Electrochemical approaches to environmental problems in ...ecaaser3.ecaa.ntu.edu.tw/weifang/eBook/electrolysis... · oxidation:reduction processes in an electrochemical cell without

Electrochimica Acta 45 (2000) 2575–2594

Electrochemical approaches to environmental problems inthe process industry

K. Juttner a,*, U. Galla b, H. Schmieder b

a Karl-Winnacker-Institut, DECHEMA e.V., Postfach 15 01 04, D-60061 Frankfurt, Germanyb Forschungszentrum Karlsruhe, ITC-CPV, PO Box 3640, 76021 Karlsruhe, Germany

Papers received in Newcastle, 20 December 1999

Abstract

Electrochemical processes can provide valuable contributions to the protection of the environment throughimplementation of effluent treatment and production-integrated processes for the minimisation of waste and toxiccompounds. As examples of effluent treatment, electrochemical reactors for removal of metal ions from waste water,anodic destruction of organic pollutants and new electrochemical abatement techniques for the purification of fluegases will be described. Further examples consider salt splitting by membrane techniques. As examples forproduction-integrated industrial processes fluidised bed electrolysis for metal recovery in the cellulose acetateproduction, the membrane process for industrial chlor-alkali electrolysis, and the electroreduction of dichloraceticacid are considered. © 2000 Elsevier Science Ltd. All rights reserved.

Keywords: Electrochemical processes; Effluent treatment; Organic pollutants

www.elsevier.nl/locate/electacta

1. Introduction

An increasing world population with growing indus-trial demands has lead to the situation where the pro-tection of the environment has become a major issueand crucial factor for the future development of indus-trial processes, which will have to meet the require-ments of sustainable development. Electrochemistryoffers promising approaches for the prevention of pol-lution problems in the process industry. The inherentadvantage is its environmental compatibility, due to thefact that the main reagent, the electron, is a ‘cleanreagent’. The strategies include both the treatment ofeffluents and waste and the development of new pro-cesses or products with less harmful effects, often de-noted as process-integrated environmental protection:� Cathodic and anodic treatment of effluents and waste.

This includes all techniques where toxic material is

removed from gases, liquids or even solids at thefinal stage of an industrial process.

� Process-integrated environmental protection. This in-cludes recycling of valuable material and substitutionof waste-producing processes by a cleaner electro-chemical technology with little or no wasteproduction.Removal and destruction of pollutant species can be

carried out directly or indirectly by electrochemicaloxidation/reduction processes in an electrochemical cellwithout continuous feed of redox chemicals. In addi-tion, the high selectivity of many electrochemical pro-cesses helps to prevent the production of unwantedby-products, which in many cases have to be treated aswaste.

Attractive advantages of electrochemical processesare generally:1. Versatility — direct or indirect oxidation and re-

duction, phase separation, concentration or dilu-tion, biocide functionality, applicability to a varietyof media and pollutants in gases, liquids, and solids,and treatment of small to large volumes from mi-crolitres up to millions of litres.

* Corresponding author.E-mail address: [email protected] (K. Juttner)

0013-4686/00/$ - see front matter © 2000 Elsevier Science Ltd. All rights reserved.

PII: S 0013 -4686 (00 )00339 -X

K. Juttner et al. / Electrochimica Acta 45 (2000) 2575–25942576

2. Energy efficiency — electrochemical processes gen-erally have lower temperature requirements thantheir equivalent non-electrochemical counterparts,e.g. thermal incineration. Electrodes and cells can bedesigned to minimise power losses caused by inho-mogeneous current distribution, voltage drop andside reactions.

3. Amenability to automation — the system inherentvariables of electrochemical processes, e.g. electrodepotential and cell current, are particularly suitablefor facilitating process automation.

4. Cost effectiveness — cell constructions and peri-pheral equipment are generally simple and, if prop-erly designed, also inexpensive.

Despite these advantages, electrochemical processesare of a heterogeneous nature, which means that thereactions are taking place at the interface of an elec-tronic conductor (the electrode) and an ion conductingmedium (the electrolyte). This implies that the perfor-mance of electrochemical processes may suffer fromlimitations of mass transfer and the size of specificelectrode area. Another crucial point is the chemicalstability of the cell components in contact with anaggressive medium and in particular the long termstability and activity of the electrode material.

The application of electrochemistry for the protec-tion of the environment have already been the topic ofseveral books and reviews [1–8]. Besides the process-oriented benefits, electrochemistry is also playing a keyrole in sensor technology. Electroanalytical techniquesfor monitoring and trace level detection of pollutants inair, water and soil as well as of micro-organisms areneeded for process automation. Sensors for environ-mental applications have been reviewed by other au-thors [2,4] and will not be considered here. Aninteresting view on the role of electrocatalysis for elec-trochemistry and environment has recently been givenby Trasatti [9].

2. Cathodic water and effluent treatment

Waste water from different industrial sectors, such aselectroplating, photographic development, printed cir-cuit board production or battery technology, requiresspecial treatment to remove toxic metal ions or recy-cling of valuable material. Due to increasingly stricterregulations for the discharge of effluents, the permissi-ble concentration of metal ions has been strongly de-creased in recent years [10–12]. Table 1 shows thedevelopment of the limiting values for different metalions in Germany from 1983 until 1991 and for compari-son the latest values in Switzerland. The concentrationswhich can be obtained by the conventional technique ofhydroxide precipitation are also listed in Table 1. Obvi-ously, the limits for effluent discharge can only be metfor a few metals like Cu or Zn by the conventionaltechnique. The precipitated metal ions are replaced byalkali or alkaline earth metal ions and in many casesthe salt content prevents the recycling and reuse of thewater. Furthermore, the cost of disposing sludge con-taining toxic metals has also been drastically increasedduring recent years [13]. Costly measures are also neces-sary to avoid the contamination of ground water withmetal ions leached out by the rain. Ion exchangertechniques as an alternative to hydroxide precipitationare gaining increasing importance [14]. However, thismethod is often too costly for many types of effluentsand suitable ion exchanger resins are not available forall metals. This situation has been the incentive for thedevelopment of new and efficient processes for thetreatment of waste water.

2.1. Optimisation of cell design

Since most of the metal ions can be removed bycathodic deposition, electrochemical processes havebeen developed, some of which are already commer-

Table 1Effluent limits and metal concentrations obtained by hydroxide precipitation [12]

Metal concentrationEffluent limits in ppmMetalafter precipiationat pH 8 in ppmD D CH

ATV A115 Abwasser-VwV 1991Jan 1991Jan 1983

Pb 2 210.5 0.50.5 1500Cd 0.10.22 0.5Cu 0.5 1

Ni 3 0.5 2 340Hg 0.05 0.05 0.01 –

2 0.1Ag 0.1 –2.65 2 2Zn

5 –2Sn 2

K. Juttner et al. / Electrochimica Acta 45 (2000) 2575–2594 2577

cialised and being now used in the industry [15–18].The removal of metal ions Mez+ from waste water isbased on the cathodic metal deposition:

Mesolz+ +ze−UMe (1)

From a thermodynamic point of view, the Nernst equa-tion predicts that it should be possible to decrease theMez+ concentration in solution to an arbitrarily lowlevel, if the potential E of the Me/Mez+ electrode ismaintained sufficiently negative with respect to thestandard potential E°Me/Mez+:

cMez+ =c+·exp� zF

RT(E−E°Me/Mez+)

n(2)

However, at extremely low concentrations, the rate ofthe mass transport controlled process strongly de-creases. In practise, electrolysis at concentrations below0.05 ppm is no longer economical due to increasingelectrolysis time and unacceptable space–time yield.Since the current densities at low metal ion concentra-tions are small, the specific energy demand for anelectrochemical waste water purification process is gen-erally quite low, typically in the order of 0.05 DM m−3

[10]. More relevant are the specific investment costs,which are inversely proportional to the space–timeyield r of the reactor. For the cathodic metal deposi-tion in an electrochemical reactor of volume V, thespace–time yield is defined as the amount of metal dmdeposited in a time interval dt :

r=1V

dm

dt(3)

According to Faraday’s law, dm is proportional to theelectrolysis charge Af ei dt :

dm=Af ei dtM

zF(4)

where f e is the current efficiency, A the actual electrodearea, and M the molecular weight of the metal de-posited. The optimal process conditions are met whenthe rate of the heterogeneous reaction attains its maxi-mum at the limiting diffusion current density, i= il:

il= −zFDMez+

cMez+

d= −zFkmcMez+ (5)

where DMez+ is the diffusion coefficient, d the Nerstdiffusion layer thickness, and km=D/d is the masstransport coefficient. On replacing the current density iin Eq. (4) by il, the space–time yield can be expressedas:

r=f eMaekmcMez+ (6)

This is the key formula for the design and constructionof an electrochemical reactor for waste water treatmentand metal recovery. For a given metal ion concentra-tion cMe

z+ a high mass transfer coefficient km and a large

specific electrode area ae=A/V are essential to obtainhigh space–time yields.

The reactor performance is not independent of thewaste water properties. Therefore, Kreysa [19,20] intro-duced the definition of a normalised space velocity rn

s

for the characterisation of the reactor performance fora given degree of conversion:

r sn=

If e

(ci−ce)VzFlog

ci

ce

(7)

This figure gives the volume of waste water in cm3

treated by reducing the inlet concentration ci by afactor of 10 (ce=0.1 ci) within 1 s in a reactor volumeof 1 cm3.

2.2. Electrochemical reactors and their application

Different types of cell constructions have been de-signed during recent years. Efficient cell design has beendirected towards optimising of the space time yield r

according to the key formula (Eq. (6)), i.e. high specificelectrode area ae and/or large mass transport coefficientkm. With respect to these criteria, the electrochemicalcells may be classified by the following three groups[10]:� Improved mass transport and thus increased current

density by setting the electrodes in motion or byapplying turbulence promoters, but with a relativelysmall electrode area in a given cell volume. Examplesare the Pump cell [21,22], the Chemelec cell [23], theECO cell [24–26], the Beat rod cell [27,28] and cellswith vibrating electrodes or electrolytes [29].

� Attempts to accommodate large electrode area in asmall cell volume resulted in developments such asthe Multiple-cathode cell [30], the Swiss-role [31] orthe ESE (extended surface electrolysis) cells [32].

� Improved mass transfer coefficients and enlarged spe-cific electrode area are provided by the use of three-dimensional electrodes. Examples are the Porousflow-through cell [33], the RETEC cell [2], thepacked-bed cell [34–36] the fluidised bed cell [37–41]and the rolling tube cell [42,43].The design of the three-dimensional electrodes origi-

nates from the fact that in contrast to the planarconfiguration the electrode area is distributed through-out all three dimensions. Due to the specific fluiddynamic conditions inside such a three-dimensional bedelectrode, this arrangement provides not only a largespecific electrode area but also large numbers of themass transport coefficient. Some example of cell con-structions are shown schematically in Figs. 1–10.

The Pump cell, which is the rotating analogue to thecapillary gap cell [44], is shown in Fig. 1. Although thiscell has not yet been employed on an industrial scale itis interesting for its operation principles. It consists oftwo static disc electrodes with current connections and


Fig. 1. Sketch of the Pump cell.Fig. 4. Cascade configuration of the ECO cell.

Fig. 2. Sketch of the Chemelec cell.

a bipolar rotating disk mounted in between. The elec-trolyte flows radial from the central tube to the outercircumference. The advantage of this design is that themass transfer coefficient can be controlled indepen-dently of the electrolyte flow and the residence time ofthe solution.

The Chemelec cell in Fig. 2 (manufactured by BewtWater Engineers, Alcaster, England) uses a fluidisedbed of glass spheres as turbulence promoters to im-prove the mass transfer to the electrodes consisting of aseries of closely spaced gauze or expanded metal sheetelectrodes. The residence time and the degree of conver-sion per pass are relatively low, because the electrolyteflow has to exceed a minimum fluidisation velocity ofthe bed. This cell is therefore suitable for pre-treatmentor recycle operations and is commonly used in the

Fig. 3. Removal of Ni2+ ions from rinsing water and recycling to the galvanic bath.


Fig. 5. Beat rod cell with concentration–time curve for the removal of silver.

Fig. 6. Sketch of the Swiss role cell with foil electrodes and plastic separator mesh.

electroplating industry for maintaining a moderatemetal ion concentration of ca. 50 ppm in a re-circulatedwash-water of a rinsing tank as shown in Fig. 3.

Another configuration that achieves high mass trans-port coefficients are cells with rotating cylinder elec-trodes (RCE) and a small gap between the cylindricalcathode and anode. High rates of mass transport areexperienced in the turbulent flow regime, so that RCEreactors allow metal deposition processes to take placeat high speeds, even from dilute solutions. RCE reac-tors have been operated at a scale involving diametersfrom 5–100 cm, with rotation speeds from 100 to 1500rpm and currents from 1 A to 10 kA [45]. It is possibleto design RCE reactors with scrapping devices to re-move the deposited metal continuously in powder formfrom the cathode surface. The so-called ECO cell (Fig.4) designed in the UK, has already been produced andapplied industrially. In the cascade configuration, a

single rotating cylinder of about 50 cm diameter servesas the cathode for all chambers. The segments throughwhich the solution successively flows eliminate back-mixing effects to a great extent so that a high degree ofconversion can be achieved.

A cell with a special cell design is the Beat rod cell(Fig. 5) which is mainly used in small electroplatingshops. Cathode rods supported in the electrolysis tankare slowly rotating within an annular type of chamber.The rods are striking one another so that the metal

Fig. 7. Flow-through and flow-by arrangements of cells with3D electrodes.


Fig. 8. Packed bed cell-design of the enViro cell.

permeable plastic mesh are wrapped around a centralcore. Inhomogeneous potential distribution due toohmic drop within the thin foil is largely compensatedby contacting the electrodes at opposite sides. Theelectrolyte flows axially through the electrode pack.When the cell is filled with deposited metal, it is regen-erated by treatment with acid similar to common ionexchanger resins. Large space–time yields are realisedand mass transport conditions are favourable since theseparator functions as a turbulence promoter. The ca-pacity of the cell can be further increased by usingperforated electrodes, thus achieving a truly three-di-mensional electrode structure. Mesh electrodes wrappedaround a perforated winding core, lead to anotherversion of the Swiss role cell with radial flow.

Another design which uses mesh electrodes instead offoils is known as ESE cells. These cells have beensuccessfully applied on an industrial scale for the purifi-cation of copper containing effluents. A quite efficientcommercial system for waste water treatment is theRETEC cell (Eltech System Corp., Cardon, Ohio). Thedesign is similar to a simple tank house cell, but con-tains 6–50 three-dimensional cathodes constructed asflow-through metal sponge electrodes [2]. Thesecathodes have an active surface area of approximately15 times their geometrical area. Between each pair ofcathodes is placed an inert, dimensionally stable anode(DSA), which is an oxide-coated titanium mesh. Thecell is used in closed water recycling systems in theelectroplating industry with applications similar tothose of the Chemelec cell.

Packed bed cells with three-dimensional electrodesare obtained when using electrodes consisting of elec-tron conducting particulate material through which theelectrolyte can flow. Numerous versions of these porouscells have been described in the literature [17,18,33,46–48]. Two principle arrangements with respect to thedirection of the current and electrolyte flow are possi-ble. They are denoted as flow-through and flow-byarrangements, Fig. 7. For technical applications theflow-by arrangement is preferred due to minimisationof ohmic losses and limitations in the thickness of thethree-dimensional electrode. With increasing distancefrom the counter electrode, the local current density of

deposit is made to settle as a powder which can bedischarged at the bottom. This reactor type can operateonly batch-wise and is restricted to small throughputsonly.

An efficient way of fitting a large planar electrodearea in a small volume is the design of the Swiss rolecell, Fig. 6. Thin metal foils separated by an electrolyte

Fig. 9. Fluidised bed electrolysis cell according to Fleischmannand Goodridge. Fig. 10. The Rolling tube cell.

K.

Juttneret

al./E

lectrochimica

Acta

45(2000)

2575–

25942581

Table 2Industrial applications of the enViro cell [11]

Inlet concentration Outlet concentration Energy consumption Anode areaMetal TroughputApplication field(ppm)(ppm) (kWh m−3)(m3 h−1) (m2)

0.05 1.2 1HgProduction of measuring instruments 0.3 3001Film processing 0.151.0Ag 150.2

0.07Pb 10.5 2 0.1Salt productionCd 0.18 10.2 20Electroplating 0.1

3500 1.70.010.08Battery production Hg/Cd0.08Cu 4020 20 1.9Cellulose acetat production

50 0.19 5CuPickling (recycling of solution) 3 1502.0 4.0 906Dyestuff production 400Cu

15Hg 0.05 2.52 4Dyestuff production

Table 3Operation data for different types of waste water electrolyis cells [11]

C1 (ppm) VR (l) VD (l h−1) i (A cm−2) I (A) UZ (V) feProcess Scale Removed metal C0 (ppm)

8000 1000 0.059 122 0.65275100CuInd.ECO cell12.5 125 – 5Beat rod cell 0.15Ind. Ag 6000 4 2000.7 0.55 0.46×10−3 1.56 0.380.3Swiss-roll cell 25380CuLab.0.64 0.45 3.5×10−3 1.46 0.95Porous flow-through cell Lab. Cu 800 0.2 2.6

50 3.2 0.76×10−3 1.9 0.664.80.150Packed bed cell Ind. Cu192Ind. 7000 600 – 3.1 0.71Cu 77 5Fluidized bed cell

0.4 40 6.9 15 – 6.05 0.005Rolling tube cell Ind. Au 81


Table 4Specific energy consumption Ee

s and normalised space velocityrn

s , for different electrolysis cells for industrial applications[11]

EesProcess E t

s rns

(l lh−1)(kWh m−3)(kWh m−3)

1.5ECO cell 3 2050Beat rod cell 60 0.2

–1.23 20Swiss-roll cell0.91.03Porous flow- –

throughcell

Packed bed 0.12 – 28cell

0.27 – 30Fluidized bedcell

14.2 – 0.4Rolling tubecell

are gathering at the lower part of the cell where theycan be collected and replaced by fresh particles fromthe top. Because the height of the cell is restricted toabout 2 m for hydraulic reasons and the fluidisationvelocity has to be maintained at a relatively highlevel, short residence times and thus only limited de-gree of conversion per pass can be achieved. Continu-ous operation with solution re-circulation and cascadearrangements of cells are therefore used in practicalapplications. Fluidised bed electrolysis has been exam-ined for different applications in many laboratories[37,50–54] and has also been applied on an industrialscale for metal recovery [40,41]. A typical example isthe removal of copper from the process liquor in cel-lulose acetate production for membrane fabrication ofartificial kidneys [11].

Moving and circulating bed electrodes combinedwith a rotating cell design are described in the litera-ture [55] as moving bed or spouted bed electrodes.Fig. 10 shows an example of the rolling tube cell[42,43]. In principle this cell resembles the well-knownplating process for piece-goods, using slowly rotatingbarrels. The rotating drum is only partially filled withgraphite particles to achieve thorough agitation. Thecathode has the shape of a perforated annular-typedrum and is surrounded inside and outside by anodesto render the current distribution as homogeneous aspossible. This cell has proved to be especially usefulfor silver and gold recovery. Further developments ofmoving bed cells and their application can be foundelsewhere [56]. Three other types of circulating partic-ulate electrodes for copper recovery from dilute solu-tions — the spouted bed, vortex bed and moving bed— have been described by Scott [55].

An alternative to electrically driven cells using highsurface area electrodes such as packed or fluidisedbed electrodes is a cementation system where electro-chemical activity takes place spontaneously and amore noble metal may be deposited onto a less noblesubstrate which itself progressively dissolves [57].Thus, no electrical energy is consumed in the precipi-tation reaction although power is required for elec-trolyte circulation. Well known examples ofcommercial cementation processes are the the finalpurification step in leach liquor treatment prior toelectrolysis in zinc electrowinning where zinc dust isused as precepitant, and the recovery of silver fromspent photographic solutions using steel substrates.Power and Ritchie [58] have contributed a valuablegeneral review of this topic, and Rickard and Fuer-stenau [59] among others have undertaken fundamen-tal mechanistic studies of the copper/iron system.Wragg and Bravo de Nahui [57] have recently studiedthe influence of process parameters on copper cemen-tation onto a packed bed of iron particles in flow

the bed approaches almost zero. The potential distri-bution and the optimum penetration depth of thecurrent density has been calculated as a function ofthe effective particle (kp) and solution conductivity(ks) for different cell geometries. For a rectangularcell under limiting current conditions, the followingrelation for the optimal bed depth is obtained [10,49](kp�ks):

hopt=' 2nks Dh

aezFkmcMez+

(8)

The over-voltage range Dh is determined by the po-tential region of the limiting current density plateauin the micro-kinetic polarisation curve and n is thevoid of the three-dimensional electrode structure(ks=k s

on). According to Eq. (8), the optimal beddepth increases with decreasing metal ion concentra-tion cMez+. This has led to the enViro cell designdepicted in Fig. 8, which can operate satisfying atlow metal ion concentrations. For example, in a sin-gle-pass operation, the concentration can be reducedfrom 10–50 ppm to less than 1 ppm. This cell has ahigh space time yield at a residence time of a fewminutes only. The metal ion concentration can be re-duced by up to a factor of 1/1000 in a single pass.The enViro cell has found wide spread industrial ap-plication as summarised in Table 2.

The principle of fluidised bed electrolysis originatesfrom Backhurst et al. [38]. The electrolyte flows frombottom to top through a loose bed of particles, Fig.9. The advantage is that the particles are held insuspension, thus avoiding bed blockage as found infixed-bed cathodes and allowing continuous metal re-covery. The fluidised particles are cathodicallycharged via a feeder electrode. The growing particles


through and batch recycle systems. Also reductivede-halogenation of chlorinated hydrocarbons such aslindane and chloroform has been demonstrated in a bedof iron scrap or zinc particles [60]. Similar toprecipitation, these metals served as sacrificial anodesproviding electrons for the reductive destruction anddichlorination of the hydrocarbons.

2.3. Operation data of electrochemical cells

For a better comparison of the different reactors andcell design described above, a summary of operation datais given in Table 3. Since these data are obtained fromcells of different sizes and process conditions, they areintended to give a rough guide rather than a reliable basisfor assessing the individual reactors [10].

Another way of comparing the different processesconsidered so far is given by the data in Table 4 [10,12].These data are closely related to the economic efficiency.The specific electric energy demand E e

s for the treatmentof 1 m3 of effluent is relevant for the calculation of theenergy costs. For cells containing rotating parts, the totalenergy consumption E t

s is also specified. Another factorwhich strongly affects the total costs is the normalisedspace velocity rn

s , which is also given in Table 4.Considering both the specific electrical energyconsumption and the normalised space velocity ofprocesses that have already found industrial application,the packed bed and fluidised bed electrolysis seem toperform best.

3. Anodic treatment of aqueous effluents and organic waste

By an electrochemical treatment of wastes either apartial (reduction of toxicity) or a complete decomposi-tion of the pollutants can be achieved. Complete de-composition of organic material means the oxidation tocarbon dioxide and as consequence a relatively highenergy consumption for large organic molecules. Theattempts for an electrochemical oxidising treatment ofwaste water or wastes can be subdivided in twocategories:� direct oxidation at the anode and� indirect oxidation using appropriate anodically formed

oxidants (Cl2, hypochlorite, peroxide, ozone, Fen-ton’s reagent, peroxodisulphate).A special case with the indirect oxidation frequently

called mediated electrochemical oxidation (MEO) is theuse of metal ions with high oxidation potential (e.g.Ag(II) and Co(III)) which can be electrochemicallygenerated in a closed cycle largely avoiding emissions.

3.1. Direct oxidation at the anode

For the direct oxidation two main properties are

required for a suitable anode: high oxygen overpoten-tial and corrosion stability.

A model described in Ref. [61] for the course of theanodic oxidation of an organic molecule assumes threemajor steps:

1- Discharge of water forming an adsorbed hydroxylspecies:

S[ ]+H2O=S[OH]+H++e−

2- The adsorbed OH is the ‘activated state’ of waterin O-transfer reactions to the organic molecule R:

S[OH]+R=S[ ]+RO+H++e−

3- Co-evolution of O2 by oxidation of water dimin-ishing the current efficiency:

S[OH]+H2O=S[ ]+O2+3H++3e−

Phenol and derivates are the mostly investigated ex-amples of the electrochemical studies. Main intermedi-ates found are benzoquinones, maleic acid and hydroxydiphenyls, which can be further oxidised to CO2 andsmall amounts of CO. A voltammetric study at aplatinum rotating ring disk electrode with phenols,chlorophenols, organic acids and a surfactant (fluori-nated octanoic acid) is presented in Ref. [62]. Differentanode materials like platinum as reference [63], nickel,glassy graphite and titanium supported oxide electrodes(IrO2, RuO2, PbO2 and SnO2 [64]) and Bi as well as Fedoped PbO2 [61] were tested. At phenol concentrations�1 mM fouling of the noble metal anodes was ob-served by films of polymer material. A high rate wasachieved with lead dioxide anodes using a packed-bedcell [65,66]. The oxidation of phenol was complete atoptimum conditions of pH and temperature in theseexperiments; the main product was carbon dioxide andbenzoquinone and maleic acid were identified as by-products.

Comninellis has investigated the oxidation of numer-ous aromatic compounds and for industrial waste watertreatment the use of a pilot plant equipped with Ti/SnO2 anodes is reported [67].

Other applications have been researched for directanodic oxidation are: sugars, alcohols, distilleryeffluents, dyes, aromatics etc. [68]. In the oxidation ofazodyes beside CO2, N2 and sodium sulphate can beformed aromatic esters, phenols, aliphatic carboxylicacids, cyclic and aliphatic hydrocarbons, aromaticamines etc. In the presence of chloride the oxidation ofazodyes happens mainly through the formation of ‘ac-tive’ chlorine [68].

Scott [69] has investigated the anodic oxidation ofaromatic and aliphatic compounds using batch andflow electrolysis in divided and undivided cells. Theresults confirm the suitability of several anode materialsfor the oxidation of phenols. For the intermediates(especially aliphatic species) the anodes are, however,


Fig. 11. Principle of the MEO process.

3.2. Indirect oxidation

The most used oxidant electrochemically generatedbelonging to this category is chlorine or hypochlorite.Besides peroxide, Fenton’s reagent [74] and peroxidisul-phate, ozone is the other prominent oxidant which canbe produced electrochemically. ABB Ltd and OxyTechLtd use hollow, cylindrical fluorocarbon-impregnatedcarbon anodes are for ozone generation in a solid-poly-mer electrolyte cell [68].

As mentioned already above the use of metal ions withhigh oxidation potential like Ag(II), and Co(III) cancompletely oxidise organic compounds to CO2. There-fore, this frequently called mediated indirect oxidationmethod was developed for the treatment of hazardousorganic wastes with only a limited amount of water wherethe oxidant is continuously electrochemical generated ina closed cycle. The method was originally used in thenuclear industry to dissolve refractory plutonium dioxidein nitric acid [75,76]. Ag(II) was the most efficient oxidantfor this purpose. This experience initiated the treatmentof mixed (radioactive) wastes (e.g. tributyl phosphate,solvent of the purex process) at the Savannah RiverLaboratory, USA and at AEA Dounreay, GB. Thesuccessful experiments were the kick-off for the processdevelopment activities at the Lawrence Livermore Lab-oratory [77], at AEA Dounreay [78] and at theForschungszentrum Karlsruhe [79], where the first timeintermediates and reaction paths for the oxidation ofphenol and chlorinated phenols were identified, to applythe method for a plurality of hazardous wastes. Theprinciple of this process is shown in Fig. 11. For totaloxidation only redox couples with a high oxidationpotential are suitable, for example Co(III)/Co(II) with1.82 V (NHE) and especially Ag(II)/Ag(I) with 1.98 V(NHE). Ag(II) is generated anodically from Ag(I) in anaqueous nitric acid solution using an electrochemical cell,which is divided by a membrane.

In the cathodic compartment nitric acid is reducedaccording to the Vetter mechanism [80] to NO which canbe regenerated to nitric acid by the well-known oxidativeabsorption in columns [81]. Experiments with electro-chemically generated Mn(III) in sulphuric acid [82],which generation was intensively investigated byComninellis for production of aldehydes [83], showed incounter-current columns a very fast and complete oxida-tion of NO to HNO3.

In this procedure, using two electrochemical cells, onefor destruction of organics and the other for nitric acidrecovery, hydrogen is the final cathodic reaction product.Recent investigations of Forschungszentrum Karlsruhetogether with Eilenburger Elektrolyse- und Umwelttech-nik have combined these two electrochemical cells intoone, using silver containing nitric acid as anolyte andsulphuric acid as catholyte, avoiding the NO formationat the cathode.

less effective. For the oxidation of aliphatic acids (oxalicand glyoxylic acid) platinum and DSA are not or lesseffective. Graphite and tin oxide anodes are also suitablefor the oxidation of formic acid.

For the direct oxidation of SO2 in solution Lu andAmmon [70] have carried out a kinetic study using Pt,Pd, Ru, Ir, Rh, Re and Au. The DSA PdOx/Ti showedthe highest electrocatalytic activity in these experimentswhilst RuOx�TiO2/Ti and IrOx�TiO2/Ti were inactive.For the sulphite to sulphate oxidation respectable currentefficiencies were also achieved using Pt/Ir coated titaniumelectrode in an undivided cell [69]. A problem for theoxidation of SO2 is to avoid the formation of elementalsulphur at the cathode. Ebonex® has been tested ascathode in the experiments [69].

The direct anodic oxidation of SO2 as a method forflue gas treatment has been studied on several electrodes(see e.g. [71]). At graphite electrodes are the productssulphate and dithionate. A continuous neutralisation ofthe generated acid is required in the process.

Numerous fundamental investigations and even benchscale experiments have been done for the anodic treat-ment of aqueous effluents containing further inorganicpollutants, e.g. cyanide [72], thiocyanate [69] and sodiumdithionite [73].

Fig. 12. Bench scale plant at Forschungszentrum Karlsruhe.


Table 5Destruction rates of selected model substances

TOC in anolyte (ppm)Substance Destruction rate (%)T (°C)

Benzyl alcohol 2–540–70 \99870 \99Chlorobenzene5–101,2,4-Trichlorobenzene \9940–70570 \991,3,5-Trichlorobenzene10 \991,2,3,4-Tetrachlorobenzene 901090 \991,2,3,5-Tetrachlorobenzene

90Hexachlorobenzene 12 000 2590Pentachlorophenol (PCP) B2 \99

1070–90 \99Lindane (g-HCH)701,2-Dichloroethane 2 \99

370 \991,1%-Dibutylsuyi-de6Heptafluoro butyric anhydride \99701470 99Trichlorfona

90Chlophenb 10 \99

a 2,2,2-Trichloro-1-hydroxy-ethylphosphonic acid dimethyl ester.b Transformer oil, containing tetra- to octachlormated biphenyls (PCBs).

The formation kinetics of the oxidising state of themediators were investigated by RDE experiments [84].For the Ag(II) system under favourable flow conditionsvery high current densities of more than 5 kA m−2

could be obtained. The formation of Ag(II) is mainlytransport controlled; an average value for b of 0.488cm s−1/2 was determined. Studies made on Co(III)generation showed that anodic oxidation of water andCo(III) formation are probably simultaneous reactions.The kinetics of the oxidation of water by the mediatoras function of temperature — the most important sidereaction diminishing the electrochemical efficiency —was also investigated to find the optimum temperaturefor the process [84].

For the process development destruction experimentswith different model substances and real wastes wereperformed in a laboratory plant and additionally in abench scale plant, Fig. 12. The experiments in labora-tory scale were done in a small filter press cell. Anodeand cathode compartment were divided by a cationexchange membrane. Both electrode areas were 11 cm2

and the volume of anode and cathode compartmentwas 6 cm3 each [85]. The bench sale plant is similardesigned than the laboratory plant with a scale upfactor of 100. The central part of this plant consists ona monopolar electrochemical cell delivered by theSwedish company Electrocell AB or alternatively abipolar cell from the German company EilenburgerElektrolyse- und Umwelttechnik GmbH can be used.Both cells have anode and cathode areas of 0.12 m2

each. A platinum foil anode is used while titanium andcarbon, respectively, are used as cathode. The maxi-mum current density is limited by the Nafion 450membrane to 5 kA m−2 corresponding to a maximumcurrent of 600 A. The anodic and cathodic cycles are

equipped with magnetically coupled pumps. Liquid or-ganic waste is fed into the anodic cycle by a pump,solid waste is fed into the anodic glass vessel from thetop. Due to high flow rates of the electrolyte up to 1500l h−1 the organic waste is fine dispersed in the elec-trolyte to minimise the transport inhibition at theaqueous/organic boundary layer [86].

In case of degradation of chlorinated organic com-pounds intermediate precipitations of silver chlorideoccur, which are redissolved in presence of an excess ofAg(II) by the oxidation of chloride to perchlorate,which is enriched in the anolyte and can be separatedfrom time to time by cooling crystallisation as potas-sium perchlorate [87].

Cathodic formed nitrogen oxides are oxidised backto nitric acid in a bubble column, which is continuouslyfed with a sulphuric acid Mn(III) solution generated atthe anode of a second electrolytic cell, while hydrogenis generated as a useful by-product at the cathode ofthis cell.

In the laboratory plant the destruction of a largevariety of different, mostly chlorinated, model sub-stances were investigated. In all experiments the elec-trolyte consisted of 0.5 M AgNO3/7 M HNO3. Theexperiments were carried out with current densitiesbetween 1 and 5 kA m−2 and reaction temperatures of40–90°C. The destruction rates were calculated fromthe residual organic carbon contents (TOC) in theanolyte at the end of the reaction, as shown in Table 5.

All selected model substances with the exception ofhexachlorobenzene could be destroyed to more than99%. As reaction products always CO2 and smallamounts of CO were formed. To study the behaviour ofchemical warfare agents mustard gas was simulated byan equimolar mixture of 1,2-dichloroethane and 1,1%-


Table 6Destruction of obsolete pesticides

Fig. 14. Destruction of industrial waste in bench scale plant(continuous mode).

The release of CO2, CO, and O2 was determined on-line. After the reaction, the remaining organic com-pounds were extracted with toluene and analysed byGC/MS. The amount of total organic carbon (TOC)was determined in the anolyte. The chemical formulasof the active compounds and the results of the de-struction experiments are illustrated in Table 6.

The first four samples were decomposed completely,from the fifth sample a residue, fine white flakes, wereobserved after the experiment and small amounts ofDDT were analysed in the flakes. An experiment per-formed with pure DDT showed a slow destructionunder the same experimental conditions, but noresidue and no DDT could be detected. Therefore, itcan be assumed that the DDT residue found duringthe destruction of sample c5 is bound in the matrix(wax or polymer) which shields the DDT against elec-trochemical attack.

Moreover several organic multi-component wastemixtures from the chemical industry were treated inthe laboratory and bench scale plants. Fig. 13 showsas an example the current efficiencies for the singlereaction products analysed in the anodic off-gas ofthe laboratory cell during a batch experiment of sucha waste, containing more than 30 different compo-nents. The main components were benzoylchloride,chloro benzoylchlorides, chloro benzaldehyde, benzotrichloride and benzoic anhydride. The experimentwas performed at 60°C using a current density of 4kA m−2. Starting the experiment CO2 evolution oc-curs and current efficiencies up to 100% wereachieved. After 60 min CO2 evolution decreases andO2 formation, caused by water oxidation, increases.From the point of intersection of the both curves thecurrent demand of 4.35 kA h−1 kg−1 for the carbonoxidation was calculated. Afterwards the same wastewas treated in the bench scale plant in continuousmode (Fig. 14) dosed referring the current demand

dibutylsulfide. Soman or Sarin were simulated by thepesticide Trichlorfon, which has a similar chemicalstructure but does not have the physiological effectiveC�F bond, and heptafluoro butyric anhydride asfluorinated compound. As pure pesticide amongTrichlorfon the prevailed Lindane (g-hexachlorocyclo-hexane) was treated. All these substances are be de-stroyed completely.

Five partly decomposed pesticide samples (Pakistan,15 years old) containing different active substancesand matrices, sampled from the German Gesellschaftfur Technische Zusammenarbeit, were treated in thelaboratory plant. The electrolyte solution consisted of0.5 M AgNO3 in 7 M HNO3 solution in all cases.

Fig. 13. Destruction of industrial waste in laboratory plant(batch experiment).


Fig. 15. Mobile demonstration plant.

determined before. From Fig. 14 it is evident, thatdirectly after stopping the organic waste flow CO2

formation decreases and O2 formation increases. Thisindicates, that the current demand determined from thebatch experiment in the laboratory is applicable to thescaled up plant in good accordance. In average thecarbon balance is closed better than 95%.

Based on the described experiences and founded by theEuropean Commission in the frame of a CRAFT projecta mobile demonstration plant was designed and built,which is operated on an industrial site since March 1999(Fig. 15). In case of destruction chlorinated organicwastes, the plant offers both possibilities, separation ofsilver chloride from the anolyte and recovering silver oroxidising the organic chlorine to perchlorate and winningof potassium perchlorate. Sulphuric acid is used ascatholyte, generating hydrogen at the cathode as avaluable by-product. The migration of silver ions fromthe anolyte through the membrane into the catholyte wasdetermined with only 0.03 g (Ah)−1, which is reducedand deposited at the cathode, flushed out periodically,separated from the catholyte and recycled to the anolyte.

4. Electrochemical gas purification

4.1. General aspects

An increasing demand for off-gas purification, partic-ularly smaller scale power plants, heating combustionunits or chemical plants, has encouraged the develop-ment of new concepts of electrochemical gas purifica-tion techniques. Many gaseous pollutants, such aschlorine, hydrogen sulphide, nitrous oxides, or sulphurdioxide permit electrochemical conversion in anaqueous environment, since the standard potentials ofthe corresponding reactions are all within the stabilityrange of aqueous electrolytes:

2Cl−��Cl2+2e− E°=1.360 V

H2S��S+2H++2e− E°=0.142 V

NO+2H2O��NO3

−+4H++3e− E°=0.957 V

NO2+H2O��NO3

−+2H++3e− E°=0.775 V

SO2+2H2O��SO4

2− +4H++2e− E°=0.138 V

Different concepts of electrochemical gas purificationcan be found in the literature [11].

Fig. 16. Electrochemical gas purification by inner cell andouter cell processes.


Fig. 17. An electrochemical absorption column with 3D parti-cle bed electrode.

dissolved pollutant is directly converted at the surfaceof the packed bed electrode. Such a device exhibits highspace–time yield and has been successfully tested forelectrochemical absorption of sulphur dioxide and chlo-rine [89,92,93].

Indirect electrochemical processes with homogeneousredox mediator are the so-called ‘Peracidox’ processdeveloped by Lurgi [94] using peroxodisulphate as theredox mediator for SO2 oxidation, and the modified‘Mark 13 A’ process developed at the Joint EuropeanResearch Center Ispra [95], using bromine as mediatorfor the indirect oxidation of SO2. In both cases theelectrochemical regeneration of the redox mediator isperformed by an outer cell process in a separate electro-chemical cell.

Electrochemical gas purification processes with het-erogeneous redox mediators are the lead dioxidecatalysed oxidation of SO2, suggested by Strafelda andKrofta [96] and the copper catalysed process for thecombined electrochemical and catalytic removal of SO2

at a Cu2O/Cu2+ catalyst [97]. Porous gas diffusionelectrodes have been employed for the conversion ofSO2 using pyrolised Co phtalocyanines as electrocata-lyst [98]. Electrode technology borrowed from the fuelcell technology have also been used for effluent gastreatment. A cell construction similar to that of amolten carbonate fuel cell with porous electrodes and amolten Na2S�Li2S electrolyte at 623 K has been exam-ined for removal of SO2 and H2S [99–102]. High tem-perature solid state electrochemical cells based onyttrium stabilised zirconium dioxide YSZ and porouspalladium electrodes have been studied by Hibino [103]for the successful electrochemical removal of both NOand CH4 at 650–750°C in presence of an oxidisingatmosphere. NO is reduced to N2 and CH4 is oxidisedto COx..

4.2. New process de6elopments

Since flue gases not only contain SO2 but to a certainextent also NOx, the development of processes for thesimultaneous removal of both components has been thesubject of recent investigations [93,104–107], which ledto the following alternative reaction schemes:

The cerium(IV)-assisted process [104,105] is an indi-rect outer cell process with Ce4+ as homogeneousredox mediator for the simultaneous oxidation of SO2

and NOx to sulphuric acid and nitric acid, respectively.The process scheme and the main reactions are depictedin Fig. 18. The acidic solution is continuously fed downto the anode compartment of an electrochemical cellwhere Ce4+ is regenerated. When the solution hasreached a critical concentration of ca. 30–40% HNO3

and H2SO4, separation of the nitrate effluent from thesulphate solution and recycling of Ce3+ are performedin a separate unit, where the nitric acid is first separated

In any case, the initial step is an absorption of thepollutant species into a liquid phase. Since the solubilityof several gases in aqueous solutions is too small, thetransfer from the gas phase into the solution phasemust be supported by a reaction, which converts theprimarily dissolved species permanently to a more solu-ble one. In the simplest case, this can be either thedirect conversion at the electrode of an electrochemicalcell [88–90], or indirectly via chemical reaction with aredox mediator, which can be electrochemically regen-erated in a follow up step. Homogeneous and heteroge-neous catalysts have been applied as redox mediators.The advantage of heterogeneous mediators, such asoxides, is that a separation of the reaction productsfrom the mediator is not necessary. Both, the electro-chemical conversion of the pollutant and the electro-chemical regeneration of the redox mediator can beachieved either by an inner cell or an outer cell process,as illustrated in Fig. 16 [91].

Direct electrochemical conversion of gases by aninner cell process can be carried out in a speciallydesigned electrochemical absorption column, whichconsists of a three-dimensional packed bed electrode ofconducting particles in contact with a cylindrical feederelectrode and a counter electrode separated by a porousdiaphragm or ion exchanger membrane [88–90].

Fig. 17 shows schematically the construction of sucha device, which usually operates under counter currentflow conditions of the gas and the liquid phases. The


Fig. 18. The Cerium(IV)-assisted process.

by distillation and the excess of sulphuric acid is thencontinuously removed by a liquid/liquid extractiontechnique [108,109].

The lead dioxide dithionite process [93,106,107] com-bines direct and indirect conversion of SO2 and NOx,respectively, in two steps with dithionite as homoge-neous redox mediator for the indirect reduction of NOx

and lead dioxide as heterogeneous catalyst and media-tor for the direct oxidation of SO2. The gas mixtureenters first an absorption column where NO, as themain component of nitrous oxides in flue gases, isabsorbed by a complex forming agent Fe(II)EDTA andsimultaneously reduced with dithionite, S2O4

2−. Theredox mediator dithionite is continuously regeneratedby cathodic reduction of SO3

2− in an electrochemicalcell. The second component, SO2, passes the NO ab-sorption column without reaction and enters an electro-chemical cell, for example as shown in Fig. 17, where itis oxidised directly to sulphuric acid at the anode. Thefirst stage of this process, the indirect electrochemicalconversion of NOx, is shown schematically in Fig. 19.A pilot plant for the treatment of 100 Nm3 h−1 of fluegas with 600 ppm NO inlet concentration had beentested on an industrial site[Brite Project 2026].

Microkinetic studies of NOx absorption revealed thatthe degree of NO conversion can be significantly im-proved, if Fe(II)EDTA is present in the absorptionsolution at a low content. This is demonstrated forexample in Fig. 20. On addition of Fe(II)EDTA at t1,the partial pressure of NO at the outlet is drasticallydecreased and remains at lower level as long as dithion-ite is present in the solution. Using gas phase andsolution analysis, several reaction products, such asNH4

+ and amidosulphonic acid in addition to SO32−

could be identified, while the formation of N2 wasfound to be negligible small. Optimisation of this pro-cess requires further investigations into the complexchemistry of the nitrogen oxides [110–112].

Another interesting aspect of electrochemical promo-tion of environmentally important heterogeneous gasphase reactions is the so called NEMCA effect (non-

faradaic electrochemical modification of catalytic activ-ity). NEMCA, which has been described for over 60catalytic systems [113–115], significantly increases thecatalytic activity and selectivity of metal and metaloxide catalysts deposited on solid electrolytes such asY2O3-stabilised ZrO2(YSZ) or b%%-Al2O3 by electro-chemical polarisation of the catalyst electrode. The rateof the catalytic reaction is typically increased by afactor of 5–100. Technical applications are still in theinitial stage of development. The positive influence ofNEMCA on the catalytic reduction of NO by C3H6 inthe presence of O2 leading to the formation of CO2 andN2 has recently been reported [116].

5. Process-integrated environmental protection

5.1. Recycling of process streams

In the chemical industry and particularly in the gal-vanic branch, the concept of zero effluent technologyhas been widely accepted, leading to clean and wellconditioned process streams which can be recycled tothe manufacturing process. Examples have already beengiven in Section 2. The different devices for electrolyticeffluent treatment and metal recovery from the rinsingwater of electroplating processes can also be used torecycle the process water as well. In fact, recycling turnsout to be easier than a final treatment of the effluent,

Fig. 19. Indirect electrochemical NOx removal.


Fig. 20. Effect of Fe(II)EDTA (1 mM) on the absorption ofNO with dithionite (0.16 M S2O4

2−); pH 6.3; gas flow rate 3.9l min−1, NO inlet concentration 7200 ppm, addition ofFe(II)EDTA at t1.

products of hydrogen and oxygen. Experiments haveproved that a large part of the energy can be savedby using gas diffusion electrodes [118]. Either a hy-drogen anode works with the hydrogen evolved at thecathode or an oxygen electrode consumes the oxygenproduced at the anode. The first alternative is themore favourable avoiding the high overvoltage of theoxygen electrode. The use of bipolar membranes of-fers another, energetically favourable way of conver-sion salts to acid and base, as shown in Fig. 21. Thebipolar membrane BM can be considered as a combi-nation of a cation and an anion exchanger mem-brane. Water in the membrane dissociates at theinterface where the H+ ions are forced to migratethrough the cation exchanger part and the OH− ionsthrough the anion exchanger part of the membrane.Theoretically, the energy consumption for water split-ting in the bipolar membrane, as proposed by AlliedSignal [121], is about 40% of that needed by electrol-ysis with gas evolution. Sulfuric acid concentrationsbetween 5 and 10% and sodium hydroxide concentra-tions between 12 and 16% can be achieved [117].

Combinations of ion exchanger membranes andconventional ion exchanger resins for the treatment ofwater and process solutions with continuous regenera-tion of the ion exchanger have been described in theliterature [112]. This concept significantly reducesstack resistance, power consumption and increases theavailable surface area of conventional electrodialysisprocesses. The ion exchanger materials are continu-ously regenerated electrochemically by H+ and OH−

ions that are produced by water splitting in an ap-plied electrical DC field. The major application ofthis technology is the production of ultra-pure waterand the removal and recovery of heavy metals andprecious metal ions from industrial effluents [122,123].

since recycling requires to keep the metal concentra-tion level at about 50 ppm in a rinsing tank whiledischarging needs to reduce the concentration below 1ppm. Furthermore, by recycling it is easier to selectthe most favourable point in the process to conditionthe stream.

In other sectors of the process industry, large vol-umes of highly concentrated sodium salts resultingfrom neutralisation of acidic or alkaline solutions areeconomic motivations to reinforce the demand forzero effluent technology [3,92]. As a result, there is anincreasing interest in technologies to split the sodiumsalts back to sodium hydroxide and acid. Since rever-sal of the neutralisation reaction is accompanied by apositive change of the free enthalpy, such a processcan only be realised by an electrochemical route. Dif-ferent approaches based on cation and anion ex-changer membranes are competing for this market.An example is the electrolytic splitting of sodium sul-phate in a three-compartment cell with anion andcation exchanger membranes [117–120]. Protons H+

and hydroxyl ions OH− are produced during oxygenand hydrogen evolution at the anode and cathode ofthe cell. The cation exchanger membrane KM permitsselective transport of Na+ ions and the anion ex-changer membrane AM that of SO4

2− so that rela-tively pure solutions of sodium hydroxide andsulphuric acid are formed in the cathode and anodecompartments, respectively. This process is limited bythe performance of the anion exchanger membrane,which can tolerate only a limited concentration ofacid in the anolyte before back-diffusion of protonsbecomes significant, leading to a decrease in the cur-rent efficiency. Economic aspect of these processeshave been discussed previously [117]. A disadvantageis the high energy consumption resulting from theelectrolysis of the water leading to the undesired by-

Fig. 21. A membrane cell for salt splitting, using a bipolarmembrane.


Fig. 22. CuCl2 recovery from spent solution using electrodialysis combined with continuous electrochemical regeneration of cationand anion exchanger beds.

A review of the technology was given by Ganzi et al.[124]. Three different concepts can be distinguished.The Ionpure (Millipore) and the Electrodiaresis (Euva)concept are employed for de-ionisation of pure water[122,123]. The third concept which is directed towardsthe separation, concentration and recovery of metalions from solution is shown schematically in Fig. 22 forCuCl2 recovery as an example. The feed solution ispassed through separate beds of cation- and an anionexchangers. Under the action of an applied electricalfield, protons and hydroxyl ions, which are producedby electrolysis, are passing the ion exchangermembranes for continuous regeneration of the resinbeds, replacing the Cu2+ and Cl− ions which arereleased into the central chamber of the concentrate.The process works with a current efficiency of 30% andwas tested in a galvanic line for several weeks [122].

5.2. Chloralkali membrane electrolysis

The conventional process for chloralkali electrolysisis the amalgam process [125–127]. This technology isassociated with the emission of mercury into the air andmercury-containing waste waters. With the develop-ment of fluorinated ion exchanger membranes a newprocess became available which represents a muchcleaner technology [128]. In recent years new plantsbased on the membrane technology have replaced theold technology. The new technology also allows toapply the zero gap cell concept which is characterisedby electrodes pressed directly onto the diaphragm inorder to minimise the IR drop caused by the gasevolution. Such membrane cell technology has beencommercialised by AZEC in Japan. Another interestingimprovement to the process has been achieved by ICIwith the FM 21 electrolyser which is a plate and framecell block only 40 cm in height and up to 2 m wide. The

low height of the electrolyser avoids gas accumulationand contributes to further energy savings. The progressin decreasing the specific energy consumption for chlo-ralkali electrolysis by the invention of these new pro-cesses is depicted in Fig. 23.

5.3. Monochloracetic acid production

Monochloracetic acid is produced by chlorination ofacetic acid [129].

CH3COOH+Cl2�CH2ClCOOH+HCl (9)

CH2ClCOOH+Cl2�CHCl2COOH+HCl (10)

Dichloracetic acid is formed as an undesired by-productthrough further chlorination of the main reactionproduct. In the conventional process dichloracetic acidis reconverted into monochloracetic acid by a catalytichydrogenation:CHCl2COOH+H2�

PdCH2ClCOOH+HCl (11)

Fig. 23. Specific energy of different processes for alkali chlo-ride electrolysis.


In a final reaction step, Cl2 is recovered by HClelectrolysis:

2HCl�Cl2+2H++2e− (12)

A more elegant way of recovering monochloracetic acidfrom dichloracetic acid consists in the reversal of thedirection of the side reaction. Reversal of a sponta-neous reaction (DG\0) is only possible by forcedelectrolysis. Reactions (11) and (12) can therefor becarried out in a single electrochemical device with thefollowing anode and cathode reactions:

cathode:

CHCl2COOH+2e−+2H+�CH2ClCOOH+HCl(13)

anode: 2HCl�Cl2+2H++2e− (14)

overall cell reaction:

CHCl2COOH+HCl�CH2ClCOOH+Cl2 (15)

During the development in the laboratory this processhas caused serious corrosion problems of different elec-trode materials like magnetite, graphite and lead. But itwas found that the reaction can be catalysed at graphiteelectrodes in the presence of traces of different metalions (Pb2+, Cu2+, Au3+). By means of such metal ioncatalysis the current density at graphite electrodes couldbe enhanced by a factor of almost 10 without anysignificant corrosion of the graphite electrodes. Thisprocess is a good example to demonstrate how electro-chemistry can contribute to improved production-inte-grated environmental protection. It shows how easy itis to reverse the production of an undesirable follow-upproduct by a forced electrochemical reaction.

6. Conclusions

A number of selected electrochemical processes anddevices, which have been developed in the processindustry in recent years for environmental protection,has been presented and was briefly discussed. Some, butnot all, of them have been tested successfully in thelaboratory or even at a pilot scale, and some havealready reached commercialisation. Due to the specificadvantages in a number of applications, electrochemicalprocesses for the treatment of waste and prevention ofpollution will find increasing acceptance in future devel-opments. In particular, there are many examples whereindustrial processes involve chemical oxidation/reduc-tion steps, which could easily be replaced by a directelectrolysis step thus avoiding costly space, time andenergy consuming separation processes leading to acleaner and process integrated environmentaltechnology.

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