in A).

The Comninellis’ group [60] has proposed a kinetic model to explain the

change in COD and ICE when a pure organic of formula CxHyOz undergoes

electrochemical incineration in a recirculation plant at constant current. This

model is based on the variation of the limiting current density during the

electrolysis, which is determined from Eq. (25):

(25)

where ilim(t) is the limiting current density (in A m-2

) at a given time t, km is

the average mass-transport coefficient in the electrochemical reactor (in m s-1

)

and COD(t) (in mol m-3

) is the chemical oxygen demand at the same time t.

Figures 3a and b show the two regimes (A and B) identified depending

on the ratio = iappl/ilimo, where iappl is the applied current density and ilim

o is

the initial limiting current density:

(i) Regime A (iappl < ilim(t)): the electrolysis is under current control, COD

decreases linearly with time and ICE is 100%.

(ii) Regime B (iappl > ilim(t)): the electrolysis is under mass transport control

and secondary reactions (such as O2 evolution) start. Both COD removal

and ICE follow an exponential trend due to mass-transport limitation.

A good agreement between experimental and predicted values from the

above kinetic model has been found for both COD and ICE evolutions

during the EO with a BDD anode of different organic compounds like acetic

acid, isopropanol, phenol, 4-chlorophenol and 2-naphthol in 1 M H2SO4

[61].

When TOC removal of a solution containing a pure organic is

monitored, the mineralization current efficiency (MCE, in %) at constant I

(in A) and electrolysis time t (in h) can be estimated from the expression

[62]:

(26)

where n is the number of electrons exchanged in the mineralization process

of the organic compound, 4.32×107 is a conversion factor (= 3,600 s h

-1

12,000 mg carbon mol-1

) and m is the number of carbon atoms of the

molecule under study.


Energy parameters are essential figures-of-merit to assess the viability of

the electrochemical treatment for industrial application. Operating at

constant I, the energy consumptions per unit volume (EC) and unit TOC

mass (ECTOC) are obtained from Eqs. (27) and (28), respectively [51]:

(27)

(28)

where Ecell is the average potential difference of the cell (in V). A similar

equation to (28) is applied to determine the energy consumption per unit

COD mass (ECCOD, in kWh (g COD)-1

).

Figure 3. Evolution of (a) COD and (b) instantaneous current efficiency with

electrolysis time. (A) represents the charge transfer control and (B) represents the

mass transport control. Symbols: COD0 = initial chemical oxygen demand (in mol

O2 m-3), = iappl/ilim

o, VR = reservoir volume (in m3), km

= mass-transport coefficient

in the reactor (in m s-1) and A= electrode area (m2). Adapted from ref. [60].

4. Effect of operating parameters

It has been reported that the mineralization rate, current efficiency and

energy consumption for the EO treatment of organic pollutants in a BDD

electrochemical reactor depend on several operating parameters that include

primordially the applied current (or current density), temperature, flow rate,


pollutant concentration, pH and the electrolyte type and its concentration.

For each electrolysis, a careful study of these variables is necessary to find

the best experimental conditions to be applied.

The current density is a key parameter since it regulates the amounts of

ROS and other weaker electrogenerated oxidants that are able to destroy the

pollutants. As an example, Figure 4a shows the influence of current density

on the TOC abatement with the specific charge Q (in A h L-1

) determined for

Figure 4. Influence of operating variables on TOC decay for the degradation of 1.8 L

of 178 mg L-1 mecoprop solutions in 0.05 M Na2SO4 of pH 3.0 by EO using a flow

reactor with a 20 cm2 BDD anode. (a), current density: () 150 mA cm-2, () 100

mA cm-2 and () 50 mA cm-2, temperature 40 ºC, flow rate 130 L h-1. (b), current

density 150 mA cm-2, temperature: () 15 ºC, () 40 ºC and () 60 ºC, flow rate

130 L h-1. Plot (c): Current density 150 mA cm-2, temperature 40 ºC, flow rate: ()

75 L h-1, () 130 L h-1, () 170 L h-1 and (▲) 230 L h-1. Adapted from ref. [19].


the EO degradation of 1.8 L of 178 mg L-1

mecoprop solutions in 0.05 M

Na2SO4 at pH 3.0 and 40 ºC using a plant like that of Figure 2a, but only

equipped with a flow reactor with a 20 cm2 BDD anode, at flow rate of

130 L h-1

[19]. An increase in current density from 50 to 150 mA cm-2

caused a rise in specific charge from 14 to 27 A h L-1

for reaching overall

decontamination, although TOC was more rapidly reduced since the time

needed for total mineralization dropped from 25 to 17 h. The faster TOC

removal with time at higher current density can then be ascribed to the

concomitant greater production of BDD( OH) from reaction (1), thus

accelerating the oxidation rate of all organics. However, the higher specific

charge required for total decontamination indicates a gradually slower

accumulation of BDD( OH) due to the progressive acceleration of its

parasitic non-oxidizing reactions. These wasting reactions involve the

primary oxidation of this radical to O2 by reaction (6) or its reaction with

H2O2 (generated via reaction (7)) to produce hydroperoxyl radical (HO2 ) by

reaction (29) [2,19,34,51]. Moreover, the parallel increase in rate of O3

(reaction (8)) and peroxodisulfate (reaction (9)) formation can also reduce

the accumulation of BDD( OH) at the anode.

(29)

The above behavior is also reflected in the concomitant decay in current

efficiency and the higher energy consumption. At 3 h of electrolysis, for

example, MCE dropped from 23% at 50 mA cm-2

(Q = 1.67 A h L-1

) to 8.9%

at 150 mA cm-2

(Q = 5 A h L-1

). In contrast, the final EC was raised from

149 kWh m-3

at 50 mA cm-2

in 25 h to 700 kWh m-3

at 150 mA cm-2

in 17 h.

The increase in temperature is expected to cause a greater mass transport

toward the BDD anode due to the decrease of medium viscosity. This

phenomenon can explain the tendency found when the temperature of the

above 178 mg L-1

mecoprop solution varied from 15 to 60 ºC operating at

150 mA cm-2

and flow rate of 130 L h-1

. Under these conditions, Figure 4b

shows a more rapid TOC destruction at higher temperature, decreasing the

required specific charge for total decontamination from 35 A h L-1

at 15 ºC

to 20 A h L-1

at 60 ºC. The mecoprop mineralization was then accelerated

because of the faster reaction of pollutants with BDD(●OH) as a result of the

inhibition of its wasting reactions since more amounts of organics could be

transported toward the BDD surface. This mass-transport limitation

undergone by organics to arrive to the anode was confirmed by increasing

the flow rate of the solution, as shown in Figure 4c for the same solution

electrolyzed at 150 mA cm-2

and 40 ºC. A progressively faster TOC decay

was observed when the flow rate was raised, then dropping the specific


charge for total mineralization from 34 A h L-1

(20 h) at 75 L h-1

to 18 A h L-1

(11 h) at 230 L h-1

.

The aforementioned results indicate an enhancement of the EO process

with a BDD anode upon increase of temperature and flow rate, yielding

higher current efficiency and lower energy consumption. The same behavior

is found when the pollutant concentration increases. Figure 5a shows that

total mineralization of the herbicide mecoprop was feasible when its

concentration rose up to 643 mg L-1

(near to saturation), although longer

time or higher Q was required. Figure 5b highlights that the current

efficiency became greater for higher initial herbicide concentration. An

MCE value as high as 100% can be observed at the early stages of the

treatment for the most concentrated solution at 50 mA cm-2

, whereas it was

reduced to 18% for 73 mg L-1

. This trend can be related to the enhanced

mass transport of organics toward the anode when the initial concentration

increases, thus favoring their reaction with BDD( OH) and inhibiting the

parasitic non- oxidizing reactions of this radical. Moreover, Figure 5b shows

Figure 5. Dependence of (a) TOC and (b) mineralization current efficiency on

specific charge for the EO treatment of 1.8 L of mecoprop solutions in 0.10 M

Na2SO4 at pH 3.0, 50 mA cm-2, 40 ºC and flow rate of 230 L h-1 in a flow reactor

with a 20 cm2 BDD anode Initial herbicide concentration: () 643 mg L-1,

() 356 mg L-1, () 178 mg L-1 and (▲) 73 mg L-1. Adapted from ref. [19].


a dramatic decay of the current efficiency when all electrolyses were

prolonged up to achieve total decontamination. This behavior has been

explained by the progressive decrease of organic matter with formation of

more hardly oxidizable by-products [2]. At the end of the degradation of the

643 mg L-1

mecoprop solution after consumption of 12.2 A h L-1

(22 h) by

applying 50 mA cm-2

at 40 ºC and flow rate of 230 L h-1

, MCE = 29% and

EC = 143 kWh m-3

(ECTOC = 0.401 kWh (g TOC)-1

) were obtained.

While high temperature, flow rate and organic concentration and low

current density are desirable for improving the performance of the EO

treatment with a BDD anode, the effect of pH seems related to the effluent

composition. In the presence of Na2SO4, several authors [19,28] have

reported a slight influence of pH in the range 2-12 on the degradation of

organics, even when varying the electrolyte concentration from 0.05 to 0.50 M

[19]. In contrast, Weiss et al. [22] have described that solutions of sodium

dodecylbenzenesulfonate, a common surfactant, treated in a system like that

of Figure 1b exhibited lower TOC removal rates in alkaline media of pH 12

than in acidic or neutral solutions, due to concurrent oxidation of dissolved

carbonates from reaction (10) at potentials less positive than that required for

water oxidation. When Cl ion is present in the medium, the active chlorine

species formed via reactions (12)-(15) also attack the organic matter and the

degradation is enhanced in acidic media compared to the alkaline ones

because of the higher standard potential of Cl2/HClO than of ClO ion,

which predominates in the latter effluent. This behavior has been confirmed,

for example, by Scialdone et al. [27] for the electrochemical incineration of

250 mL of 100 mg L-1

oxalic acid in Na2SO4 (pH 2) or NaOH (pH 12)

solutions using a flow plant and an electrochemical reactor like that of

Figure 1b after injection of 7,000 C at 39 mA cm-2

, 25 ºC and flow rate of

0.2 L min-1

. In the absence of NaCl, 81-83% mineralization with 56-57%

current efficiency was found in both media. In contrast, by adding 10 g L-1

NaCl, the mineralization increased up to 93% with 65% current efficiency at

pH 2, but it was only reduced by 72% with 48% current efficiency at pH 12

as a result of the greater oxidation ability of active chlorine species

generated in the former medium, as stated above.

5. Conclusion

A variety of BDD electrochemical reactors are available nowadays for

water treatment. Robust equipments based on the commercial DiaCell® have

been successfully integrated in swimming pools and spas for ensuring an

effective disinfection. The conventional FM01-LC reactor, as well as

innovative microfluidic cells based on the ElectroCell AB, have also been well


proven for batch or continuous decontamination. However, the preponderant

systems for the EO treatment of organic pollutants involve purpose-made

filter-press reactors, which have shown a high ability for the complete

degradation of pesticides, dyes, pharmaceuticals and common industrial

contaminants contained in synthetic solutions and real wastewaters. BDD,

when used as an anode, is able to produce a very oxidizing mixture, which

depending on the composition of the electrolyte can contain different ROS

( OH, H2O2, O3), active chlorine and peroxosalts. In such media, the organic

pollutants and their reaction by-products can be even completely mineralized.

Furthermore, if current density and flow rate, which tend to become the two

key operating parameters, are accurately chosen, a 100% TOC removal can be

attained along with a high current efficiency and moderate energy

consumption when highly concentrated solutions are treated. The integration

of EO with BDD anodes in existing water treatment facilities, as well as

further development of coupled processes recently proposed can definitely

enhance the economic viability of the BDD technology, finally allowing its

real implementation for treating wastewaters at real scale.

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Engineering

Boron doped diamond electrodes