Shellsol D70 Research Articles

Separation and Recovery of Palladium (II) and Platinum (IV) from Automotive Catalysts by Solvent Extraction Using Tri-n-butyl Phosphate and Aliquat 336

Pt, Pd, Fe, Ni, and Cr are among the metals found in spent automotive catalysts, which are both environmentally critical and economically beneficial. In this study, solvent extraction methods were utilized to separate and recover palladium (Pd) and platinum (Pt) from wasted car catalyst. The synthetic leach liquor employed in this investigation contains (mg/L): Pd-140, Pt-500, Ni-800, Fe-2000 and Cr-150. The variation of tri-n-butyl phosphate (TBP) in Shellsol D70 from 0.1 to 2 M revealed selective and quantitative Pd extraction even at the lowest extractant concentration. Stripping Pd from loaded organic (LO) phase using 1 M thiourea and 1 M HCl yielded stripping efficiency of 99%. The use of Aliquat 336 at concentrations ranging from 0.05 to 0.5M in the extraction of Pt from Pd free raffinate demonstrated selective extraction of Pt, reaching 99.7% at 0.2M. The use of a McCabe-Thiele plot in 0.2M Aliquat 336 is expected to result in the extraction of more than 99% of Pt in two stages when an A/O phase ratio of 3 is used. The stripping efficiency of Pt from loaded organic (LO) at O/A ratio 6 with 0.5 M thiourea (tu) and HCl was 99.9 percent.

Conversion of Lithium Chloride into Lithium Hydroxide Using a Two-Step Solvent Extraction Process in an Agitated Kühni Column

A significant consequence of the green transition is the growing demand of lithium-ion batteries (LIBs), as they are essential for electrical vehicles. In turn, the demand for the raw materials that are needed to produce LIBs is increasing. A common LIB cathode type for electrical cars is lithium nickel manganese cobalt oxide (NMC). Since cobalt is currently considered as a critical raw material, nickel-rich NMC cathodes are now designed with lower cobalt contents. The synthesis of these new NMC types requires LiOH instead of Li2CO3, which was used for Co-richer NMC materials in the past. Most production routes of LiOH start from Li2CO3 or Li2SO4. However, LiCl could also be a potential precursor for LiOH, as it could be obtained from various lithium sources. A two-step solvent extraction process (SX) was developed for direct conversion of LiCl into LiOH, using a phenol (butylhydroxytoluene or BHT) and a mixture of quaternary ammonium chlorides (Aliquat 336) in an aliphatic diluent (Shellsol D70) as the solvent. The SX process was validated in counter-current mode using a rotary agitated Kühni extraction column. The use of a column instead of mixer-settlers reduced the CO2 uptake by the final product (LiOH), which prevented the partial conversion of LiOH to Li2CO3. A total of 75 L of LiCl feed solution was processed in the Kühni column to obtain a solution of LiOH with a final purity of more than 99.95%, at a yield of 96%.Graphical

Recovery of Lithium from Simulated Nanofiltration-Treated Seawater Desalination Brine Using Solvent Extraction and Selective Precipitation

ABSTRACT The world's seas and oceans contain vast amounts of lithium, but the low concentration hereof renders solvent extraction impractical for its recovery. By contrast, seawater desalination brine, after treatment by nanofiltration, contains a roughly tenfold greater concentration of lithium than raw seawater. Hence, lithium can be effectively recovered from such streams using solvent extaction. Compared with other techniques to sequester lithium from dilute solutions, solvent extraction offers the advantages of simple operations, robust and well-established technology and high recovery yields. Thus, we propose a solvent-extraction based process to recover lithium from seawater desalination brine, treated by nanofiltration. The first step comprises the removal of magnesium and calcium using methyltrioctylammonium neodecanoate in p-cymene. This is followed by a lithium extraction step using the extractants Mextral 54–100 and Cyanex 923 in Shellsol D70 diluent. The lithium extract is then scrubbed with water and stripped with hydrochloric acid. Subsequently, residual alkaline earth metals are removed with sodium hydroxide in ethanol and finally lithium is precipitated using sodium carbonate. The solvent extraction, scrubbing and stripping steps were demonstrated on mini-pilot scale in continuous countercurrent mode (in mixer-settlers), while the precipitation steps were demonstrated in batch. The process was found to have an overall yield of 74%, affording a lithium carbonate product with a purity of 97 wt%.

Conversion of Lithium Chloride into Lithium Hydroxide by Solvent Extraction

A hydrometallurgical process is described for conversion of an aqueous solution of lithium chloride into an aqueous solution of lithium hydroxide via a chloride/hydroxide anion exchange reaction by solvent extraction. The organic phase comprises a quaternary ammonium chloride and a hydrophobic phenol in a diluent. The best results were observed for a mixture of the quaternary ammonium chloride Aliquat 336 and 2,6-di-tert-butylphenol (1:1 molar ratio) in the aliphatic diluent Shellsol D70. The solvent extraction process involves two steps. In the first step, the organic phase is contacted with an aqueous sodium hydroxide solution. The phenol is deprotonated, and a chloride ion is simultaneously transferred to the aqueous phase, leading to in situ formation of a quaternary ammonium phenolate in the organic phase. The organic phase, comprising the quaternary ammonium phenolate, is contacted in the second step with an aqueous lithium chloride solution. This contact converts the phenolate into the corresponding phenol by protonation with water extracted to the organic phase, followed by a transfer of hydroxide ions to the aqueous phase and chloride ions to the organic phase. As a result, the aqueous lithium chloride solution is transformed into a lithium hydroxide solution. The process has been demonstrated in continuous counter-current mode in mixer–settlers. Solid battery-grade lithium hydroxide monohydrate was obtained from the aqueous solution by crystallization or by antisolvent precipitation with isopropanol. The process consumes no chemicals other than sodium hydroxide. No waste is generated, with the exception of an aqueous sodium chloride solution.Graphical

Solvent extraction equilibrium modeling for the separation of ammonia, nickel(II), and copper(II) from the loaded LIX84-I

A novel solvent extraction equilibrium model has been developed to separate nickel and copper from the loaded LIX84-I (active component: 2-hydroxy-5-nonylacetophenone oxime dissolved in a nonpolar diluent, ShellSol D70) formed by the extraction from ammoniacal alkaline solutions. The targeted process consists of the scrubbing of ammonia and subsequent selective stripping of nickel(II) followed by stripping of copper(II), in which sulfuric acid is used in each step. The present model considers the equilibria of extraction of copper, nickel, and ammonia; distribution of metal-extractant complexes; and solution equilibria in aqueous and organic phases and was compared with the experimental results under varying sulfuric acid concentrations and phase ratios. As a result, this model was shown to accurately predict the equilibria after scrubbing and stripping from the initial operational conditions. This model is also applicable to the extraction of copper and nickel from an ammoniacal solution and thus facilitates the design, optimization, and control of the ammoniacal process for the separation of copper and nickel.

Scale-Up of Membrane-Based Zinc Recovery from Spent Pickling Acids of Hot-Dip Galvanizing.

Zinc recovery from spent pickling acids (SPAs) can play an important role in achieving a circular economy in the galvanizing industry. This work evaluates the scale-up of membrane-based solvent extraction technology aimed at the selective separation of zinc from industrial SPAs as a purification step prior to zinc electrowinning (EW). The experiments were carried out at a pilot scale treating SPAs batches of 57 to 91 L in a non-dispersive solvent extraction (NDSX) configuration that simultaneously performed the extraction and backextraction steps. The pilot plant was equipped with four hollow fiber contactors and 80 m2 of total membrane area, which was approximately 30 times higher than previous bench-scale studies. Tributylphosphate diluted in Shellsol D70 and tap water were used as organic and stripping agents, respectively. Starting with SPAs with high Zn (71.7 ± 4.3 g·L−1) and Fe (82.9 ± 5.0 g·L−1) content, the NDSX process achieved a stripping phase with 55.7 g Zn·L−1 and only 3.2 g Fe·L−1. Other minor metals were not transferred, providing the purified zinc stripping with better quality for the next EW step. A series of five consecutive pilot-scale experiments showed the reproducibility of results, which is an indicator of the stability of the organic extractant and its adequate regeneration in the NDSX operation. Zinc mass transfer fluxes were successfully correlated to zinc concentration in the feed SPA phase, together with data extracted from previous laboratory-scale experiments, allowing us to obtain the design parameter that will enable the leap to the industrial scale. Therefore, the results herein presented demonstrate the NDSX technology in an industrially relevant environment equivalent to TRL 6, which is an essential progress to increase zinc metal resources in the galvanizing sector.

Ammoniacal Solvoleaching of Copper from High-Grade Chrysocolla

The copper silicate ore chrysocolla forms a large potential copper resource, which has not yet been fully exploited, due to difficulties associated with its beneficiation by flotation and metallurgical processing. Direct acid leaching of chrysocolla causes silica gel formation. Therefore, in this work, the feasibility of solvometallurgical methods to leach copper from high-grade chrysocolla while avoiding issues with silica gel formation was assessed. Ammoniacal solvoleaching was performed with a solvent comprising the chelating extractant LIX 984 N or the acidic extractant Versatic acid 10 in an aliphatic diluent (ShellSol D70 or GTL Fluid G70), combined with a small volume of aqueous ammonia. In the three-phase system, aqueous ammonia dissolves copper from milled and sieved chrysocolla, while copper is simultaneously extracted to the organic phase, releasing ammonia that can be reused for further extraction. The best results were obtained with LIX 984 N as extractant: using a 50 vol% LIX 984 N solution, about 75% of copper could be extracted after 60 min of leaching at 25 °C. The stripping of copper from the pregnant leach solution was optimized. Quantitative stripping of copper was achieved with 1.89 M sulfuric acid and the final aqueous solution of copper sulfate had a concentration of 33 g L−1. Experiments in a leaching reactor (1 L) and small battery of mixer-settlers (3 stages, 35 and 143 mL effective volume in the mixer and the settler, respectively, per stage) were successfully conducted and allowed to recover copper with a purity of 99.9%. A conceptual flow sheet has been developed.Graphical

Microvolume Screening of Extraction and Phase Behavior in a Liquid-Liquid Microsystem.

Spontaneous formation of a third immiscible phase during liquid-liquid solvent extraction presents an enormous technical challenge for industry. Insight from current empirical investigations is greatly limited by the lack of methodologies that simultaneously report the progress of the extraction, third-phase onset time, and chemical and physical nature. The microfluidic strategy presented here answers this challenge by supporting an optically transparent submicroliter organic-phase film in a micropillar array surrounded by the aqueous phase. To demonstrate, we used 1 M Cyanex 572 in Shellsol D70 (organic phase) to extract Yb3+ and Dy3+ from a pH 2 aqueous phase. Real-time optical tracking confirmed that the visual onset of third-phase formation is consistent with the cessation of extraction (at the loading limit). Spectroscopic analysis of the solid-like third phase was carried out successfully. The new analytical approach offers a step change in speed and efficiency for reagent development, process control, and fundamental studies of complex phase behavior in reactive multiphase systems.

Preg-robbing verification and prevention in gold chloride-bromide leaching

The use of cyanide is becoming more challenging due to the related environmental risks, increasing the interest in cyanide-free gold leaching. Chloride-bromide leaching is one of the widely investigated alternative lixiviants for gold recovery. However, gold is known to adsorb or precipitate readily on the carbonaceous matter present in the ore, being one of the challenges in chloride process development.The current study presents an equation indicating gold loss due to preg-robbing into solids in chloride-bromide media. The equation was verified using a naturally preg-robbing gold concentrate. It was shown that the graphite naturally present in gold ore can be an equally active preg-robbing agent for gold chloride complex as industrially manufactured and commercially available activated carbon. In order to study preg-robbing prevention during gold leaching in chloride-bromide media, carbon blinding, and Cl2-pretreatment were applied. Blinding with ShellSol D70 was shown to improve gold recovery only slightly whereas Cl2-pretreatment was shown to increase gold recovery from 2% to 45% for a naturally preg-robbing gold concentrate.

Uranium recovery from strong acidic solutions by solvent extraction with Cyanex 923 and a modifier

Uranium stripping with strong acid solution is always highly desired due to its simple operation and less pollution. However, intensive acid neutralisation for uranium precipitation in the subsequent step limited its application. A new solvent extraction process has been developed to transfer uranium from strong to weak sulphuric acid solutions suitable for uranium precipitation without intensive neutralisation. An organic system consisting of 10% Cyanex 923 and 10% isodecanol as the modifier in ShellSol D70 was optimised for the process. It was found that uranium was extracted efficiently from 4 to 6M H2SO4 solutions with the organic system, and it could be efficiently stripped with 0.2–0.5M H2SO4 solutions. Both extraction and stripping kinetics of uranium were very fast, reaching the equilibrium within 0.5min. Temperature between 30 and 60°C has slight effect on uranium extraction and stripping. Four theoretical stages could effectively extract more than 98% uranium from a solution containing 17.5g/L U and 6.0M H2SO4 at an A/O ratio of 1:1.5, and it could generate a loaded organic solution containing about 12g/L U. More than 99% U could be stripped from the loaded organic solution containing 14.6g/L U with 0.5M H2SO4 using five stages at an A/O ratio of 1:3. As a result, the loaded strip liquor containing more than 40g/L U would be obtained which is suitable for uranium recovery by precipitation using hydrogen peroxide. A conceptual process has been proposed for uranium transfer from strong to weak sulphuric acid solutions for its recovery.

Role of diluents in the comparative extraction of Th(IV), U(VI) and other relevant metal ions by DHOA and TBP from nitric acid media and simulated wastes: Reprocessing of U–Th based fuel in perspective

Keeping reprocessing of 233U–Th based fuel in perspective, solvent extraction of Th(IV), U(VI) and some relevant metal ions has been carried out from nitric acid media and simulated high-level waste (HLW) typical of a fast-breeder reactor using dihexyloctanamide (DHOA) in commercially available alternate diluents, viz. Solvesso 100 and ShellSol D70. Acid uptake constant (KH) and distribution ratios (D) of metal ions as a function of nitric acid, thorium ion and extractant concentrations are determined. Slope-analysis results showed various types of species extracted into the organic phase. For comparison, a similar set of experiments were carried out with tri-n-butyl phosphate (TBP) in both diluents. A few clear and interesting trends were observed in the variation of D values of metal ions. Interference of other relevant metal ions likely to get co-extracted along with Th(IV) and U(VI) was also studied by ICP-AES using a simulated HLW.

Separation of lithium from sodium in chloride solutions using SSX systems with LIX 54 and Cyanex 923

A novel synergistic solvent extraction (SSX) system consisting of LIX 54 and Cyanex 923 in ShellSol D70 to recover lithium and separate it from sodium in chloride solutions has been developed. With the SSX system, over 97% of lithium was extracted at pH11 from a synthetic solution with almost no extraction of sodium in a single contact, resulting in a separation factor of lithium over sodium of 1560, indicating very good separation. The co-extracted sodium could be readily scrubbed from the loaded organic solution using water. Lithium was readily stripped with acidic solution and complete strip was achieved at an equilibrium pH of 3.5. The lithium extraction McCabe Thiele diagram showed that three theoretical stages could be used to completely extract lithium from the synthetic solution using an aqueous/organic (A/O) ratio of 2:1. Based on the lithium strip McCabe Thiele diagram, two theoretical stages are required with a 0.5M HCl strip solution at an A/O ratio of 1:3. The lithium extraction and strip kinetics were very fast. The extracted lithium species was found to be LiAB, where A and B represents the deprotonated LIX 54 (A) and Cyanex 923, respectively.

Uranium Solvent Extraction and Separation From Vanadium in Alkaline Solutions

Quaternary amine could be used to extract uranium from alkaline leach solutions, but third phase formation prevented its application. In this paper, isodecanol was used as the modifier for Aliquat 336 in Shellsol D70 to extract uranium from a carbonate leach solution. The formation of third phase was eliminated. More than 98% of uranium was extracted using 3% (w/v) Aliquat 336 and 3% (w/v) isodecanol in Shellsol D70 from a carbonate leach solution containing 95 mg/L U and 25 mg/L V at pH 10.3 and room temperature in a single contact. The separation factor of uranium over vanadium increased with increasing pH of the auqueous solution and reached 280 at pH 11, indicating a good separation of the two metals. The co-extracted vanadium was scrubbed from the loaded organic solution at pH 11 using a solution containing 50 g/L Na2CO3. Over 90% of uranium was stripped from the loaded organic soluiton using an acidic solution containing 150 g/L ammonium sulphate in a single contact. It was also found that to obtain high uranium extraction, the chloride concentration should be controlled to less than 1 g/L in the aqueous phase.

Recovery of scandium from synthetic red mud leach solutions by solvent extraction with D2EHPA

The extraction of scandium from an Australian red mud by selective acid leaching was explored and preliminary leaching tests showed that diluted sulphuric acid can be used to leach scandium from the red mud. The recovery of scandium from a synthetic leach solution of the red mud using solvent extraction was studied. A number of extractants were investigated for the extraction of scandium and its separation from the other metals in the synthetic leach solution. It was found that amongst the three acidic organophosphorus extractants studied, D2EHPA performed best. With the organic system consisting of 0.05M D2EHPA and 0.05M TBP in Shellsol D70 under an A/O ratio of 5:1 at pH 0.25 and 40°C, over 99% scandium was extracted and almost no iron and aluminium were co-extracted. The scandium extracted can be stripped from the D2EHPA/TBP system with 5M NaOH to obtain Sc(OH)3 product. A conceptual flowsheet for the recovery of scandium from red mud is proposed.

Solvent Extraction Equilibrium of Zinc from Nitrate Solutions with 2-Ethylhexylphosphonic Acid Mono-2-ethylhexyl Ester

The solvent extraction equilibrium of zinc with 2-ethylhexylphosphonic acid mono-2-ethylhexyl ester (EHPNA) in an alkane diluent, Shellsol D70, was studied at a constant ionic strength of 1 mol/L from nitrate solutions considering the correction of the extractant nonideality in the organic phase by Alstad’s method. On the basis of the slope analysis of the obtained data, a molar ratio of 1:2 for the zinc to the EHPNA dimer in the extracted complex was indicated. Apparent extraction equilibrium constants obtained at each temperature from 283 to 323 K showed the endothermic nature of the extraction reaction and gave the apparent standard enthalpy and entropy changes. These two constants enable the calculation of the zinc distribution ratio at each temperature which correlates the experimental data with good accuracy.

Modeling of the Extraction Equilibrium of Copper from Sulfate Solutions with Acorga M5640

The extraction equilibrium of copper from sulfate media with the aldoxime Acorga M5640 in ShellSol D70 has been investigated. The distribution results were interpreted by taking into account the nonideality of the aqueous phase. The activity of copper and hydrogen ions in the target systems CuSO4/H2SO4/Na2SO4 and CuSO4/H2SO4/Fe2(SO4)3/ZnSO4 were calculated through the speciation of the aqueous solutions and by applying the Pitzer model. The experimental pH values were found in good agreement with the predicted pH values. A model considering the dimerization of the aldoxime extractant was proposed to predict the distribution ratio and the copper loading isotherms. The extraction constant at infinite dilution and the apparent dimerization constant were evaluated from the experimental data and were found to be 103.06 ± 0.07 and 51 ± 9 M−1, respectively, at 25°C.

Separation and recovery of copper, nickel, cobalt and zinc in chloride solutions by synergistic solvent extraction

The synergistic solvent extraction (SSX) system consisting of LIX63, Versatic 10 and TBP in Shellsol D70 can be used to separate copper, nickel, cobalt and zinc from strong chloride solutions and to transfer copper and nickel to sulphate media for electrowinning.In the first circuit with the SSX system, over 95% Cu was extracted after a single contact, with the co-extraction of some 12% Ni and negligible Co, Mn and Ca. The loaded organic solution could be scrubbed using water to produce the loaded strip liquor suitable for copper conventional electrowinning. In the second circuit with an SX system consisting of Alamine 336 and TBP in Shellsol D70, over 99% Zn, 87% Co and 43% Mn were extracted at pH 3.5 from the copper depleted raffinate after a single contact. The co-extraction of Ni and Ca was negligible. After stripping, the zinc, cobalt and manganese in the loaded strip liquor could be separated from each other using conventional solvent extraction or ion exchange techniques. In the third circuit for nickel with the same SSX system, over 98% Ni was extracted from the copper–cobalt–zinc depleted raffinate in a single contact. The co-extraction of chloride, manganese and calcium was small and was scrubbed. After stripping using spent nickel electrolyte, the loaded strip liquor would be suitable for conventional electrowinning. The nickel extraction and stripping kinetics were reasonably fast for industrial operation. A process flowsheet to recover copper, cobalt, zinc and nickel from strong chloride leach solutions has been proposed.

Equilibrium analysis of solvent extraction of yttrium (III) and europium (III) from hydrochloric acid with P 507

For the accurate prediction of equilibrium distribution ratios of rare earth metals during solvent extraction under non-ideal conditions, the extraction equilibria of yttrium (III) and europium (III) from the hydrochloric acid solutions with P 507 in Shellsol D70 were studied. A chemically-based model was established and the extraction equilibrium constants were determined by the nonlinear least squares method. The proposed model involves the cation exchange reaction and the solvation extraction in the low and high acidity regions, respectively. In the model, the nonideality of the aqueous phase and ( HR ) 2 ¯ was corrected by considering the complexation of the metals with Cl − and by replacing [ ( HR ) 2 ¯ ] with its effective concentration, respectively. This model was verified by fair agreement between the calculated metal distribution ratios and those experimentally obtained in the single metal systems. The feed concentrations for the systems are in wide ranges of the metal (up to 0.1 mol/L), hydrochloric acid (0.07-3.00 mol/L) and the extractant (0.25-1.00 mol/L). The model enables the engineering prediction of the equilibrium distribution ratios with good accuracy in a binary metal system.

Separation of cobalt and zinc from concentrated nickel sulfate solutions with Cyanex 272

AbstractBACKGROUND: Removal of cobalt and zinc from concentrated nickel solutions separately using two Cyanex 272 circuits has been practised in the nickel industry. However, no detailed study has been conducted and data are scarce for further improvement. This study aims to optimise the operating conditions and to simplify the process flowsheet.RESULTS: With a synthetic solution containing 100 g L−1 Ni, 1.4 g L−1 Co and 0.8 g L−1 Zn and the organic solution containing Cyanex 272 and TBP in Shellsol D70, the operating conditions of extraction, scrubbing and stripping were optimised. McCabe–Thiele diagrams were constructed to determine the theoretical extraction and stripping stages and a flowsheet to separate cobalt and zinc from nickel was proposed. With this flowsheet, more than 99% cobalt and zinc could be separated, resulting in a pure nickel solution with less than 10 mg L−1 of cobalt and zinc.CONCLUSIONS: The current study shows that Cyanex 272 can be used to separate cobalt and zinc in one Cyanex 272 circuit effectively from concentrated nickel solutions to obtain very pure nickel solutions suitable for nickel electrowinning or hydrogen reduction. The cobalt and zinc in the loaded strip liquor were concentrated over 10 times and can be separated readily in another much smaller solvent extraction circuit. Copyright © 2010 Society of Chemical Industry

Recovery of nickel and cobalt from laterite leach solutions using direct solvent extraction: Part 2: Semi- and fully-continuous tests

In Part 1 of this paper, two synergistic solvent extraction systems consisting of Versatic 10/LIX63/TBP and Versatic 10/4PC were assessed in batch tests for the separation and purification of nickel and cobalt from synthetic laterite leach solution after iron removal. In Part 2, semi- and fully-continuous tests are reported for the Versatic 10/LIX63/TBP system, with conditions optimised for separating nickel and cobalt from manganese, magnesium and calcium. Semi-continuous extraction tests were conducted using the synergistic organic system consisting of 0.50 M Versatic 10, 0.45 M LIX63 and 1.0 M TBP in Shellsol D70. With a pH profile of 5.5/6.1/6.5 for the three stages EX1/EX2/EX3 at 40 °C, the nickel and cobalt extractions were 99.9% with only 5 mg/L nickel and < 1 mg/L cobalt left in the raffinate. With two stages of scrubbing and a pH profile of 5.4/5.0 at 40 °C, about 2 mg/L manganese and less than 1 mg/L magnesium and calcium were left in the scrubbed organic solution. With two stripping stages and an O/A ratio of 10 at 40 °C using 50 g/L H 2SO 4 as strip solution, the stripping efficiencies of nickel and cobalt were over 95%. A fully-continuous pilot plant was operated for 280 h. With an O/A ratio of about 2 and a pH profile of 5.5/5.8/6.0/6.3 for the four stages EX1/EX2/EX3/EX4 at 40 °C, both nickel and cobalt were almost completely extracted. The nickel and cobalt concentration in the raffinate was lower than detection limit of 0.2 mg/L. The manganese, magnesium and calcium concentrations in the loaded organic solution were 34, 8 and 1 mg/L, respectively. Using a pH profile of 5.4/5.0 for SC1/SC2 at an O/A ratio of 10 and 40 °C, the manganese scrubbing efficiency was over 96% and the concentrations of manganese and magnesium in the scrubbed organic solution were < 5 mg/L and that of calcium 1 mg/L. Using three strip stages and a strip solution containing 50 g/L H 2SO 4 and 55 g/L Ni at an O/A ratio of 10 and 40 °C, over 98% Ni and 99% Co were stripped with only 64 mg/L Ni in the stripped organic solution. The nickel concentration in the loaded strip liquor was 86 g/L, giving a ΔNi of 31 g/L. The loaded strip liquor contained less than 1 g/L acid.