Ž .Hydrometallurgy 56 2000 13–31www.elsevier.nlrlocaterhydromet

Simulation of a SX–EW pilot plant

Hossein Aminian a,), Claude Bazin a, Daniel Hodouin a,Claude Jacob b

a Department of Mining and Metallurgy, LaÕal UniÕersity, Quebec, Canada G1K 7P4b Gaspe DiÕision, Noranda Mines and Exploration, MurdochÕille, Canada G0E 1W0´

Received 20 January 1999; received in revised form 15 December 1999; accepted 17 December 1999

Abstract

Ž .Solvent extraction followed by electrowinning SX–EW is an economical option for theprocessing of low-grade and oxidized copper ore. Phenomenological models were developed tosimulate the copper SX and the EW processes. The two models were linked together to make aSX–EW simulator. The simulator is used to predict the operation of the Mines Gaspe pilot plant´put in operation in 1996. Results confirm the potential of the simulation that would subsequentlybe used for student training, process optimisation and to assess the performance of controlstrategies. q 2000 Published by Elsevier Science B.V. All rights reserved.

Keywords: Solvent extraction; Electrowinning; Modeling; Copper; Iron

1. Introduction

Ž .Solvent extraction followed by electrowinning SX–EW has become in the last twodecades a key process for the recovery of copper from solution obtained by leachinglow-grade copper ore. Several operations have been commissioned in North and South

w x w xAmerica in the last decade 1 . In Canada, Gibraltar Mine 2 operated a SX plantfollowed by EW to produce copper cathode from leached low-grade ore dumps. From1996 to 1997, Mines Gaspe operated a pilot plant to test the economic feasibility of´leaching an oxidized copper ore, followed by SX and EW of copper. Despite the fact

) Corresponding author. Fax: q1-418-656-5343.Ž .E-mail address: haminian@gmn.ulaval.ca H. Aminian .

0304-386Xr00r$ - see front matter q 2000 Published by Elsevier Science B.V. All rights reserved.Ž .PII: S0304-386X 00 00063-3

( )H. Aminian et al.rHydrometallurgy 56 2000 13–3114

that copper SX–EW processes have the reputation of being low-cost operations, plantengineers should always look at ways to make the plant more efficient. Process modelsprovide an economic platform to learn, control and ultimately optimize the operation ofa plant.

w xSeveral models have been independently proposed to simulate the SX process 3–9w xand the EW process 10–12 . However, few papers, if any, deal with the integration of

the two process models into a single simulator of a SX–EW plant. This paper describesthe development of an integrated simulator for the SX–EW process and shows somesteady-state simulation runs of an actual SX–EW pilot plant. The paper consists of foursections. Section 2 shows the general layout and principles of operation of a SX–EWplant. Section 3 describes the SX model while Section 4 deals with the process modelused for the copper EW process. The last section shows how the two process models arelinked to simulate the operation of the Gaspe pilot plant.´

2. The copper SX–EW process

Fig. 1 shows a possible layout of a SX–EW process used to recover copper fromsolutions produced by ore leaching. The process of Fig. 1 consists of two extraction

Fig. 1. Flowsheet of Gaspe SX–EW pilot plant.´

( )H. Aminian et al.rHydrometallurgy 56 2000 13–31 15

stages followed by a stripping stage in closed circuit with an EW unit. The pregnantsolution from the leached ore piles is mixed in a first mixer–settler unit with an organicphase carrying an extractant for copper. The mixer fragments one of the two phases,usually the organic one, into droplets whose surfaces are used as sites for the transfer ofions from the aqueous to the organic phase. The aqueous and organic phases are thenallowed to separate in the settler. The discharged aqueous solution of the first extractionstage feeds a second extraction mixer–settler where it is allowed to contact a stripped

Ž .organic solution see Fig. 1 . The organic and aqueous phases are separated in thesettler. The loaded organic flows into the first extraction stage while the dischargedaqueous solution or raffinate is pumped back to the leach ore pads.

The organic solution from the first extraction stage feeds the stripping mixer–settlerŽ .Fig. 1 where it is mixed with an acidic solution that reverses the extraction process andregenerates the extractant in the organic phase by releasing the extracted copper into theaqueous phase. The aqueous phase, or strong electrolyte, produced by the stripping unitfeeds the EW circuit while the organic is returned to the second extraction stage. Thecopper in the strong electrolyte is plated onto cathodes in the EW circuit and thedischarged solution or spent electrolyte returns to the stripping mixer–settler.

3. SX model

The proposed model describes the steady-state operation of mixer–settler units thatare mostly used in copper SX plants. The proposed model is only concerned with themixer unit, as observations in the studied pilot plant clearly demonstrated that the settlerprovides sufficient residence time to allow the complete separation of the aqueous andorganic phases. Observations also confirm that the transfer of ions between the aqueousand organic phases in the settler is negligible. Future experimental and modeling workwill, however, account for the operation of the settler in order to properly represent thedynamics of the process.

Ž .The extraction process is due to the ion transfer between the aqueous a and organicŽ .o phases described by:

Mqn i qn HRlM R qn Hq 1Ž .i i i ni i

where M stands for the ith extracted species, n the ionic valence and HR the extractanti i

molecule. For instance, the copper extraction reaction is given by:

CUq2 q 2HR l CuR q 2Hq 2Ž . Ž . Ž . Ž . Ž .a o a2 o

Ž q3 .while ferric iron Fe is extracted according to:

Feq3 q 3HR l FeR q 3Hq 3Ž . Ž . Ž . Ž . Ž .a o a3 o

where subscripts a and o stand for aqueous and organic phases, respectively. The modelfor the extraction process should account for the conservation of each species involved

( )H. Aminian et al.rHydrometallurgy 56 2000 13–3116

in the reaction. For that purpose, the species are divided into extracted species andŽ .exchangers, the extractant HR and hydrogen ion being the exchangers in the process.

Assuming a constant volume mixer hold-up, the volume conservation for the aqueousand organic phases is given by:

Q in qQ in sQout qQ in 4Ž .a o a o

Ž in . Ž out.where Q stands for a volumetric flow rate at the feed Q and discharge Q of themixer. Since the transfer of ions has little impact on the phase densities, one can assumethat in steady-state operation:

Q in sQout sQ , Q in sQout sQ 5Ž .a a a o o o

The extracted species mass conservation is given by:

Q x in yQ x out yAf s0 for is1 to n 6Ž .a i a i i

Q y in yQ yout qAf s0 for is1 to n 7Ž .o i o i i

Ž .where x is the species i concentration molrL in the aqueous phase, y the species ii iŽ .concentration in the organic phase molrL and f is the rate of transfer of species ii

from the aqueous to the organic phase expressed as molrmin cm2. The total interfacialsurface, A between the aqueous and organic phases is given by:

6VoAs 8Ž .

dB

where V is the volume of the organic phase in the mixer and d the mean diameter ofo B

spherical organic droplets estimated from the impeller geometry and speed of rotationw xusing the relationship proposed by Godfrey et al. 13 .

The mass conservation equations for Hq and RH are given by:

Nin outQ x yQ x qA n f s0 9Ž .Ýa H a H i i

is1

Nin outQ x yQ x yA n f s0 10Ž .Ýo RH o RH i i

is1

q Ž .where n is the number of RH or H involved in the transfer of species i see Eq. 1 .i

The model for the extraction process accounts for the chemical reaction and the masstransfer that lead to the recovery of species from the aqueous phase to the organic one.

Ž .Fig. 2 shows a schematic view of the extraction process. The aqueous species Mi

diffuse from the aqueous phase to the organicraqueous interface to react with theextractant that has diffused from the bulk organic phase to the interface. The producedhydrogen proton diffuses back to the aqueous phase and the extracted species diffuses

Ž .from the interface to the bulk organic see Fig. 2 .

( )H. Aminian et al.rHydrometallurgy 56 2000 13–31 17

Fig. 2. The extraction process.

After an initial transition period, the rate of transfer of the species from the aqueousphase is balanced with the rate of species consumption by the chemical reaction. If thediffusion process is slow, then the species activities at the interface will be close to thoseof the thermodynamic equilibrium conditions of the extraction reaction. On the otherhand, if the chemical reaction is slow, the species activities at the interface will be closeto those of the bulk solution. In practice, both mechanisms control the process and theinterfacial species activities lie between the thermodynamic equilibrium species activi-ties and the species activities of the organic and aqueous phases.

Ž .The rates of transfer of the species extracted species and hydrogen from or to thebulk aqueous phase are modeled using:

r a sma x yx int 11Ž .Ž .i i i i

r a sma x yx int 12Ž .Ž .H H H H

while the transfer rates in the organic phase are given by:

r o smo y yy int 13Ž .Ž .i i i i

r o smo y yy int 14Ž .Ž .RH RH RH RH

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where m stands for the mass transfer coefficient for species i. The superscript intiŽ .indicates the interface see Fig. 2 .

As the species reach the organicraqueous interface, they react according to Eq. 1.Ž . w xThe rate of forward k reaction for the species is given by 14 :f i

m nint intx yŽ . Ž .i HRfc sk 15Ž .pi f i intxŽ .H

where m, n and p are the partial reaction orders. The relationship between equilibriumw xconstant and rates of reaction is given by 15 :

zniint intk y xŽ .f i i Hz

K s s 16Ž . Ž .i n iint intk ž /r i x yŽ .i RH

Assuming zs1, the rate of the chemical reaction occurring at the droplet interface giw xmay be derived as 8 :

m n niypint int int intx y x yŽ . Ž . Ž . Ž .i RH H ig sk K y 17Ž .pi r i i niyn 1ymint int intž /x y xŽ . Ž . Ž .H RH i

Since at steady-state the overall rate of reaction is equal to the rate of diffusion and tothe rate of the chemical reaction, one can write:

r a sr o sr sg s f 18Ž .i i i i i

N

f s f s n f 19Ž .ÝRH H i iis1

Žwhere f is the overall transfer rate of species i from one phase to the other phase Eqs.i.6, 7, 9 and 10 .

The equilibrium constant, the mass transfer coefficients, and the forward and reverserate constants are estimated from laboratory tests and plant sampling according to a

w xprocedure described elsewhere 16 .The simulation of a mixer–settler unit requires the estimation of the bulk solution

composition by solving simultaneously the mass balance equations and the extractionŽ .equations Eqs. 6, 7, 9 and 10 . The problem is non-linear and an iterative search is used

w xto find the solution. The detailed procedure for the simulation is given by Aminian 16 .The proposed mixer–settler model is valid for the extraction process as well as for thestripping one.

When a SX circuit consists of several interconnected mixer–settler units, as for theŽ .Gaspe process Fig. 1 , an iteration process has to be conducted over the extraction and´

stripping units and the results of each simulation are used to estimate a new recycleorganic phase composition that is re-injected into the circuit feed data and the simulationis repeated until convergence is obtained for the plant.

( )H. Aminian et al.rHydrometallurgy 56 2000 13–31 19

4. EW model

The technical aspects of copper EW from a solution enriched by SX are discussed byw xBiswas and Davenport 17 . The EW cell may consist of up to 40 anode–cathode

couples suspended in a tank through which the copper-rich solution circulates andŽ .overflows at the opposite side of the feed Fig. 3 . The anodes are usually made of lead

Ž . w x Ž .Fig. 3. EW cell. a Industrial EW Cell, 17 . b Anode-cathode pair.

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doped with silver and antimony to reduce the overvoltage of the oxidation reaction. Themain oxidation reaction at the anode is water decomposition:

H O™1r2O q2Hqq2ey 20Ž .2 2

and in the presence of ferrous ions:

Feq2™Feq3 qey 21Ž .

The cathodes are usually made of stainless steel from which the plated copper is easilypeeled off. The main reduction reactions in the winning of copper are:

Cuq2 q2ey™Cu 22Ž .

and the reduction of ferric ions:

Feq3 qey™Feq2 23Ž .

Reduction of hydrogen according to:

2Hqq2ey™H 24Ž .2

is minimum during the winning of copper because of the hydrogen reversible potentialw xand overvoltage 18 .

As for the SX model, the development of the EW model is presented in two sections.The first section lists the material balance equations used to describe the macroscopicoperation of the process, while the second section looks into the mechanisms occurringat the electrodes.

The EW cell shown in Fig. 3 identifies the variables used to write down the materialbalance equations. The mass conservation equation for the solution in the cell is givenby:

Nin in out outQ r yQ r q P yLs0 25Ž .Ý i

is1

where Q is a volumetric flow rate, r the fluid density, P the rate of deposition ori

consumption of species i at the electrodes and L the water losses to the surroundingsdue to evaporation and the mist produced by the oxygen released at the anodes. Models

w xhave been developed to account for such losses 12 , which are however neglected inthis study. The mass conservation equation therefore becomes:

Nin in out outQ r yQ r q P s0 26Ž .Ý i

is1

The density of the solution is estimated from the solution composition using anw xempirical model for a copper sulfate solution 19

10 grLFx F60 grLCurs1018q2.38 x q0.54 xCu H SO2 4 ½10 grLFx F180 grLH SO2 4

where the density and species concentrations are expressed in grL.

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The species mass conservation equations are written as:

Q in x in yQout x out "P s0 27Ž .i i i

where x is the species i concentration expressed as grL in the cell. Tracer testsi

conducted on a pilot electrochemical cell show that the cell behaves closely to a perfectw xmixer 20 , i.e., that the discharged stream has the same composition as the electrolyte in

the tank, so that: x sx out, where x is the species i concentration in the cell. The ratei i i

of formation or consumption of species i is deduced from Faraday’s law:

w c Ii iP s 28Ž .i z Fi

where P is in grmin, w stands for the molecular weight of species i, I the total electrici

current flowing from the anode to the cathode, F the Faraday constant of 96,500Crequivalent mole, z the number of electrons involved in the reaction and c thei i

proportion of the total electric current that contributes to the reaction leading to theformation or consumption of species i. The variable c is referred to the currenti

efficiency if species i is the electro was species.The simulation of an EW cell requires one to predict, for a given cell voltage, the

composition of the discharged solution and the current circulating within the circuit fromthe feed electrolyte composition and flow rate. The methodology followed in this studyassumes that each anode–cathode couple operates as the electrical circuit shown in Fig.4. The total cell voltage, U , is used to push the electric current through the anodic andT

Ž .cathodic busbars and connection resistances R and R , the electrolyte resistanceH,a H,c

Fig. 4. Electrical circuit of copper EW.

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Ž .R and through the reaction branches at the anode and cathode for which the voltages

drops are respectively U and U . Since the sum of the voltage drops should vanish, onea c

can write:

U s IR q IR q IR qU qU 29Ž .T H ,a H ,c s a c

Ž .where I is the total current flow in the circuit. The hardware resistances R and RH,a H,c

are usually maintained as small as possible to optimize the energy utilization. Theelectrical resistance of a copper sulfate solution is calculated using the empirical

w xrelationship 19

10 grLFx F60 grLCuy1R sd 0.134y0.00356 x q0.00249 xŽ .s Cu H SO2 4 ½10 grLFx F180 grLH SO2 4

30Ž .

Ž .where d is the distance cm between the anode and the cathode and R is expressed insŽ .ohms V .

The estimation of the voltage drops at the anode and cathode is the key to thesimulation of the EW process. As it is shown in Fig. 5, several reactions can occur at thecathode and at the anode. Only one electrochemical reaction is first considered todescribe the model used to predict the microscopic transfer and reaction occurring at theelectrode. The model is then extended to incorporate the other electrode reactions.

Fig. 5. Schematic view of an electrode surface.

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Table 1Standard potentials for some reactions

System Voltsq2 yFe q2e ™Fe y0.44

q yH qe ™1r2H 0.02q2 yCu q2e ™Cu q0.34q3 y q2Fe qe ™Fe q0.77q y4H qO q2e ™H O q1.232 2

Fig. 5 shows a schematic view of an electrode surface and of the diffusion layerbetween the surface and the bulk solution. For the presentation, the reaction:

Aqz qzeymA 31Ž .

is considered. In order to proceed, the free energy of reaction 31 has to be negative,otherwise, energy would have to be provided by an external source. The free energy ofan electrochemical reaction is related to its reversible potential calculated using theNernst’s equation:

RT aA ,E0EsE y ln 32Ž .qzF aA ,E

Ž .where R is the gas constant, T the temperature of the solution K , z the number ofelectrons exchanged in the reaction, E0 the standard potential of the reaction for whichthe reactants and products have unit activities and a the species i activity at thei,E

electrode surface. The standard potentials of some reactions are given, with reference tohydrogen which is assumed to have a zero reversible potential, in Table 1. A positivereversible potential implies that the reaction will proceed in the direction shown in Eq.31 without external voltage. The activity of a species at the electrode is estimated fromits molar concentration assuming that Raoult’s law is valid.

If the reversible potential is positive or if the voltage drop at the electrode is largerthan the reversible potential, then the reaction will proceed at a rate that is controlled bythe electron transfer rate at the electrode and the diffusion rate of reactants from thesolution to the electrode and of the products back to the solution.

w xThe rate of electron transfer in molrmin, f , is approximated using 21 :R

z F z F1Ž .o a 1ya a h y 1ya h

qzf s Si a a e ye 33Ž .RT RT½ 5R A ,E A ,EzF

Ž .where a 0Fa-1 is a charge transfer coefficient, S is the area of the electrodeŽ 2 . o Ž 2 .cm , i is the standard current density Arcm , and h the overvoltage given by:

hsU qE 34Ž .E

where U is the voltage drop at the electrode and E the reversible potential.TheE

diffusion rate of reactants to the electrode is given by:

f qsm q x q yx q 35Ž . Ž .D , A A A ,S A ,E

( )H. Aminian et al.rHydrometallurgy 56 2000 13–3124

where m is the mass transfer coefficient of species A and x is the concentration ofA A,S

species A in the bulk solution. The diffusion rate of products from the electrode to thesolution is given by:

f sm x yx 36Ž . Ž .D , A A A ,S A ,E

If the product of the reaction is plated on the electrode surface, there is no rate ofdiffusion back to the electrolyte. For a steady-state operation, considering the stoichiom-etry of Eq. 31, one should have:

f s f s f qs f 37Ž .R D , A D , A

which has to be solved to calculate the species activities at the electrode surface.When several reactions occur at an electrode, the equality of Eq. 27 has to be solved

for each electrochemical reaction. The overall rate of reaction and transfer for eachreaction j, at an electrode is given by:

P z Fj jI s f z Fs 38Ž .j j j wj

where w is the molecular weight of species j, I the current used for reaction j and Pj j jŽ .the rate grmin of production or consumption of species j. The total current passing

though an electrode is given by:

NR

Is I 39Ž .Ý jjs1

and must be equal to the current at the other electrode, i.e.,

N Na c

Is I s I s I s I 40Ž .Ý Ýanodic a , j cathodic c , jjs1 js1

where N and N are respectively the number of reactions occurring at the anode and ata c

the cathode.The EW cell model parameters, mainly the standard current density io, the massj

transfer coefficient m , and the a coefficient for each reaction are calibrated fromj

laboratory experiments as well as from tests carried out in the studied pilot plant. Thecalibration nevertheless remains very approximate, but the objective here is to propose amodel that approximates the actual mechanisms without being too empirical. Prelimi-nary simulation results show that the approach provides sufficient robustness and allowsa larger field of applications than empirical models would do.

The procedure for the simulation of the EW cell involves balancing the electrochemi-cal reactions for a given solution composition and balancing the species conservation inthe cell according to the calculated current flow. The input variables for the simulationare the strong electrolyte feed rate and composition, the cell voltage and the hardware

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resistance. The first iteration assumes that the electrolyte in the cell has the samecomposition as the feed solution. An initial value of the voltage drop at the cathode isthen used to calculate the cathodic current by solving the equality of Eq. 27. Using thecalculated current density the voltage drop at the anode is evaluated by:

U sU yU y I R qR qR 41Ž . Ž .a T c H ,a H ,c s

and the resulting value is used to estimate the anodic current. If the anodic current isdifferent from the cathodic current, the voltage drop at the cathode is modified and theiteration is repeated until the calculated anodic and cathodic currents become equal.

Using the estimated current, the rate of species production and consumption arereadily calculated allowing an estimation of the discharged electrolyte composition.Using the new electrolyte composition, another estimate of the electrical current iscalculated according to the procedure described above. The procedure is repeated untilthe electrolyte composition converges to stable values.

If the EW cells are part of a SX–EW circuit, the procedure for simulation has to berepeated until the complete circuit, i.e., the extraction, stripping and EW stages,converges to final values. Strategies to accelerate the convergence of all the calculationsare currently under study.

5. Model calibration

Model calibration requires estimation of the rate constants k and the mass transferr

coefficients m of the different species for the SX model, as well as the standard current

Fig. 6. Gaspe SX–EW flowsheet.´

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Table 2Ž .Inputs to the simulator leached solution

Ž .Copper grL 1.96Ž .Iron grL 8.54

pH 1.78Ž .Total EW cell voltage V 2

Ž .EW cell temperature 8C 55Ž .Anode–cathode distance m 0.03

2Ž .Cathode surface m 1.8

density io and the a coefficients for EW rate process. The partial reaction orders for thejŽ .SX model are assumed to be one msnsps1 and the relationships obtained by

w x w xFleming et al. 22 and Rod et al. 23 for the mass transfer coefficients for copper SX,i.e.,

m sm s2m s0.1m 42Ž .Cu RH CuR H2

are used to reduce the number of unknown parameters. The equilibrium constants werew xestimated from laboratory SX tests 8 . The rate constants, mass transfer coefficients,

standard current density and a coefficients were estimated from literature data and byusing the results of two sampling campaigns carried out on the studied SX–EW pilotplant. Fig. 6 shows the SX–EW pilot plant with the positions of the collected samplesthat were used to determine the liquid phase flow rates and compositions of thecorresponding streams. Parameter estimation was performed using the SX–EW simula-tor running with the input variables listed in Table 2 and unknown parameters providedby a non-linear optimisation algorithm. The simulator provides values for the outputvariables listed in Table 3, that are used to calculate the following least squares criterion:

2J unknown parameters sÝ y yyŽ . Ž .ˆi i

where y is a measured variable obtained from the plant sampling and y the corre-ˆi i

sponding value generated by the simulator. The estimation procedure is shown in Fig. 7

Table 3Simulated results vs. pilot plant data

Ž . Ž .Stream Cu grL Fe grL

Observed Simulated Observed Simulated

Ž .Feed 1 1.96 1.96 8.54 8.54Ž .Extraction 1 aqueous 2 0.96 0.87 8.62 8.49Ž .Extraction 1 organic 3 2.36 2.06 0.04 0.037Ž . Ž .Extraction 2 aqueous raffinate 4 0.18 0.42 8.66 8.29Ž .Extraction 2 organic 5 1.48 1.37 0.04 0.037

Ž .Strong electrolyte 6 51.7 51.5 1.44 0.032Ž .Spent electrolyte 7 46.7 45.6 1.5 0.032Ž .Stripped organic 8 0.88 0.81 0.04 0.003

( )H. Aminian et al.rHydrometallurgy 56 2000 13–31 27

Fig. 7. Optimisation procedure for calculating the mass transfer and back reaction coefficients.

and the calibrated parameters are given in Table 4. The estimated values of theconsidered output variables are given in Table 3. The observed results are in goodagreement with the simulated data. However, the observed concentrations of iron in thespent and strong electrolytes differ from those of the simulation results. The lowersimulated iron content can be explained by the fact that the iron can be transferred fromthe extraction stage to the stripping mixer–settler by the crud or by the droplets of

Ž .aqueous phase captured in the organic solution of the first extraction stage see Fig. 1 .This phenomenon is not incorporated in the simulator. The transfer of iron from theaqueous to the organic phase in the SX process is minimum as it was shown by

w xlaboratory tests and pilot plant data 8 .

6. Process simulation

The calibrated simulator was used to simulate the SXrEW circuit of Gaspe mines for´operating conditions prevailing during sampling campaigns that were not used for the

( )H. Aminian et al.rHydrometallurgy 56 2000 13–3128

Table 4Estimated model parameters

Parameter Value

( )SX model cmr sy3m 3.77=10Cuy4m 6.14=10Fey3m 3.03=10FeR 3

y1k 4.2=10r Cuy3k 5.8=10rFe

EW modela 0.62Cu

q3a 0.1Fe

q2a 0.1Fe

a 0.2H O2y2 Ž .m 12.2=10 cmrsCu

y2 Ž .q3m 3.01=10 cmrsFey2 Ž .m 2.01=10 cmrsH O2y2 Ž .q2m 4.03=10 cmrsFe

o y3 2Ž .i 2.3=10 ArcmCuo y7 2Ž .q2i 1.2=10 ArcmFeo y6 2Ž .q3i 3.1=10 ArcmFeo y8 2Ž .i 2.06=10 ArcmH O2

model calibration. The input variables are listed in Table 5 and the simulated results arecompared with the experimental ones in Table 6.

The simulation gives acceptable results despite the assumptions made for the masstransfer and rate of the chemical reaction. The simulation predicts that iron will only bemarginally extracted by the organic phase and this result is confirmed by pilot plant andlaboratory experiments.

The validated simulator was used to analyse the influence of different parameterssuch as the concentration of copper in the leached solution, the cell voltage and the ratio

Ž .of phases flow rate of aqueous phaserflow rate of organic phase on the behaviour ofthe process. Fig. 8 shows the effect of the concentration of copper in the leach solution

Table 5Inputs for the simulator validation

Ž .Copper grL 2.22Ž .Iron grL 10.48

pH 1.46Ž .Total EW cell voltage V 2

Ž .EW cell temperature 8C 55Ž .Anode–cathode distance m 0.03

2Ž .Cathode surface m 1.8

( )H. Aminian et al.rHydrometallurgy 56 2000 13–31 29

Table 6Simulated results vs. pilot plant data

Ž . Ž .Stream Cu grL Fe grL

Observed Simulated Observed Simulated

Ž .Feed 1 2.22 2.22 10.48 10.48Ž .Extraction 1 aqueous 2 1.46 1.24 10.4 10.23Ž .Extraction 1 organic 3 1.96 2.03 0.04 0.04Ž . Ž .Extraction 2 aqueous raffinate 4 0.48 0.65 10.38 10.37Ž .Extraction 2 organic 5 1.4 1.38 0.04 0.03

Ž .Strong electrolyte 6 51.4 51.09 0.3 0.03Ž .Spent electrolyte 7 43.6 45.9 0.3 0.03Ž .Stripped organic 8 0.88 0.91 0.51 0.003

on the concentration of copper in the electrolyte and on the current efficiency. Resultsshow that any increase of copper content in the aqueous feed solution of the SX circuitleads to an increase of the copper concentration in the strong and spent electrolyteswhich leads to higher current efficiency. The influence of voltage on the concentrationof copper in the spent electrolyte and on the current efficiency is shown in Fig. 9. Theresults show that with an increase of the voltage, the concentration of copper in the spentelectrolyte decreases while the current efficiency increases. Finally, Fig. 10 shows that

Fig. 8. Influence of copper concentration in the leached solution.

Fig. 9. Influence of cell voltage in the EW circuit.

( )H. Aminian et al.rHydrometallurgy 56 2000 13–3130

Fig.10. Influence of phase ratio on the copper concentration in raffinate.

the concentration of copper in the raffinate increases with the aqueous to organic phaseratio, a result consistent with plant observations.

7. Conclusion

Models for the copper SX and EW processes were developed and calibrated tosimulate an integrated circuit of SX and EW. The unknown parameters of the modelshave been estimated using an optimisation routine operating on the SX–EW simulator.The simulated results agree well with the observed data. Work is in progress to

Ž .incorporate a model for ore leaching LIX to the SX–EW simulator as the raffinate thatis re-circulated to the leached piles may become saturated with iron leading to adifferent behaviour of the process. Once a complete LIX–SX–EW model has beendeveloped and calibrated, it will be used for process simulation, optimisation, design andultimately control. In the meantime, the SX–EW simulator is used to improve ourunderstanding of the process and to study the influence of process variables on thecurrent efficiency and recovery of copper and iron from the leached solutions.

Acknowledgements

The authors are grateful to Noranda, Mines Gaspe for their collaboration and´permission to present the paper. The financial support of the Natural Science andEngineering Research Council of Canada is also acknowledged.

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