recovery of sulfuric acid from copper tank house electrolyte bleeds

Ž .Hydrometallurgy 56 2000 293–307www.elsevier.nlrlocaterhydromet

Recovery of sulfuric acid from copper tank houseelectrolyte bleeds

Ken Gottliebsen, Baruch Grinbaum, Dehong Chen,Geoff W. Stevens)

Department of Chemical Engineering, The UniÕersity of Melbourne, ParkÕille, VIC 3052, Australia

Received 12 October 1999; received in revised form 10 February 2000; accepted 19 February 2000

Abstract

Many hydrometallurgical processes produce large amounts of acid waste. For example, copperŽ .SXrEW process plants typically bleed a concentrated sulphuric acid stream electrolyte from

their tank house in order to limit the build up of impurities in the electrowinning stage. A processbased on solvent extraction, to selectively recover up to 90% of the sulphuric acid from electrolytebleed streams, has been developed. This is recovered as a pure aqueous acid stream at aconcentration of up to 130 grL that can be recycled back into the tank house circuit thus reducingboth neutralisation and acid make up costs. The acid concentration of the electrolyte is reducedfrom 180 grL to as low as 18 grL. The extraction system involves the use of a branched long

Ž . Ž .chain aliphatic tertiary amine tris 2-ethylhexyl amine TEHA as the extractant, Shellsol 2046 asthe diluent and octanol as a modifier. The equilibrium data and simulation results were alsocompared with an alternative extractant, CYANEX 923. q 2000 Elsevier Science B.V. All rightsreserved.

Keywords: Sulphuric acid; Copper; Tank house

1. Introduction

In the solvent extraction and electrowinning process for the production of copper, aŽ .175–180 grL sulphuric acid solution electrolyte is used to strip the copper from the

loaded organic. It is then sent to the electrowinning cells where copper is electrolyticallydeposited onto stainless steel cathodes. The spent electrolyte from the electrowinning

) Corresponding author. Tel.: q61-93-44-6631; fax: q61-93-44-4153.Ž .E-mail address: [email protected] G.W. Stevens .

0304-386Xr00r$ - see front matter q2000 Elsevier Science B.V. All rights reserved.Ž .PII: S0304-386X 00 00081-5

( )K. Gottliebsen et al.rHydrometallurgy 56 2000 293–307294

Fig. 1. Sulphuric acid recovery process.

tank house is returned to the SX process and used to strip the organic again. Thus theelectrolyte is continuously regenerated. During normal operation, impurities such as ironbuild up in the electrolyte, mainly from aqueous entrainment in the organic. Impuritiessuch as iron affect the current efficiency of the electrowinning process as well as thepurity of the deposited copper. In order to control the concentration of impurities in theelectrolyte, a bleed stream is often required. This bleed has to be neutralised before itcan be released to tails. An equivalent quantity of fresh acid is added to the circuit asmake up to ensure that an acid concentration of 175–180 grL is maintained. If the bleedelectrolyte could be purified and reused in the electrolyte circuit, this would reduce thecosts associated with neutralisation and make up acid.

The proposed recovery process is based on solvent extraction, which has theadvantage of being a well established technology in the mining industry. The processalso does not increase the environmental impact of the copper recovery plant as alladditional chemicals can be completely recycled. A simplified diagram of the process isshown in Fig. 1. The process involves contacting the bleed stream with an organicsolvent which will selectively extract the sulphuric acid, leaving any contaminating ironand other impurities in the electrolyte.

Ž . Ž .The reaction between the tertiary amine extractant, tris 2-ethylhexyl amine TEHAŽ .and sulphuric acid can be described by two equations. The first, represented by Eq. 1

Ž .occurs at low acid concentrations below 1 M and involves the formation of the aminesulphate salt, combining two amines for every sulphate.

H SO q2R N m R NH SO 1Ž . Ž .2 4Žaq . 3 Žorg . 3 4Žorg .2

Ž .The second equation, represented by Eq. 2 involves the formation of aminebisulphate. This occurs at higher concentrations of sulphuric acid, when the concentra-tion of amine sulphate is above 0.02 M

H SO q R NH SO m2 R NH HSO 2Ž . Ž . Ž .2 4Žaq . 3 4Žorg . 3 4Žorg .2

( )K. Gottliebsen et al.rHydrometallurgy 56 2000 293–307 295

The loaded organic solvent is then contacted with pure water at an elevatedtemperature to strip the solvent of acid and produce a pure sulphuric acid stream that canbe recycled back into the copper process. The stripping reaction is simply the reverse of

Ž . Ž .Eqs. 1 and 2The solvent is made up of two other constituents in addition to the TEHA, a diluent

and a modifier. The diluent used in the experiments was Shellsol 2046, which isnecessary to reduce the high viscosity of the loaded solvent. A modifier is required toavoid the formation of a third phase. The third phase is formed at a high solvent loading.Octanol was used as modifier in the initial investigation and it was found that theminimum concentration required to prevent the third phase formation was 20%. Later,tridecanol was used as it has a higher flash point and therefore more cost effective inany industrial application. Tridecanol and octanol demonstrated similar properties asmodifiers.

The conditions, such as extraction and stripping temperature, modifier concentrationand solvent to aqueous phase ratio, have been optimised for the conditions described inFig. 1. These results show that 90% of the sulphuric acid can be recovered.

An alternative to basic extractants, such as amines, is neutral or solvating extractantssuch as liquid phosphine oxides. Two alternative phosphine oxides have been proposedfor the extraction of sulphuric acid. The first being tri-n-octylphosphine oxide withcarbon tetrachloride as the diluent that was tested for the extraction of mineral acids

w xsuch as hydrochloric acid, nitric acid and sulphuric acid 1 .The second is CYANEX 923 which is a mixture of four trialkylphosphine oxides

w x2,3 . Rickelton proposed CYANEX 923 as a possible extractant for the recovery ofsulphuric acid. He found that CYANEX 923 displayed a good compromise between itsability to extract sulphuric acid and to be effectively stripped with water. Rickelton’sexperiments also observed that CYANEX 923 has a very high selectivity for acids inpreference to both copper and nickel. Rickelton’s experiments were performed without adiluent. This extractant has the potential to be used in the recovery of acid.

w xMore recently, Alguacil and Lopez 3 looked at the effect of diluents such as decaneand toluene on the equilibrium of the CYANEX 923 extraction system. They found thatthe diluent did not seem to influence the acid extraction although they observed theformation of a third phase with aliphatic diluents at CYANEX 923 concentrations of10% to 20% and above an aqueous acid concentration of 3 M. It was also found that theextent of extraction decreased at higher temperatures for sulfuric acid.

The extraction of mineral acids by CYANEX 923 under the conditions used in thisŽ .investigation can be represented by the general reaction described by Eq. 3 .

mHqqX myqL mH XL 3Ž .aq aq org m or g

2. Experimental

The equilibrium experiments were performed with two types of technical gradeŽ . w Žextractants: TEHA greater than 95% TEHA, Fluka and CYANEX 923 93% trialkyl-

.phosphine oxide . For TEHA extraction system, solvents were prepared by making 1 M


Ž . Ž .43% solutions of TEHA in varying amounts of octanol and Shellsol 2046 Shell Aust .The concentrations of octanol varied from 14.25% to 57%. For the CYANEX 923system, neither a diluent nor modifier was used. All experiments were done at an ArOratio of 10:1 to ensure that the aqueous acid concentration did not change significantly.Aqueous solutions were made from analytical grade sulphuric acid in distilled water. Allorganic chemicals were used without further purification. Organic and aqueous phases

Ž .were contacted in either separating funnels at ambient temperature or in sealed jarsŽ .placed in a water bath at elevated temperatures and agitated for approximately 3 min to

allow equilibrium to be reached. The mixture was then allowed to settle after which thetwo phases were separated. The organic phase was allowed to stand overnight and wasthen placed in a centrifuge at 140=g for at least 5 min. Any entrainment was removedand the solvent was titrated against a standardised solution of NaOH. All NaOHsolutions were standardised with oxalic acid. The final aqueous phase was also titratedagainst NaOH.

3. Equilibrium data

The process of acid recovery requires a compromise between extraction and strip-ping. Under certain operating conditions, the equilibrium isotherm is such that theloaded solvent cannot be stripped with water. For this reason, it was necessary to obtain

Fig. 2. Equilibrium isotherms for the extraction of H SO using 43% wrw TEHA with octanol as a modifier.2 4


equilibrium isotherms over a wide range of conditions. This would enable optimumoperating conditions to be identified.

3.1. TEHA extraction system

It was found that both temperature and octanol concentration had a significant effectw xon the equilibrium. Previous equilibrium experiments performed by Eyal et al. 4–6

Ž . Ž .Fig. 3. Equilibrium isotherms for the extraction of H SO using 43% TEHA in a 14% wrw octanol, b2 4

28.5% wrw octanol.


show that without the presence of a modifier TEHA extraction of acid is poor. Whenoctanol is included as a modifier, the extraction is significantly improved. Fig. 2 showsclearly the effect of octanol on the equilibrium at constant temperature. The three curvesin Fig. 2 at each octanol concentration represent experiments performed at threedifferent temperatures. Four different octanol concentrations were used: 14.25%, 28.5%,43% and 57% by weight. At 57 wt.%, the octanol was acting as both modifier anddiluent. In the other three cases, Shellsol 2046 was used as a diluent. The best extractionwas observed at 57% octanol. The ability of the solvent to extract the sulphuric aciddecreases as the concentration of the modifier decreases.

A similar effect is also observed with temperature change. The three temperaturesused at each octanol concentration were 208C, 358C and 608C. These temperatures werechosen for the following reasons:

Ø 208C — Ambient laboratory temperature.Ø 358C — Typical ambient temperature at many copper mine sites around AustraliaØ 608C — Elevated temperature chosen to give larger range.

Fig. 3 shows clearly the effect of temperature at two different octanol concentrations.At each octanol concentration, an increase in temperature lowered the ability of thesolvent to extract the sulphuric acid. The effect of temperature was more significant atthe lower octanol concentration. This temperature effect is common in many solventextraction processes.

Fig. 4. Equilibrium isotherms for the extraction of H SO with CYANEXw 923 using both industrially2 4

supplied acid bleed and pure laboratory acid at 208C.


Fig. 5. Comparison between experimental and literature equilibrium isotherms for the extraction of H SO2 4

with CYANEXw 923 using industrially supplied acid bleed solutions.

( )3.2. Liquid phosphine oxide CYANEX 923 system

w xThe aim of this work was to confirm the results obtained by Rickelton 2 whoproposed the use of CYANEX 923 as a possible extractant for the recovery of sulphuricacid. Fig. 4 represents the equilibrium data for the CYANEX 923 system for both pureacid in the aqueous phase and on a solution obtained from an operating plant containing2 grL Fe, 37 grL Cu, as well as as a number of other metal ions. The data shows thatall three isotherms are almost the same, indicating that the large amounts of copper andiron in solution do not have a significant effect on the equilibrium. Fig. 5 compares the

w xresults of Rickelton 2 with extraction and stripping curves. The graph shows that thereŽis no significant difference between the experimental extraction equilibrium data per-

. Ž .formed at 208C and extraction data determined by Rickelton at 408C . However, theŽ .strip isotherm at 508C obtained by Rickelton is significantly different from the

experimental strip isotherm.

4. Modelling the system

The TEHA equilibrium data consisting of 127 points over a range of octanolconcentration, temperature and aqueous acid concentration, was fitted to the following


Table 1Value for parameters in regression model

Parameter Valuey3A 7.5000=10

B 1.0771F 0.71743

y2G 1.3029=10y4H 1.3434=10

J 2.5231K y130.00M y4.4006

y2N 7.7782=10y4P y7.7055=10

equation:

BysAq 4Ž .DX

1q ž /ž /C

This form of equation was chosen because it has an asymptote at higher aqueous acidŽ .concentrations i.e above 2 M . This meant that any extrapolation of the data would not

result in significant error. It was found that parameters A and B did not significantlychange with octanol concentration or temperature. C was a function of both temperatureand octanol concentration while the parameter D is only a function of octanol concentra-tion. By a further curve fitting exercise, relationships for both C and D were chosen

Ž . Ž .according to the Eqs. 6 and 7 .Ž .The final model is shown below Table 1 .

Bw xH SO sAq 5Ž .2 4 sol Dw xH SO aq2 4

1q ž /ž /C

2y6 w x w xCsFqG=Tempq HqJ=10 =Temp P K= Oct q Oct 6Ž .Ž . Ž .2w x w xDsMqN= Oct qP= Oct 7Ž .

w x Ž .where: H SO sconcentration of sulphuric acid in aqueous phase molrkg ;2 4 aqw x Ž . w xH SO sconcentration of sulphuric acid in organic phase molrkg ; Oct s2 4 sol

Ž . Ž .concentration of octanol % wrw ; Temps temperature 8C . This resulted in a fit withan average error of "6%.

5. Process optimisation

Ž . Ž .Eqs. 5 – 7 in addition to steady state material balances around equilibrium stages,were used to determine the optimal process conditions for the extraction of 180 grL


H SO , to produce as high as possible pure sulphuric acid using water as a stripping2 4

agent. The following six parameters were optimised to maximize both the amount ofacid recovered and product acid concentration.

Ø Octanol concentrationØ Temperature of extracting stages

Ž . Ž . Ž .Fig. 6. Effect of octanol concentration on a % H SO recovered, b product H SO concentration grL .2 4 2 4


Ø Temperature of stripping stagesØ Extraction and strip OrA ratioØ Solvent to strip water ratioØ Number of extraction and stripping stages

Ž . Ž .Fig. 7. The effect of the number of extraction stages on a % H SO recovered, b product H SO2 4 2 4Ž .concentration grL .


Fig. 6 shows that both the total amount of H SO recovered and the product H SO2 4 2 4

concentration rise as the concentration of the octanol is decreased. Both these parameterslinearly increase until a maximum is reached at just below 20% octanol. The dotted lineat 20% octanol represents the minimum concentration of modifier to prevent third phaseformation. The trend results from a more stable extracted complex at higher octanol

Ž . Ž .Fig. 8. The effect of the number of strip stages on a % H SO recovered, b product H SO concentration2 4 2 4Ž .grL .


Ž . Ž .Fig. 9. The effect of extraction and strip OrA ratio on a % H SO recovered, b product H SO2 4 2 4Ž .concentration grL .

concentrations, above 20%, making it harder to strip the loaded solvent efficiently withpure water. Twenty percent octanol concentration is close to the level of third phaseformation. Thus, an octanol concentration of around 25% was chosen for this study toensure that there is no chance of a third phase while still producing close to themaximum recovery and product concentration. The effect of temperature differencebetween extraction and strip stages on the optimum octanol concentration is alsorepresented in Fig. 6. Three temperature differences were investigated ranging from 08C

Ž . Ž .Fig. 10. The effect of solvent to strip water ratio on a % H SO recovered, b product H SO concentration2 4 2 4Ž .grL .


Table 2Optimum operating conditions for acid recovery process

Ž .Octanol concentration % 25%Extraction temperature 208CStripping temperature 608CNumber of extraction stages 7Number of stripping stages 9Extraction OrA ratio 2.05Strip OrA ratio 1.78

to 408C. As the temperature difference is increased, so does the optimum octanolconcentration at which the maximum recovery and product concentration occurs.

The number of extraction and strip stages have an effect on the amount of H SO2 4

recovered and the product H SO concentration. Fig. 7 shows the effect of the number2 4

of extraction stages using a solvent with 25% octanol. At low extraction and stripŽ .temperatures 208C , only three extraction stages are required; however, the recovery

and product acid concentration are both very low at 63% and 105 grL, respectively. Ifthe temperature is increased to 408C in both extraction and stripping, a higher percentagerecovery and product acid concentration is achievable. For optimum results at thistemperature, at least five extraction stages are required. When a temperature differencebetween extraction and stripping is introduced, then an even further increase in recoveryand acid concentration is possible. The optimum number of extraction stages in this caseis between six and seven to give a recovery and product concentration of above 80% and135 grL, respectively. Fig. 8 shows that the process by which the acid is stripped fromthe loaded solvent is not nearly as efficient as the extraction of H SO , although the2 4

same trends in temperature exist. Thus, to overcome the inefficient stripping it isnecessary to increase the number of stages in the stripping section to between 8 and 10.

Ž .The organic to aqueous ratio OrA in the extraction and strip stages can beoptimised to maximize amount of acid recovered and product concentration. Fig. 9shows a clear maximum in both percent recovery and product acid concentration at justover an OrA ratio of 2. Fig. 10 shows the effect of the strip OrA ratio when the OrAratio in the extraction stages in maintained at 2. The higher the strip OrA ratio, thehigher the percent recovery which approaches a maximum at a strip OrA ratio of 1.5.However, as the strip OrA ratio is increased, the product concentration decreases. Thisis due to the larger volume of strip water diluting the acid product.

Thus, the optimum operating conditions for the conditions considered are summa-rized in Table 2. This results in a H SO recovery of 90%, a product acid concentration2 4

130 grL and a raffinate acid concentration of 18 grL.

6. Comparison between TEHA and CYANEX 923 extraction system using indus-trial aqueous acid feed

Using this approach, it is also possible compare 1 M TEHA in 25% wrw isotride-canol to CYANEX 923 using industrial acid solutions. The extraction and stripping


Table 3Simulation results for the comparison of TEHA and CYANEX 923

CYANEX 923 TEHA

Modifier concentration No modifier 25% IsotridecanolExtraction temperature 208C 208CStripping temperature 208C 208CNumber of extraction stages 6 6Number of stripping stages 4 6Extraction OrA ratio 1.9 1.9Stripping OrA ratio 1.9 1.9

Ž .Product acid concentration grL 131.7 125.3Percent recovery 79.2 75

equilibrium isotherms for both solvents at 208C were used. The optimum recovery andproduct acid concentration at an operating temperature of 208C are shown in Table 3.

The results show both CYANEX 923 and TEHA could be used for an acid recoveryprocess. Both extractants have the ability to produce a product acid concentration of wellover 120 grL and a percent extraction close to 80% with no iron or copper detectable inthe product. It is further encouraging that these results were obtained without using anelevated stripping temperature which if used could provide even greater productconcentrations and percent recoveries. Both TEHA and CYANEX 923 require furtherpilot plant experiments to fully assess their viability.

7. Conclusions

The following conclusions result from this work:

Ø The modifier concentration has a significant effect on the equilibrium for TEHAextraction system.

Ø The temperature can significantly affect the equilibrium and so alter the optimumrecovery conditions for both TEHA and CYANEX 923. The recovery of the acid issignificantly improved if stripping is performed at an elevated temperature.

Ø Both TEHA and CYANEX 923 extraction systems could be used for an acidrecovery process.

Acknowledgements

Ž .Support from the Australian Research Council and Zeneca Specialties UK for theirfunding of this project is gratefully acknowledged. The authors would also like to thank

Ž .the Advanced Mineral Products Special Research Centre AMPC for their support anduse of equipment and Gunpowder Copper Co for supply of industrial solutions, ShellChemicals Australia for supply of the Shellsol 2046 and Cytec for the CYANEX 923.


References

w x Ž . Ž .1 D.M. Petkovic, M.M. Kopeni, A.A. Mltrovic, Solvent Extraction and Ion Exchange 10 4 1992685–696.

w x Ž .2 W.A. Rickelton, ISEC’93 2 1993 731–734.w x Ž .3 F.J Alguacil, F.A. Lopez, Hydrometallurgy 42 1996 245–255.w x Ž .4 A.M. Eyal, A.M. Baniel, Solvent Extraction and Ion Exchange 9 1991 195–210.w x Ž .5 A.M. Eyal, B. Hazan, R. Bloch, Solvent Extraction and Ion Exchange 9 1991 211–222.w x Ž .6 A.M. Eyal, B. Hazan, R. Bloch, Solvent Extraction and Ion Exchange 9 1991 223–236.

recovery of sulfuric acid from copper tank house electrolyte bleeds

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