Experimental and theoretical studies on a new type of electrochemical reactor for waste-water treatment

Chemical Engineering and Processing. 31 (1992) 195-203

195

Experimental and theoretical studies on a new type of electrochemical reactor for waste-water treatment H. Bergmann, Institute

K. Hertwig

for Chemical Engineering,

and F. Nieber Technical

University

Korthen,

O-4370 Kwthm

(Germany)

(Received November 4, 1991; in final form December 18, 1991)

Abstract In this paper results of experimental investigations and the problems of mathematical modclling for a new electrolytic cell-a drum construction with a vertically moving particle bed electrode-are presented. Advantages of the so-called VMPB reactor, compared with other existing cells, are reported. The new cell is used for the deposition of one or several metals from diluted solutions and can be an alternative to conventional clcctrochemical reactors for waste-water treatment, especially in the pollutant concentration range from 0. I ppm to 20 g of metal per litre. Using a complex non-isothermal reactor model, including balances of mass, heat and electrical flow, mass transfer coefficients are calculated and scale-up tasks rcdhzed.

Introduction Efforts to lower environmental pollution have also led to new developments in reactor design in the field of advanced electrochemistry. The advantage of the combination of a high mass transfer rate with a very large electrode surface has especially favourcd reactors with particle electrodes for processing weakly concentrated liquid wastes contaminated by heavy metals or organic compounds. Such reactors-depending on their working conditions, structure, geometry and orientation of current and mass flow-can be divided into monopolar and bipolar cells, fixed bed and moving bed reactors and cells with cubic, cylindrical, how-by and flowthrough electrodes. Famous examples of industrially used fixed bed cells are the reactors developed by Bennion [ 1, 21 or Kreysa [ 31 (Fig. l(a)). The essential disadvantage of these constructions is the frequent regeneration of the bed as a result of metal plating, so, to

prolong their lifetime, they are only used to operate on low concentrations. Fluidized bed cells [4-61, without or with outer side particle recirculation, are used to avoid this problem [7]. However, because they are uneconomical, few applications are known. Another way of preventing the particles from growing together is the use of rotating drums containing particle electrodes The basic constructions were often taken from electroplating, where small goods are galvanically plated in drum apparatus This can be seen in the reactors proposed by Kammel and Lieber [X] (Fig. l(b)) and in the ‘rolling layer cell’ of Hein and Schab [Y, lo] (Fig. l(c)). The exploitation of the first reactor, however, is often directed to the production of poorly adhesive metal powder, which must be discharged by special means, or to metal recuperation with high current densities and a continuous output of growing particles. Both cells are unsuitable for ultrafine cleaning of wastes down to concentrations lower than 0.5 ppm. The extremely small space-time yield of the rolling layer reactor is a decisive disadvantage. As well as the reactor variants described above, many other cells with particle bed electrodes are found in the literature [ 1 1 - 141. Most of them have never been used in practice.

A new cell concept (a)

(b)

Fig. I. Typical cells for waste-water treatment.

0255-2701/92/$5.00

Cc)

The following requirements the vertically moving particle

(0 1992 -

should be realized with bed (VMPB) reactor:

Elsevier Sequoia. Al1 rights reserved.

196

(I) waste-water treatment in the concentration range 0. I - IO 000 ppm (relating to the metal ions); (2) high space-time yield of the particle drum; (3) continuous transport of particles and electrolyte through the cell. (4) independence of cell behaviour from the effects of evolved gas; (5) low cell voltages and little increase in electrolyte temperature; (6) low manufacturing and exploitation costs. A schematic view of the VMPB reactor is given in Fig. 2. The main part of the reactor is a rotating drum filled with electrode particles, characterized by a drum diameter to electrode thickness ratio of 20:1-3O:l. The back of the drum consists of a rotating disk and a central non-rotating part, through which the particle transport, electrolyte and current feeding are realized. A perforated plate covers the drum on the side opposite the anode. The anode, made from graphite or titanium, is in a box, one side of which consists of a PVC diaphragm to avoid mixing of anolyte and catholyte. Thus the distance between the anode and the cathodic drum is minimized. If hydrogen gas develops, it can easily escape through the perforated plate. This is aided by the fact that the gas evolution zone of the bed lies in the vicinity of the perforated plate. The drum is held in motion by a V-belt or a moving shaft, depending on the technical design. Up to now three independent variations of the VMPB reactor have been created. The parameters of the laboratory cell are given in Table 1. Figure 3 shows the mixing behaviour inside the particle bed for various rates of rotation and degrees of filling of the drum, VF. At low degrees of filling the particle bed behaves as in the rolling layer cell of Hein

TABLE I. Cell parameters Drum diameter Particle bed thickness

Particle size Rate of rotation

0.03 m 0.0015-0.003 tn 1.5-15 min-’

Current density (with respect to separator area) Anode material

10-150Am~Z Coated titanium, graphite

Separator System Acidity Electrical power Electrolysis

Motor, pumps

fixed arafe

ct_

wrticle

CoihOUE

current

fme,-

cutler

Fig. 2. Schema of the VMPB reactor.

Impregnated PVC diaphragm CuSO,/H,SO, 0.01-0.12 mol H,SO, lb’ 2-200

w

300 w

and Schab, whereas for V, = 0.5 the conditions of ideal mixing are nearly achieved. With increase in V, the particle mixing is reduced and only two moving zones, a rolling layer and a zone near the reactor axis, become visible. The angle of inclination of the bed depends on the rate of rotation and increases continuously with rising rotation rate. It can be concluded that the most unfavourable bed behaviour occurs for low degrees of filling, because the concentration profile inside the bed is equalized, which can lead (especially at limiting current conditions, when the process of metal plating can be compared to a first-order reaction) to a lowering of the reaction rate. This disadvantage is compensated by a higher mass transport rate at a filling degree of 0.5. However, more intensified mixing causes other negative side effects of mechanical, chemical and electrochemical particle corrosion. This is why other cell constructions with moving bed electrodes are not sufficiently suitable for ultrafine cleaning of liquid wastes. So we can draw

VF - 0.33, n =.2,5mm”

with

0.3 m

Fig. 3(a).

!+ = 033,

i-,-5m-’

197

0

2

4

6

8

lo

12

14

h

16

tFig. 4. Depletion

vr

=0,5,n-

IOmn”

vf=o.5.

experiment

(V

=

0.035 mX).

the theoretical depletion. From these curves the current efficiency-concentration curve can be obtained (Fig. S), where (p represents the average current efficiency of a bed with certain parameters. For small amounts of copper ion concentration the dependence of current efficiency is nearly linear, which is a sign of limiting current conditions for the process of copper plating. The constant value of (p in the region of higher concentration, where theoretically (p = 1 is expected, is caused by mechanical, chemical and electrochemical side effects inside the moving bed, counteracting the copper deposition, and is typical for rotating drum cells [ 151. It is remarkable that one can work with relatively high current efficiency down to very low electrolyte concentration. And it is also clear that to obtain a very low output or end concentration an extremely small amount of current efficiency must be tolerated, at least in some parts of the bed. In a reactor cascade the last reactor works under poor current efficiency conditions, however, the loss of energy, due to the small necessary reaction current, would be negligible. When using a high electrolysis current at low current efficiency, deviations from the linear Cp-c behaviour can occur. A more

n-t5mmP

(b)

I

t Fig. 3. Study of particle bed mixing behaviour: V, = 0.5; (c) V, = 0.67.

curve for a batch

(a) Vk. = 0.33; (b)

,

I,&’

I

I

I

I

1

’ ’ ’

I%

the conclusion that the most efficient reaction conditions exist for high filling degrees of the drum and low rotation rates.

Results CC”?. -

A typical depletion curve for a certain volume of electrolyte is shown in Fig. 4. The broken line describes

Fig. 5. Average current efficiency as a function of electrolyte concentration with the average current density as parameter.

198 1

0.75 I 1% 0.5

a25

D

Fig. 6. Average

current

efficiencyPelectrolyte

concentration

curve.

detailed picture is shown in Fig. 6. Gas evolution probably influences the mass transport and consequently the current efficiency [ 161. Table 2 contains average amounts A@7/AccUz+in % ppm-’ for various rates of rotation and cell currents (the constant @-values are in parentheses). It can be certified experimentally that a higher rotation rate influences the current efficiency positively, especially at low cell currents. To prevent fast mechanical wear, however, rotation rates below 5 min-’ should be chosen. With increasing cell current the slope of the (p -c curve normally decreases, because the profile of the main current density distribution across the bed thickness does not change significantly, which can be seen indirectly from Table 3 and is demonstrated theoretically later.

TABLE n (mix’)

2.5 4.4 8.0

TABLE VP

0.67 0.67 0.67 0.67

2. Slope of the p-concentration Current

curve

(A)

Table 3 indicates a special dependence of the process efficiency on the rotation rate and the electrolyte acidity. Normally, the current efficiency decreases as the acidity of the electrolyte decreases, because the electrode potentials drift into more negative regions due to the worse conductivity. However, in many experiments it has been shown that for an acidity of approximately 0.02 mol HzSO, 1-l (p increases before falling again as the pH rises. With rising acidity no significant changes in experimental results occur after reaching an acid concentration of 0.06 mol 1-l. The influence of impurities on the current efficiency was investigated in a series of experiments. It appears that chloride ions, citric acid and a special surfactant up to concentrations of 100 ppm do not influence the cp-value.

Reactor modelling

For the determination of parameters and scale-up, detailed modelling is necessary. Only in rare exceptions, if the geometry and process parameters of the laboratory cell correspond to the practical conditions, is a direct scale-up possible. For such a case, for example for a batch electrolytic cell under limiting current conditions, the time-dependent concentration is given by ccu2+

=c$&+ exp( -%)

where k is the slope of the (p-c curve. Equation (1) is an expression analogous to the concentration-time behaviour of a first-order reaction in a batch reactor. In our model the mass and energy balances were connected with the electric flow balance. The following main reactions were considered:

1.25

2.5

5

10

cathode

2.8 (76) 3.9

I.8

0.9 (97) 1.0

0.5 (74) 0.5

Cu”+ + 2e + Cu

(2)

(100) I .9

2H-+2e-+H,

(3)

(66) 1.1 (82)

(55) 0.6 (88)

(78) 4.4 (80)

3. Copper

(76)

deposition

current

for various

conditions

0.06 0.06 0.06 0.02

Hz0 ~ 2e+$O,

of exploitation

+ 2H+

(4)

(ccU2+ = I5 ppm)

1, ( 4

cs

(mol I-‘)

anode

bin-l)

2.5 4.4 8.0 8.0

1.25

2.5

5

10

0.60 0.77 0.88 1.10

0.71 0.80 0.88 1.oo

0.66 0.78 0.90 1.10

0.65 0.80 0.89 1.00

199

With the help of a rotating disk electrode the kinetic behaviour of the anode and cathode was found. The polarization of the cathode was described by addition of the partial current densities for reactions (2) and (3):

The cell voltage

i, = i, + i,

In accordance with the mixing behaviour particle bed the electrolytic cell was modelled ideal mixed reactor in the batch regime,

(5)

The average

current

efficiency

is given by

(p = I, /z

to the equation

uz = u, - u,/, =hK + A@, + A@, + A@+, + A@,, + A@, (18) of the as an

(6)

where the main current I, can be obtained tion across the bed depth b,:

and in continuous

s

(7)

The partial current the cathodic potential

densities i, and i, are functions of at every point of the particle bed:

VsF, b,c

(19)

by integra-

hc I,=-.-.-

corresponds

i,(y) dy

exploitation

by

.,a = 0

The corresponding mc

heat balances

are

p!!T=&+&+hcar

(21)

dr

uk=@,-@r

(8)

The following expressions electrode kinetics:

i _

to describe

the

1

--+$exp(-~,iK.,j

-i,,2exp(-~vK,2)

U, = CT; + aT + b-r In iA

(9)

d2Gp p= IcF dy2

-iE(y)

(10) (11)

(12) FE

(13)

with @FIy-o- - @“F

(14)

‘pPI, =hK = 0

(15)

d% dy

v=bK =

d% dy

=o y=o

0

- To) = 0,

+ 0,

(16) (17)

+ QGas

(22)

The modified reaction heat includes both the reaction heat effects and the heat effects due to ohmic voltage losses: +(1

with the concentration-dependent exchange current densities i,,, , and i,,*, the limiting current density ip’, the overpotentials r~k,, and ?~k,~ according to reactions (2) and (3) the experimentally determined symmetry factors CX,and x2, the equilibrium anode potential U);, the temperature-dependent factors aT and b,, and the anodic current density ia. The equations for the current flux through the bed are d2@, ~ = in(y) FE Icp dv’

and tic,(T

-io,,exp(-f$fkl)

I-

&=

were used

-@)A

Affix

4F

2

-u,

1 z

(23)

where AHK., and AH,,, are the enthalpy changes of reactions (2) and (3), coupled with reaction (4). 0, is the energy flow, caused by heat transfer through the wall, 6,

= G,AD(Tu

- T)

(24)

and dcrs is the heat flow as a result of evaporation and convective gas transport, assuming saturation of hydrogen and oxygen gases. The mathematical model was simulated by a numeric algorithm. The mass transfer coefficient p necessary for ip’ in eqn. (9) can be determined by comparison of the measured slope of the (p-c curve with calculated values for various p. Figure 7 demonstrates this procedure. In the current density range lo- 150 A m-* (relating to the particle contacted area of the perforated plate) for rotation rates between 1 and 8 per minute and a particle diameter of 0.001-0.003 m, mass transfer coefficients between 0.5 and 1.5 x IO-‘rn s-’ were obtained. Figure 8 compares the real depletion behaviour (points) with calculated depletion (curves) for p = 0.5 x lo-’ m s-‘. Knowing p, a precise scale-up becomes possible. The calculation of the current density distribution across the particle bed is an important modelling problem. Figure 9 shows that parts of the bed do not take part in the reaction (i, = 0). At higher cell currents, no significant changes in the i, distribution occur. An

Fig. 9. Main current distribution across the particle bed (V,=O.67, cs=O.Ol rnol I-‘, ccUz+=25ppm. p=1.5x IO-‘ms-‘, d,= 1.5x IO-‘m, b, =0.05m).

(b) Fig. 7. Determination

Fig. 8. Comparison

of the mass transfer

of real and calculated

coefficient

depletion

increase of cell current mainly results in more intensified hydrogen evolution. From Fig. 10 it can be seen that for certain conditions (in general, a low ratio of particle-phase to fluid-phase conductivity, low concentration, low voltage drop across the bed) a relatively high main current density appears across the whole bed, which can lead to metal deposition at the current feeder. Such an effect is also possible during the startup period of the reactor. In times of standstill, surface

Fig. IO. Total rents.

current

density

distribution

for various

cell cur-

oxidation can considerably reduce the effective particle conductivity to values below 10 S m-‘, which can be tested experimentally. Figure 11 shows the total current density distribution for several particle-phase conductivities. In Fig. 12 the average current efficiency is shown as a function of the cell current and pH-value of the catholyte. An example for a non-isothermic modelling is given in Fig. 13. It shows the warm-up curves for a

201

Fig. 13. Predicted increase experiment (V = 0.25 mi).

TABLE

4. Reactor

of electrolyte

cascade

temperature

in a batch

parameters

If

WY XP

5

Fig. Il. Change of current particle-phase conductivity.

distribution

with changing

effective

0.6 m 0.06 m 0.67 0.06 mol H,SO,

Drum diameter Bed depth Degree of filling Acidity Inlet concentration (first reactor)

TABLE 5. Outlet cascade (in ppm)

500 ppm

concentration

d,(mm)

Km s-‘)

Reactor

2.5 2.5 1.5 1.5

1.0 1.5 1.0 1.5

103 93 87 57

x x x x

I-’

of copper

10-S IO-5 10m5 10-S

The influence of particle coefficient is evident.

of copper

1

ions

Reactor 13.0 6.0 5.0 2.5

diameter

2

in the reactor

Reactor

3

4.00 0.40 0.30 0.09

and mass

transfer

Conclusions

Fig. 12. Current efficiency as a function of cell current (average current density) and pH-value (c,-~-~ = 50 ppm, cs = 0.05 mol I-‘, h,=0.04m,~=0.Sx10-5ms~‘).

technical batch cell under normal and extreme current loads. Thus, negative effects on cell material can be predicted. Finally, Table 5 presents scale-up examples for a cascade of three VMPB reactors with the parameters described in Table 4.

With the help of batch experiments for depleting copper containing acid electrolytes, while varying the pH-value, degree of filling, rotation rate of the drum, and the concentration of impurities, it can be shown that the new reactor construction is suitable for wastewater treatment down to metal ion concentrations below 0.5 ppm. A high degree of filling and a rotation rate of less than 5 min-’ were most effective. The influence of impurities is negligible in the concentration range up to 100 ppm. For a scale-up, mathematical modelling becomes necessary. The influence of cell current on current density and electrolyte temperature changes, the

202

choice of optimal bed depth or the influence of particle size and mass transfer conditions can only be quantified by mathematical modelling, as can be shown in simulations of non-isothermal models for technical single cells or reactor cascades. A decisive model parameter, which must be known in advance, is the mass transfer coefficient. Maximal values of 1.5 m s-’ were obtained for p.

Acknowledgements The authors thank W. Schnaubelt (Leipzig) for their generous support work.

and the CAL Ltd. of the development

Nomenclature A

particle contact cross-section of perforated plate, m heat transfer surface, m constant in eqn. (1 I), V depth of bed, m constant in eqn. ( 1 I), V(A m-‘) -I concentration of copper ions, ppm initial, inlet concentration, ppm average heat capacity, J kg-’ K-’ concentration of H,SO,, mol I-’ particle diameter, m Faraday constant, A s mol-’ specific bed surface, m2 mm3 reaction enthalpies, J mol-’ cell current, A main current, A average current density, A mm2 total current density, A mm2 main current density, A mm2 limiting current density of main reaction, A mm2 side current density, A mm2 exchange current density of main and side cathodic reactions, respectively, A mm2 heat transfer coefficient, J m-* K-’ slope of q-c curve, % ppm-’ molar mass of copper, gmol-’ mass, kg mass flow rate, kg SC’ rotation rate, min-’ number of electrons heat transfer rate, J s-’ sum of various heat effects, J s-’ modified reaction heat rate, J s-’ gas constant, J mol-’ K-’

temperature, K environmental temperature, K initial, inlet temperature, K time, s anodic potential, V cathodic potential, V cell voltage, V equilibrium anodic potential, V volume of electrolyte, m3 electrolyte flow rate, m3 s-’ degree of drum filling volume of particles, m3 coordinate, m transfer coefficients ’ mass transfer coefficient, m s-’ overvoltages of main and side cathodic reactions, respectively, V effective liquid conductivity, S m-’ effective particle conductivity, S m-’ = y/b,, dimensionless coordinate potential of liquid phase, V potential of particle phase, V liquid-phase potential at y = 0, V voltage drop across anolyte, V voltage drop across diaphragm, V voltage drop across electrolyte inside particle bed, V voltage drop across perforated plate, V voltage drop across gap between diaphragm and perforated plate, V average current efficiency

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15 E. Avci, Electrolytic recovery of copper from dilute solutions considering environmental measures, J. Appl. Ekirochem., 18 (1988) 288. 16 K. M. Takahashi and R. C. Alkire, Mass transfer in gas sparged porous electrodes, Chem. Eng. Commun.; 38 (1985) 209.