WO2021228936A1 - Method for producing battery grade lithium hydroxide monohydrate - Google Patents

Method for producing battery grade lithium hydroxide monohydrate Download PDF

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WO2021228936A1
WO2021228936A1 PCT/EP2021/062621 EP2021062621W WO2021228936A1 WO 2021228936 A1 WO2021228936 A1 WO 2021228936A1 EP 2021062621 W EP2021062621 W EP 2021062621W WO 2021228936 A1 WO2021228936 A1 WO 2021228936A1
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chloride
lithium
aqueous
hydroxide
solution
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French (fr)
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Koen Binnemans
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Katholieke Universiteit Leuven
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    • CCHEMISTRY; METALLURGY
    • C01INORGANIC CHEMISTRY
    • C01DCOMPOUNDS OF ALKALI METALS, i.e. LITHIUM, SODIUM, POTASSIUM, RUBIDIUM, CAESIUM, OR FRANCIUM
    • C01D3/00Halides of sodium, potassium or alkali metals in general
    • C01D3/14Purification
    • CCHEMISTRY; METALLURGY
    • C01INORGANIC CHEMISTRY
    • C01DCOMPOUNDS OF ALKALI METALS, i.e. LITHIUM, SODIUM, POTASSIUM, RUBIDIUM, CAESIUM, OR FRANCIUM
    • C01D15/00Lithium compounds
    • C01D15/04Halides
    • CCHEMISTRY; METALLURGY
    • C01INORGANIC CHEMISTRY
    • C01DCOMPOUNDS OF ALKALI METALS, i.e. LITHIUM, SODIUM, POTASSIUM, RUBIDIUM, CAESIUM, OR FRANCIUM
    • C01D3/00Halides of sodium, potassium or alkali metals in general
    • C01D3/04Chlorides
    • CCHEMISTRY; METALLURGY
    • C01INORGANIC CHEMISTRY
    • C01DCOMPOUNDS OF ALKALI METALS, i.e. LITHIUM, SODIUM, POTASSIUM, RUBIDIUM, CAESIUM, OR FRANCIUM
    • C01D3/00Halides of sodium, potassium or alkali metals in general
    • C01D3/04Chlorides
    • C01D3/06Preparation by working up brines; seawater or spent lyes
    • CCHEMISTRY; METALLURGY
    • C01INORGANIC CHEMISTRY
    • C01DCOMPOUNDS OF ALKALI METALS, i.e. LITHIUM, SODIUM, POTASSIUM, RUBIDIUM, CAESIUM, OR FRANCIUM
    • C01D3/00Halides of sodium, potassium or alkali metals in general
    • C01D3/04Chlorides
    • C01D3/08Preparation by working up natural or industrial salt mixtures or siliceous minerals

Abstract

The invention relates to processes for refining lithium chloride comprising the steps of : 1) dissolving a solid composition comprising less than 98 % (w/w) lithium chloride in an organic solvent thereby selectively dissolving lithium chloride; 2) adding an alkali metal hydroxide solubilised in an organic solvent; 3) optionally further purifying lithium chloride in the organic solution by non-aqueous ion exchange; and 4) removing the organic solvent.

Description

METHOD FOR PRODUCING BATTERY GRADE LITHIUM HYDROXIDE MONOHYDRATE
FIELD OF THE INVENTION The present invention relates to the lithium industry. A process is disclosed for the production of high-purity lithium hydroxide monohydrate using solid lithium chloride as input stream. The high-purity lithium hydroxide monohydrate is of battery-grade and can be used for the manufacturing of cathode materials for lithium-ion batteries, with particular attention for nickel-rich NMC chemistries (e.g. NMC622 and NMC811). The invention includes a solvent extraction process for the conversion of lithium chloride into lithium hydroxide by chloride/hydroxide anion exchange.
BACKGROUND OF THE INVENTION
Lithium is an indispensable raw material for lithium-ion batteries (LIBs), where is it being used in the electrolyte (e.g. LiPFe) and in cathode materials (e.g. LC0O2, NMC, LiFePCU). The main lithium sources are brines from salt lakes (sa/ars) and hard rock lithium ores (pegmatites containing spodumene, petalite, lepidolite, amblygonite, triphylite, zinnwaldite and other lithium minerals). Geothermal brines are another potential lithium source. Commercially available lithium compounds are lithium carbonate (U2CO3), lithium chloride (LiCI), lithium hydroxide (LiOH) and lithium hydroxide monohydrate (UOI-M-I2O). Battery-grade lithium salts are required for the manufacturing of lithium-ion batteries. Battery-grade lithium carbonate and lithium hydroxide monohydrate have a purity of at least 99.5%, and often even higher. Lithium is mixed in brines with large quantities of sodium, as well as smaller or larger amounts of potassium, magnesium and calcium.. The largest proportion of anions comprises the chloride ion (Cl ), with minor amounts of carbonate (CO32 ), sulphate (SO42 ) and nitrate (NO3 ) ions. Also boron compounds (boric acid and borates) are present. Lithium recovery from brines is mainly operated in South America (Lithium Triangle in Chile, Argentina and Bolivia), as well as in PR China (Qinghai-Tibet Plateau). The brine is pumped up from beneath the crust on a salt lake to large open- air, shallow evaporation ponds. Lithium brine is concentrated by solar evaporation and wind to a concentration of approx. 6000 ppm Li. Several evaporation steps in different ponds precipitate sodium, potassium and magnesium salts. The concentrated brine is pumped into the recovery plant. Boron is removed by solvent extraction and magnesium by precipitation of magnesium hydroxide, Mg(OH)2, with lime (CaO) and sodium hydroxide (NaOH). The concentrated brine is further treated by sodium carbonate (NazCCh) to first precipitate calcium as calcium carbonate (CaCCh) and then to precipitate lithium as lithium carbonate. Technical-grade lithium carbonate is further purified by several redissolution and reprecipitation steps. Purification of technical lithium carbonate to battery-grade is conventionally done by the carbonation process. Technical-grade lithium carbonate is subjected to carbonation with pressurized carbon dioxide gas (CO2) to solubilize lithium as lithium hydrogencarbonate (UHCO3). Lithium hydrogencarbonate is then heated to 90 °C to precipitate lithium as lithium carbonate with a purity of 99.95%.
At present, lithium carbonate is the most important lithium compound for lithium-ion battery applications, but the importance of lithium hydroxide is rapidly increasing, especially because of the rising importance of nickel-rich NMC cathode materials (NMC = nickel-manganese-cobalt) that are prepared starting from lithium hydroxide. If the metals are used in a ratio of 6 parts nickel to 2 parts cobalt and 2 parts manganese (6-2-2), or 8-1-1, rather than 1-1-1 or 5-3-2 as in the past, the synthesis requires lithium hydroxide rather than lithium carbonate. The higher temperature required to synthesize cathode materials comprising more than 60% nickel with lithium carbonate damages the crystal structure of the cathode and changes the oxidation state of nickel. However, lithium hydroxide allows fast and complete synthesis at lower temperatures. The change in demand is shifting mining projects towards developing lithium hydroxide production rather than lithium carbonate. Industrial production of lithium hydroxide is conventionally conducted in a number of ways including: 1) reaction of lithium carbonate with Ca(OH)2 and rejection of CaCCh (e.g. US patent US4,207,297), 2) reaction of U2SO4 with NaOH and rejection of Na2SC>4, 3) direct conversion of a-spodumene with Ca(OH)2.
A process for direct transformation of lithium chloride into the important precursor for nickel-rich cathode materials lithium hydroxide monohydrate, bypassing the lithium carbonate intermediate, would be advantageous from an economical and environmental point of view because this reduces the number of processing steps, leads to the consumption of less chemicals and generates less waste. However, such a process implies the preceding purification of technical-grade lithium chloride into battery-grade lithium chloride.
Lithium chloride has an exceptionally high solubility in polar organic solvents, and particularly in alcohols. This behaviour distinguishes lithium chloride from other alkali chloride salts such as sodium chloride (NaCI) and potassium chloride (KCI). For instance, at 20 °C, the solubility of lithium chloride is 43.8 g/100 g methanol and 24.3 g/100 g ethanol, whereas the solubility of sodium chloride is 1.41 g/100 g methanol and 0.07 g/100 g ethanol.
Several organic-solvent-based processes for lithium chloride purification have been disclosed. US patent US3,278,260 discloses the use of alcohols to purify lithium chloride obtained from spodumene ore. US patent US3,410,653 describes the separation of lithium salts from other alkali salts by selective dissolution of lithium salts in alcohols. US patent US4,274,834 discloses the extraction of lithium chloride from other inorganic salts by isopropanol. US patent US4,980,136 describes a process in which lithium chloride is purified with selective dissolution in isopropanol. However, these methods fail to remove calcium and magnesium impurities, because magnesium chloride and calcium chloride have a significant solubility in alcohol solvents. For instance, the solubility of MgCh is 7.4 g in 100 mL of ethanol at 30 °C, and the solubility of CaCh is 25.8 g in 100 g of ethanol at 20 °C.
Once battery-grade lithium chloride is obtained, it is converted into lithium hydroxide to obtain the targeted end product lithium hydroxide monohydrate. This can be achieved by electrolysis in a membrane cell via a process that is similar to the production of sodium hydroxide from sodium chloride (US patents US4,036,713 and US20110044882). Typically, a fluorinated cation-exchange membrane with sulfonic acid groups is being used (e.g. Nafion membrane). The main drawbacks of the membrane electrolysis process are the high price of membranes, the loss in energy efficiency due to the internal resistance of the ion-exchange membrane and the fact that the membranes are susceptible to fouling and scaling. When lithium chloride is used as electrolyte, toxic chlorine gas is released at the anode. Chlorine gas can also attack the membrane. An approach similar to membrane electrolysis is electrodialysis. The same drawbacks of membrane electrolysis hold for electrodialysis.
Lithium chloride can be converted into lithium hydroxide by a strong basic anion exchanger in the hydroxide form. Because of the low atomic mass of lithium (6.94 g mol 1), a large ion exchange capacity is required to prepare 1 kg of lithium hydroxide starting from lithium chloride by using an anion exchanger, implying that this method is not efficient.
Direct chloride-to-hydroxide conversion by solvent extraction or a similar technique would have several advantages with respect to membrane electrolysis or ion exchange for the direct conversion of lithium chloride to lithium hydroxide. However, chloride/hydroxide anion exchange via solvent extraction is challenging because of difficulties in preparing a solvent phase with a basic extractant (liquid anion exchanger) in the hydroxide form and the limited stability of such compounds due to Hoffman elimination [Dehmlow et at. (1985) Tetrahedron 41, 2927-2932]. Rothenberg et at. (2000) Chemical Communications 1293-1294, Lavi et at. (2001) Industrial & Engineering Chemistry Research 40 6045-6050 and Israelian patent IL133622 disclose a process for conversion of metal chlorides and onium chlorides into the corresponding hydroxide salts by a liquid membrane process comprising an onium chloride and an aliphatic alcohol in hexane. The mechanism is described as transfer of hydroxide ions across the liquid membrane. The authors state that the very weak acidity of the alcohol is essential (pKa > 16), and that extraction efficiency depends on the type of alcohol, in the order: diols > primary alcohols > secondary alcohols. The authors mention that the driving force is the concentration gradient, so that they use a very high sodium hydroxide concentration in the aqueous phase in the first exchange step (at least 2.5 mol kg 1, but preferably 10 to 12.5 mol kg 1). This process has been demonstrated for conversion of potassium chloride to potassium hydroxide, but it was found to be inefficient for conversion of lithium into lithium hydroxide: after seven contact cycles, only 65% of the chloride ions had been replaced by hydroxide ions in a lithium chloride solution (21% w/w, 5 mol kg 1), by using a 40% w/w (10 mol kg 1) sodium hydroxide solution.
SUMMARY OF THE INVENTION
Methods for the production of high-purity lithium hydroxide monohydrate are provided using solid technical-grade lithium chloride as input stream. The high-purity lithium hydroxide monohydrate is of battery-grade and can be used for the manufacturing of cathode materials for lithium-ion batteries, with particular attention to nickel-rich NMC chemistries (e.g. NMC622 and NMC811). The invention includes a solvent extraction process for conversion of lithium chloride into lithium hydroxide by chloride/hydroxide anion exchange. The process comprises the following steps: (1) dissolving solid technical-grade lithium chloride in an organic solvent to obtain a pre purified lithium chloride solution since the solubility of lithium chloride in organic solvents is significantly higher than the solubility of sodium chloride or potassium chloride; (2) addition of a solution of an alkali metal hydroxide in an organic solvent; (3) (optionally) purifying the organic solution by non-aqueous ion exchange; (4) removing the organic solvent from the purified lithium chloride, followed by dissolution of the lithium chloride in water; (5) converting lithium chloride to lithium hydroxide by solvent extraction; (6) crystallizing battery-grade lithium hydroxide monohydrate from the aqueous lithium hydroxide solution. A general flowsheet of the process is shown in Figure 1. The process is widely applicable and can prepare lithium hydroxide monohydrate from different lithium sources, including from salt lake brines (salars), geothermal brines, hard rock lithium ores (pegmatites containing spodumene, petalite, lepidolite, amblygonite, triphylite, zinnwaldite and other lithium minerals), as well as lithium-rich slags produced by pyrometallurgical processes for recycling of lithium-ion batteries. The only restriction is that solid technical-grade lithium chloride is available as starting material for the process. Lithium compounds other than lithium chloride can be processed, but need to be converted first into lithium chloride. Concurrently, the process allows for conversion of lithium carbonate into lithium hydroxide monohydrate.
The process described in this invention can be split into two separate processes that can be operated independently: (1) process A to purify solid technical-grade lithium chloride into solid battery-grade lithium chloride, and (2) process B to convert an aqueous solution of lithium chloride into an aqueous solution of lithium hydroxide by solvent extraction, followed by crystallizing battery-grade lithium hydroxide monohydrate from the aqueous solution. Whereas process A uses solid lithium chloride as input material, process B can use both solid lithium chloride or an aqueous lithium chloride solution as input material. Process B can also start from battery- grade lithium carbonate, after dissolution of this compound in a hydrochloric acid solution.
The invention is further summarized in the following statements:
1. A process for refining lithium chloride comprising the steps of :
1) dissolving a solid composition comprising less than 98 % (w/w) lithium chloride in an organic solvent thereby selectively dissolving lithium chloride over sodium chloride and/or potassium chloride which have a lower solubility in organic solvents than lithium chloride ;
2) adding an alkali metal hydroxide solubilised in an organic solvent thereby converting magnesium chloride into magnesium hydroxide and converting calcium chloride into calcium hydroxide; in a preferred embodiment the alkali metal hydroxide is lithium hydroxide;
3) optionally further purifying lithium chloride in the organic solution by non- aqueous ion exchange thereby removing sulphate that was not preciptated in step 1; and
4) removing the organic solvent. 2. The process according to statement 1, further comprising the steps of:
5) converting an aqueous solution comprising lithium chloride obtained in step c) into an aqueous solution comprising lithium hydroxide by solvent extraction with a mixture of a basic extractant in the chloride form and a weakly acid compound.
3. The method according to statement 2, further comprising the step of:
6) crystallizing lithium hydroxide monohydrate from the aqueous lithium hydroxide solution, or converting lithium hydroxide in the aqueous solution to lithium carbonate by reaction with carbon dioxide gas.
4. The process according to any one of statements 1 to 3, wherein step 4 comprises:
- recovering lithium chloride by evaporation of the solvent of the organic phase and crystallization of lithium chloride from the aqueous phase, or
- recovering lithium chloride aqueous phase by evaporation of the solvent of the organic phase and drying the aqueous phase comprising lithium chloride.
5. The process according to any one of statements 1 to 4, where the composition comprising lithium chloride of step 1) is obtained from salt lake brines, from geothermal brines, or from a lithium ore.
6. The process according to statement 5, where the lithium ore is selected from the group consisting of spodumene, petalite, lepidolite, amblygonite and zinnwaldite.
7. The process according to any one of statements 1 to 6, wherein the composition comprising lithium chloride of step 1) is obtained from spodumene by calcium chloride roasting, or by chlorination roasting.
8. The process according to any one of statements 1 to 6, where the composition comprising lithium chloride of step 1) is obtained from spodumene by sodium carbonate pressure leaching, followed by conversion of lithium carbonate in lithium chloride by hydrochloric acid.
9. The process according to any one of statements 1 to 6, where the composition comprising lithium chloride of step 1) is obtained from spodumene by the reaction sequence of decrepitation, sulfuric acid roasting, dissolution of lithium sulphate in water, precipitation of lithium carbonate by sodium carbonate, followed by conversion of lithium carbonate in lithium chloride by hydrochloric acid.
10. The process according to any one of statements 1 to 6, where the composition comprising lithium chloride of step 1) is obtained from spodumene by the reaction sequence of decrepitation, sulfuric acid roasting, and metathesis reaction with sodium chloride or potassium chloride to convert lithium sulphate into lithium chloride.
11. The process according to any one of statements 1 to 6, where the composition comprising lithium chloride of step 1) is from lithium-rich metallurgical slag from pyrometallurgical recycling processes of lithium-ion batteries.
12. The process according to any one of statements 1 to 12, where the organic solvent in step 1) is an alcohol, wherein the alcohol may comprise up to 5 or 10% (v/v) water.
13. The process according to any one of statements 1 to 11, wherein the organic solvent in step 1) is an alcohol selected from the group consisting of methanol, ethanol, isopropanol, 2-ethylhexanol, and 1-octanol A preferred organic solvent is 95% (v/v) ethanol.
14. The process according to any one of statements 1 to 113, where said non- aqueous ion exchange in step 4) is performed with a cation exchanger.
15. The process according to any one of statement 1 to 13 where said non-aqueous ion exchange in step 4) is performed with a strongly acidic ion exchange resin of cross-linked polystyrene with sulfonic acid functional groups.
16. The process according to any one of statements 1 to 13, where said non-aqueous ion exchange in step 4) is performed with an inorganic ion-exchange material.
17. The process according to statement 16, where the inorganic ion-exchange material ion exchanger is (poly)antimonic acid.
18. The process according to any one of statements 1 to 13, where the non-aqueous ion exchange in step 4) is performed with an ionic exchanger in the H+ form or in the Li+ form.
19. The process according to any one of statements 1 to 13, where the non-aqueous ion exchange in step 4) is performed with a weakly basic ion exchange resin of cross-linked polystyrene with tertiary amine functional groups.
20. The process according to any one of statement 2 to 19, where the solvent extraction for the chloride-hydroxide anion exchange of step 5) is carried out by a mixture of a quaternary ammonium chloride and an alcohol or phenol or a derivate thereof in a diluent.
21. The process according to statement 20, where the quaternary ammonium compound is tricaprylmethylammonium chloride.
22. The process according to statement 20 or 21, where the alcohol is selected from the group consisting of 1-pentanol, 1-hexanol, 1-heptanol, 1-octanol, 1-decanol, 1-dodecanol, 1-tetradecanol 2-pentanol, 2-hexanol, 2-octanol, cyclohexanol and
2-ethylhexanol.
23. The process statement according to statement 20 or 21, where the alcohol is a diol is selected from the group consisting of 2, 2-dimethyl-l, 3-propanediol (neopentyl glycol), 2-methyl-24-pentanediol, 2,3-dimethyl-2,3-butanediol (pinacol), 2,5-dimethyl-2,5-hexanediol, 2-methyl-2-propyl-l, 3-propanediol, 2,2,4-trimethyl-l,3-pentanediol, 2-ethyl-l,3-hexanediol, 1,2-octanediol, 1,2- pentanediol and 1,2-hexanediol.
24. The process according to statement 23, where the diol is 2,5-dimethyl-2,5- hexanediol.
25. The process according to any one of statements 20 to 24, where the phenol derivative is selected from the group consisting of 4-tert-butylphenol, 2,4-di- tert-butylphenol, 2,6-di-tert-butyl-4-methylphenol, 2,6-di-tert-butylphenol, 2,6- dimethylphenol, 3,5-dimethylphenol, 2,4,6-tri-tert-butylphenol, 4-nonylphenol, nonylphenol (isomeric mixture) and cresol.
26. The process according to any one of statements 20 to 25, where the phenol derivative is 4-tert-butylphenol or an isomeric mixture of nonylphenol.
27. The process according to any one of statements 20 to 26, where the solvent extraction process for the chloride-hydroxide anion exchange is carried out by a mixture of tricaprylmethylammonium chloride and 2,5-dimethyl-2,5-hexanediol in a diluent.
28. The process according to any one of statements 20 to 26, where the solvent extraction for the chloride-hydroxide anion exchange of step 4) is carried out by a mixture of tricaprylmethylammonium chloride and 4-tert-butylphenol in a diluent, or by a mixture of tricaprylmethylammonium chloride and nonylphenol (isomeric mixture) in a diluent .
29. The process according to statement 2, further comprising the step of recovering lithium hydroxide from the aqueous solution by antisolvent precipitation with an alcohol, such as isopropanol
30. A process for refining lithium chloride comprising the steps of : la) dissolution a composition comprising lithium carbonate in hydrochloric acid thereby obtaining lithium chloride
2a) removing the organic solvent from the purified lithium chloride, optionally with recovery of the organic solvent. Specific embodiments applicable to step 4 of the process of statement 3 are equally applicable to step 2a of the method of process 30. 31. The process according to statement 30, wherein the composition comprising lithium carbonate of step 1) has a purity of at least 99.5% (w/w).
32. A process comprising a two-step solvent extraction where in the first step a solvent comprising a mixture of a quaternary ammonium chloride and an alcohol/phenol is contacted with a concentrated sodium hydroxide solution to form in-situ a quaternary ammonium alkoxide/phenolate, followed in the second step by contact between the solvent with the quaternary ammonium alkoxide/phenolate and an aqueous solution of an alkali chloride or an onium chloride to convert the chloride salt into the corresponding hydroxide.
33. The process according to statement 30, where the alkali chloride is lithium chloride, potassium chloride, rubidium chloride or caesium chloride.
34. The process according to statement 30 or 31, where the onium chloride is tetramethylammonium chloride or tetraethylammonium chloride.
DETAILED DESCRIPTION OF THE INVENTION BRIEF DESCRIPTION OF THE FIGURES
Figure 1: Flowsheet showing the different steps of the invention: (1) dissolving technical-grade lithium chloride in an organic solvent to obtain a pre-purified lithium chloride solution; (2) addition of a solution of an alkali metal hydroxide in an organic solvent; (3) purifying the organic solution by non-aqueous ion exchange; (4) removal of the organic solvent from the purified lithium chloride, followed by dissolution of the lithium chloride in water; (5) converting lithium chloride to lithium hydroxide by solvent extraction; (6) crystallizing battery-grade lithium hydroxide monohydrate from the aqueous lithium hydroxide solution.
Figure 2: Flowsheet for the conversion of lithium chloride into lithium hydroxide by solvent extraction with mixer-settlers (MS) in countercurrent mode. The mixer- settlers could be replaced by another type of contactor.
Figure 3: Flowsheet showing the different steps of the process of the invention modified for conversion of battery-grade lithium carbonate into lithium hydroxide monohydrate.
Figure 4: McCabe-Thiele diagram for the first solvent extraction step (SX1). Organic phase: 0.65 M [A336][CI] and 2,6-di-tert-butylphenol (molar ratio = 1: 1) in Shellsol D70; Aqueous phase: 2.0 M NaOH; O/A = 1/5 to 5/1; 20 °C; 30 min Figure 5: McCabe-Thiele diagram for the second solvent extraction step (SX2). Organic phase: 0.59 M [A336][OR] in Shellsol D70; Aqueous phase: 1.64 M LiCI; O/A = 1/5 to 5/1; 20 °C; 30 min Figure 6: Individual operation of two-stage counter-current SX1 in mixer-settlers. Organic phase: 0.65 M [A336][CI] and 2,6-di-tert-butylphenol (molar ratio = 1: 1) in Shellsol D70; Aqueous phase: 2.0 M NaOH; O/A = 1/2; 20 °C; retention time in each mixer = 2-4 min; 1000 rpm
Figure 7: Individual operation of six-stage counter-current SX2 in mixer-settlers. Organic phase: 0.59 M [A336][OR] in Shellsol D70; Aqueous phase: 1.64 M LiCI; O/A = 3/1; 20 °C; retention time in each mixer= 4-9 min; 1000 rpm Figure 8: Simultaneous operation of SX1 (two stages) and SX2 (six stages) in counter-current mode using fresh organic phase. (SX1) Organic phase: 0.65 M [A336][CI] and 2,6-di-tert-butylphenol (molar ratio = 1: 1) in Shellsol D70; Aqueous phase: 2.0 M NaOH; O/A = 1/2; 20 °C; retention time in each mixer = 2 min; 1500 rpm; (SX2) Organic phase: 0.59 M [A336][OR] in Shellsol D70; Aqueous phase: 1.64 M LiCI; O/A = 3/1; 20 °C; retention time in each mixer = 4 min; 1500 rpm Figure 9: Simultaneous operation of SX1 (two stages) and SX2 (six stages) in counter-current mode using regenerated organic phase. (SX1) Organic phase: 0.65 M [A336][CI] and 2,6-di-tert-butylphenol (molar ratio = 1: 1) in Shellsol D70; Aqueous phase: 2.0 M NaOH; O/A = 1/2; 20 °C; retention time in each mixer= 2 min; 1500 rpm; (SX2) Organic phase: 0.59 M [A336][OR] in Shellsol D70; Aqueous phase: 1.64 M LiCI; O/A = 3/1; 20 °C; retention time in each mixer = 4 min; 1500 rpm
The invention relates to processes of converting technical-grade lithium chloride into battery-grade lithium hydroxide monohydrate. The process typically includes the following steps: (1) dissolving technical-grade lithium chloride in an organic solvent to obtain a pre-purified lithium chloride solution, since the solubility of lithium chloride in organic solvents is significantly higher than the solubility of sodium chloride or potassium chloride; (2) addition of a solution of an alkali metal hydroxide in an organic solvent (3) optionally further purifying the lithium chloride in the organic solution by non-aqueous ion exchange; (4) removing the organic solvent from the purified lithium chloride, followed by dissolution of the lithium chloride in water; (5) converting lithium chloride to lithium hydroxide by solvent extraction; (6) crystallizing battery-grade lithium hydroxide monohydrate from the aqueous lithium hydroxide solution.
In one application, the technical-grade lithium chloride used as input in the process of the invention is the final product of total evaporation of lithium-rich brines (salt lake brines or geothermal brines). The concentration of lithium chloride in these evaporated brines is typically between 10 and 60%.
In another application, the technical-grade lithium chloride used as input in the process of the invention is the sublimate of calcium chloride roasting or chlorination of spodumene. The concentration of lithium chloride in these sublimates is typically between 50 and 98%.
In yet another application, the technical-grade lithium chloride used as input in the process of the invention is the sublimate of calcium chloride roasting or chlorination of lithium ores other than spodumene, for example petalite, lepidolite, amblygonite, triphylite and zinnwaldite. The concentration of lithium chloride in these sublimates is typically between 50 and 98%.
In yet another application, the technical-grade lithium chloride used as input in the process of the invention is derived from lithium-containing metallurgical slags produced by pyrometallurgical processes for recycling of lithium-ion batteries. The concentration of lithium chloride herein is typically between 50 and 98%.
In specific embodiments, the process is stopped after step 4, without the further conversion to lithium hydroxide. By evaporation of water, solid lithium chloride is obtained.
After a further drying step at temperatures between 101 and 200 °C, anhydrous lithium chloride is formed and it can be used at a starting material for preparation of high-purity lithium metal by molten salt electrolysis. Alternatively, lithium chloride is obtained by evaporation of the organic solvent, without the need to remove water. In yet another embodiment, the process is stopped after step 5 and the lithium hydroxide solution is treated with carbon dioxide gas to precipitate battery-grade lithium carbonate. The lithium carbonate can be recovered from the solution by solid- liquid separation, followed by drying of the solid.
The present invention further envisages the conversion of battery-grade lithium carbonate to lithium hydroxide monohydrate by a process comprising the following steps: (1) dissolving lithium carbonate in an aqueous hydrochloric acid solution; (2) converting lithium chloride to lithium hydroxide in solution by solvent extraction and (3) crystallizing battery-grade lithium hydroxide monohydrate from the aqueous lithium hydroxide solution.
As alternatives to the methods described herein for chloride-hydroxide the methods can equally be sued used to convert alkali chlorides (e.g. potassium chloride, rubidium chloride, caesium chloride) and water-soluble onium chlorides into the corresponding hydroxides. The methods of the present invention provide different advantages.
First, the processes can be used to produce battery-grade lithium hydroxide monohydrate directly from technical-grade lithium chloride.
Secondly, the processes consume no chemicals other than sodium hydroxide if lithium chloride is used as starting material.
Thirdly, the processes reduce the number of process steps because the numerous precipitation, filtration and redissolution steps of conventional lithium purification processes can be avoided. This results in a considerable process intensification of lithium refining to battery-grade lithium hydroxide monohydrate.
Fourthly, the consumption of water can be largely reduced by the solvometallurgical process.
Finally, the process is widely applicable and can prepare lithium hydroxide monohydrate from different lithium sources, including from brines from salt lakes (salars), geothermal brines, hard rock lithium ores (pegmatites containing spodumene, petalite, lepidolite, amblygonite, triphylite, zinnwaldite and other lithium minerals), as well as lithium-rich slags produced by pyrometallurgical processes for recycling of lithium-ion batteries.
In the methods of the present invention, technical-grade lithium chloride is used as starting material for the process. Lithium compounds other than lithium chloride can be processed via the method of the present invention (e.g. lithium sulphate or lithium carbonate), when converted first into lithium chloride. These and other advantages will be apparent from this disclosure to those skilled in the art.
The processes of the present invention make use of the high solubility of lithium chloride in alcohol solvents, in combination with the very low solubility of potassium chloride, sodium chloride, magnesium hydroxide and calcium hydroxide in these solvents. The selective dissolution also allows to remove sulphate or carbonate impurities in the solid starting material.
The solubility properties of lithium chloride and the impurities is such that the organic solvent may can contain up to 5% or 10% water. Typically 95% ethanol is used in industrial applications of the present invention.
Remaining impurities after the solubilization of the material upon the use of solvents comprising e.g. 5% water are removed by the addition of the alkali metal hydroxide. This circumvents the need of organic solvents with water content below 5%.
The chloride-to-hydroxide anion exchange reaction for conversion of the aqueous lithium chloride solution into a lithium hydroxide solution is performed by a two-step solvent extraction reaction, whereby the organic phase consists of an onium salt and an alcohol or phenol dissolved in a diluent. In the first step, the organic phase is contacted with a concentrated aqueous sodium hydroxide solution so that the alcohol/phenol is transformed in an alkoxide/phenolate ion and simultaneously a chloride ion is transferred from the organic phase to the aqueous phase. Thus, an onium alkoxide or onium phenolate is formed in the organic phase. In the second step, the organic phase is contacted with the second aqueous phase containing lithium chloride. The alkoxide or phenolate ion gets protonated and simultaneously a hydroxide ion is formed and a chloride ion is transferred to the organic phase. As a consequence, lithium hydroxide is formed in the aqueous phase. Efficient transformation is achieved by performing the solvent extraction steps with two circuits of contactors in countercurrent mode.
The processes of the present invention are based on solvometallurgy and hydrometallurgy. "Solvometallurgy" is the extraction and refining of metals using non-aqueous solvents. Non-aqueous solvents do not imply anhydrous conditions, but rather a low water content. Arbitrarily, the water content is limited to maximum 50 vol.%. Solvometallurgy differs from hydrometallurgy by the absence of a discrete water phase. A "solvometallurgical process" is based on unit operations used in solvometallurgy. "Hydrometallurgy" is the extraction and refining of metals using aqueous solutions.
The term "non-aqueous" is used for organic solvents or organic solutions in the context of non-aqueous ion change comprising less than 50 vol.% of water.
The term "battery-grade lithium chloride" is used for a purity of at least 99.5% (w/w), at least 99.6% (w/w), at least 99.7% (w/w), at least 99.8% (w/w) or at least 99.9% (w/w) lithium chloride. The impurities that need to be controlled are sodium, potassium, magnesium, calcium, strontium, barium, iron, aluminium, silicon, boron, as well as sulphate impurities. For lithium recycled from lithium-ion batteries, additional impurities to be controlled are nickel, cobalt, copper and zinc. In lithium hydroxide monohydrate and lithium carbonate, the chloride content has to be limited to less than 20 ppm.
The term "technical-grade lithium chloride" is used for a solid salt mixture where the main component is lithium chloride and with other metal chlorides as impurities. Impurities of sulphate ions can be present at well. The purity of a technical-grade lithium chloride is less than 98% (w/w) LiCI, less than 95% (w/w) LiCI, less than 90% (w/w) LiCI, or less than 85% (w/w) LiCI. The purity of a technical-grade lithium chloride is typically more than 10% (w/w) LiCI, more than 25% (w/w) LiCI, more than 50%(w/w), or more than 75% (w/w) LiCI. Ranges of any of the above lower and upper limits are herewith discloses
Ranges of any upper value and any lower value of LiCI are herein disclosed.
A solid composition of a technical grade LiCI composition, is in contrast to a brine, solid at ambient conditions of temperature and pressure.
"Technical-grade" lithium chloride differs from "battery-grade" lithium chloride is the sense that the technical-grade compound requires extra purification steps before it can be transformed into lithium-containing precursors for battery materials (electrolytes or cathode materials).
"Organic solvents" as used in the present invention for the solubilization of battery LiCI and for the solubilization of alkali metal hydroxides refers to as well anhydrous solvents as well as solvents comprising less than 10% (v/v), less than 5% (v/v)or less than 1% (v/v) water.
The term "water-miscible" is used for organic solvents that form one phase in binary mixtures with water, over the whole concentration range. The term "water- immiscible" is used for organic solvents that form two phases in binary mixtures with water.
The term "onium" salts is used for compounds having the general formulae [Q+][A ], where Q+ is a cation and A is an anion. Q+ has the general formula (RI)(R2)(R3)(R4)X+, wherein X is N, P, As or Sb, but preferably X is N or P. Ri, R2, R3, R4 can be identical or different, and are a linear or branched alkyl chain, a cycloalkyl group or an aromatic group, wherein said alkyl chain, cycloalkyl or aromatic groups are optionally substituted with one or more substituents selected from the group consisting of halogen, hydroxyl, sulfhydryl, amino, cyano, carboxyl, amide, ester and nitro; and wherein said Ri, R2, R3, and/or R optionally comprise at least 1 heteroatom, each of said heteroatoms being independently selected from the group consisting of nitrogen, oxygen, sulphur, selenium and phosphorus.
The preceding is a simplified summary of the invention to provide an understanding of some features of the aspects, embodiments and configurations disclosed herein. This summary is neither an extensive nor exhaustive overview of the aspects, embodiments, or configurations. It is intended neither to identify key or critical elements nor to delineate the scope of the aspects, embodiments, or configurations, but to present selected concepts in a simplified form as an introduction to the more detailed description presented below. As will be appreciated, other aspects, embodiments, and configurations are possible utilizing, alone or in combination, one or more of the features set forth above or described in detail below.
The present invention relates to processes for purification of solid technical-grade lithium chloride to battery-grade lithium chloride, and to further convert the purified lithium chloride into solid battery-grade lithium hydroxide monohydrate. The process comprises the following steps: (1) dissolving technical-grade lithium chloride in an organic solvent to obtain a pre-purified lithium chloride solution, since the solubility of lithium chloride in organic solvents is significantly higher than the solubility of sodium chloride or potassium chloride; (2) adding a solution of an alkali metal hydroxide in an organic solvent; (3) optionally purifying the organic solution by non- aqueous ion exchange; (4) removing the organic solvent from the purified lithium chloride, followed by dissolution of the lithium chloride in water; (5) converting lithium chloride to lithium hydroxide by solvent extraction; (6) crystallizing battery- grade lithium hydroxide monohydrate from the aqueous lithium hydroxide solution. The invention comprises a solvent extraction process for conversion of lithium chloride into lithium hydroxide by direct chloride/hydroxide anion exchange.
Preparation of solid technical-grade lithium chloride
The starting material for the processes of the present invention is solid lithium chloride of a purity less than that required for the synthesis of battery cathode materials or electrolytes (i.e. "technical grade"). The solid technical-grade lithium chloride used as input material for the processes of the present invention can originate from different sources. These sources include (but are not limited to) hard rock lithium ores (pegmatites comprising spodumene, petalite, lepidolite, amblygonite, triphylite and other lithium minerals), salt lake brines, geothermal brines, and lithium-containing metallurgical slags from pyrometallurgical recycling processes of end-of-life lithium-ion batteries [WO2017121663]. The technical-grade solid lithium chloride obtained from these different resources contains different types and concentrations of impurities, depending on the type of material and the source. The main impurities are the alkali metal chlorides NaCI and KCI, and the alkaline earth metal chlorides MgCh and CaCh. The main anionic impurity is the sulphate ion. Lithium chloride obtained from brines can contain boron as impurity.
Lithium chloride can also be extracted from spodumene by calcium chloride roasting, as described in US patent US2,561,439. Herein, spodumene is mixed with calcium chloride and the mixture is heated at 800-1200 °C at a pressure of 20 mm Hg or less. The volatile lithium chloride is condensed from the vapor and collected. Calcium chloride can be mixed with calcium oxide or calcium carbonate prior to roasting. The mixture of spodumene and calcium compounds can be pelletized prior to heating under reduced pressure. Roasting of spodumene with chlorine gas gives a result similar to that of calcium chloride roasting. Lithium ores other than spodumene (e.g petalite, lepidolite, amblygonite, zinnwaldite, triphylite and other lithium minerals) can be treated via processes similar to those used for spodumene, with the difference that no decrepitation step is required.
Lithium chloride can be prepared from metallurgical slag from pyrometallurgical operations for recycling of lithium-ion batteries by processes similarly to those used for the recovery of lithium chloride from spodumene [Dang et al. (2020) Separation and Purification Technology 233, article number 116025]. Lithium chloride can be fumed from the molten slag of a pyrometallurgical recycling process by the addition of alkali or earth alkali chlorides, as disclosed in patent application W0202010464. Alternatively, lithium can be recovered from slags generated from recycling of lithium-ion batteries, a solvoleaching step with an HCI-saturated organic solvent to generate and solubilize lithium chloride from lithium oxide in the slags.
Solid technical-grade lithium chloride can be obtained from salt lake brines or geothermal brines by direct, full evaporation. The brine is first concentrated by solar evaporation. During the concentration, insoluble salts, other than lithium chloride, are formed and precipitate out of solution. The crystallized solids can be removed from the concentrated lithium chloride brine by centrifugation or filtration. Magnesium salts can be taken out of the concentrated brine by precipitation of magnesium hydroxide by addition of calcium hydroxide or sodium hydroxide. Calcium salts can be removed from the concentrated brine by precipitation of calcium carbonate through addition of sodium carbonate at a concentration low enough to avoid precipitation of lithium carbonate. The liquid phase can be acidified by hydrochloric acid and boron impurities can be taken out by solvent extraction. The remaining water can be removed from boron-free brine by evaporation in vacuo, to obtain the solid technical-grade lithium chloride.
Lithium salts other than lithium chloride can be used as starting material for the processes of the present invention, provided that the lithium salt is first converted into lithium chloride. Lithium sulphate produced from spodumene by decrepitation, followed by roasting with sulfuric acid can be converted to lithium chloride by a metathesis reaction with potassium chloride or sodium chloride in an organic solvent, as disclosed in US patent 3,278,260, or by reaction with sodium carbonate, followed by dissolution of the formed lithium carbonate in hydrochloric acid. Spodumene can be converted to lithium carbonate by pressure leaching with sodium carbonate (W02019220003). This lithium carbonate can be transformed into lithium chloride by dissolution in hydrochloric acid. In a similar way, lithium carbonate produced from salt lake brines or geothermal brines can be converted into lithium chloride by treatment with hydrochloric acid. Lithium phosphate (U3PO4) can be converted into lithium chloride by dissolution in hydrochloric acid, followed by treatment of the solution by iron(III) hydroxide to remove the phosphate as iron(III) phosphate (FePCU), and liquid/solid separation (US2019/0169038).
Step 1: Selective dissolution in organic solvent
In process step 1, the solid technical-grade lithium chloride is contacted with an organic solvent to form an organic solution comprising the dissolved lithium chloride. Step 1 allows for a major reduction of the concentrations of the alkali chloride impurities (sodium chloride and potassium chloride) in the lithium chloride. A solvent is used in which the solubility of lithium chloride is high, but that of the impurities is low. Additionally, the organic solvent has preferably a low toxicity, have a limited ecological footprint, be cheap, readily available and easy to recycle. Taking into account these consideration, a preferred organic solvent for use in the processes of the present invention is an alcohol. Examples of alcohols include methanol, ethanol,
1-propanol, isopropanol (2-propanol), 1-butanol, 2-butanol, tert-butanol, 2-methyl-
2-butanol, 1-pentanol, 1-hexanol, 2-ethylhexanol, 2-ethylisohexanol, cyclohexanol, 1-heptanol, 1-octanol, 1-nonanol, 1-decanol, isodecanol, furfuryl alcohol, ethylene glycol, propylene glycol, 1,4-butanediol, diethylene glycol, glycerol, PEG-200, PEG- 400. Preferably, the alcohol has a normal boiling point (b.p.) of less than 90 °C and/or form a two-phase liquid system with water. Preferred solvents are methanol (b.p. = 65 °C), ethanol (b.p. = 78 °C), isopropanol (b.p. = 82 °C) and the water-immiscible alcohols amyl alcohol, 2-ethylhexanol and 1-octanol. Water-free (anhydrous) solvents can be used, but are not required. The solvents do not have to be dried prior to use. Typically, the solvent contains between 1 and 5 vol.% water. Preferred solvents are 95% ethanol (v/v) and 96% ethanol (v/v). This ethanol can be bio ethanol. The water-immiscible solvents can have a water content up to saturation, i.e. water can be added until a two-phase mixture is formed. However, a low water content is advantageous to reduce the co-dissolution of impurities in the lithium chloride organic solution, because the solubility of sodium chloride, potassium chloride, magnesium chloride and calcium chloride in alcohol solvent rapidly increases with a higher water content of the solvent. To facilitate the recycling of the solvent in the process, the water-solvent mixture contains preferably only one type of organic solvent and not a mixture of alcohol solvents.
To facilitate the recovery of the alcohol from the alcohol-water mixture by distillation at the end of the process, alcohols are used with a boiling point lower than 90 °C. Alcohols that are immiscible with water have the advantage that purified lithium chloride can be easily stripped with water for conversion to an aqueous lithium chloride solution.
Because of the relatively high solubility of magnesium chloride and calcium chloride in alcohol solvents, it is advantageous to use technical-grade lithium chloride with concentrations of magnesium and calcium chloride that are as low as possible.
If lithium chloride derived from brines is used, the magnesium and calcium concentration can be reduced by precipitation of magnesium as magnesium hydroxide and calcium as calcium carbonate. Especially the magnesium concentration is to be controlled, because in brines the magnesium concentration is typically much higher than the calcium concentration. It is advisable to precipitate magnesium not by addition of calcium hydroxide solution, but via addition of sodium hydroxide solution, in order to avoid the introduction of calcium impurities. It is advantageous to use the waste stream of the solvent extraction process of step 5, because this contains a mixture of sodium hydroxide and sodium chloride.
The dissolution process of step 1 can be carried out in a batch reactor, of the type used in the chemical process industry (stirred-tank chemical reactor). Because of the high chloride concentrations, it is recommended to use batch reactors made of materials that are resistant to corrosion by chloride ions. Agitation can be done via a centrally mounted driveshaft with an overhead drive unit. Different types of impellor blades can be used. The mixing efficiency can be improved by the use of baffles in the reactor. However, mixing is not a critical factor in the dissolution process. For alcohols with a low boiling point, the reactor can be equipped with a reflux setup, especially when working with alcohols with a low boiling point.
The solid technical-grade lithium chloride is brought into contact with the organic solvent in a solid-to-liquid ratio that can vary between e.g. 1: 1, 1:5 or 1: 10 up to 1:50, 1:75 or 1:100. The solid-to-liquid ratio is typically kept as low as possible to reduce the consumption of organic solvent. The dissolution process is typically carried out at room temperature, or optionally at an elevated temperature. The maximum operating temperature is the boiling point of the solvent. Working at elevated temperatures can have the advantage of increasing the solubility of lithium chloride in the solvent and/or reducing the viscosity of the solution. High viscosities cannot only inhibit the filtration of the solution but may also slow down the mass transfer during the purification step via ion exchange, if this optional step is used. Working at elevated temperatures is especially an advantage in case of using the higher alcohols: amyl alcohol, 2-ethylhexanol and 1-octanol.
A countercurrent dissolution operation may be carried out as an alternative to dissolution in a batch reactor. The solvent containing substantial amounts of dissolved lithium chloride can be contacted with a fresh charge of solid technical- grade lithium chloride, while the solid depleted in lithium chloride may be advanced against fresh solvent containing only small amounts of lithium chloride. This approach will make use of the organic solvent in the most efficient way.
Step 2: Addition of a solution of an alkali metal hydroxide in an organic solvent To the solution produced in step 1, an alkali metal hydroxide solubilised in an organic solvent in added. The alkali metal hydroxide is preferably (anhydrous) lithium hydroxide or lithium hydroxide monohydrate, and the solvent is preferably the same solvent as used in step 1. The solution of (anhydrous) lithium hydroxide or lithium hydroxide monohydrate in the alcohol solvent is added to precipitate the dissolved magnesium chloride as magnesium hydroxide and dissolved calcium chloride as calcium hydroxide. At 20 °C, the solubility of (anhydrous) lithium hydroxide in ethanol is 2.36 g/lOOg and that of lithium hydroxide in ethanol is 2.18 g/lOOg. The lithium hydroxide monohydrate produced by the method of the present invention can be used.
The solid residue of the dissolution process is separated from the organic solution comprising the dissolved lithium chloride. After the lithium chloride has been extracted from the technical-grade lithium chloride, a solid-liquid separation (via filtration or centrifugation) is carried out to separate the lithium chloride solution from the solid residue. The solid residue can be washed after filtration with a small amount of fresh solvent and the washing solvent can be added to the purified lithium chloride solution.
Preferably, the precipitation of Mg(OH)2 and Ca(OH)2 is done before the insoluble alkali chloride and alkaline chloride salts (of step 1) are separated from the alcoholic lithium chloride solution by solid/liquid separation so that only 1 solid/liquid separation step is required in the flowsheet. The use of an alcoholic solution of lithium hydroxide or lithium hydroxide monohydrate is preferred over aqueous solutions of lithium hydroxide or lithium hydroxide monohydrate, because of the much lower solubility of Ca(OH)2 in alcohols compared to water. At 20 °C, the solubility of Ca(OH)2 in water is 1.73 g/L, and it is described as "insoluble" in alcohols. The hydroxide precipitation can also be done in a separate step after separation of the insoluble residue from the alcoholic lithium chloride solution, but this adds an extra process step to the flowsheet.
If a solvent with a low water content is being used, such as 95% (v/v) or 96% (v/v) ethanol (stepl) in combination with precipitation of Mg(OH)2 and Ca(OH)2 by addition of a solution of lithium hydroxide or lithium hydroxide monohydrate in the same solvent (step 2), the dissolved lithium chloride is already of battery-grade purity. In that case, the non-aqueous ion exchange process in step 3 can be omitted and the solution can be transferred directly to step 4. The non-aqueous ion exchange process in step 3 can be considered when solvents are used comprising between 5 and 10% (v/v) water.
Step 3: Organic solution purification by non-aqueous ion exchange In optional process step 3, eventual remaining co-dissolved impurity salts are removed from the lithium chloride in the organic solution by non-aqueous ion exchange. Although ion exchange reactions are typically performed in aqueous solutions, it is known to a person skilled in the art that this method is also applicable to organic solvents, both polar and non-polar solvents, as well to mixtures of organic solvents and water [Moody and Thomas (1968) The Analyst 93, 557-588]. Fleming and Monhemius describe how different metal chlorides can be extracted from a polar organic solvent by cation or anion exchangers [Fleming and Monhemius (1979) Hydrometallurgy 4, 159-167]. US patent US7,329,354 discloses a method for removal of metallic impurities from polar organic solvents by ion exchange resins. US patent US20170197204 discloses how metal impurities can be removed from viscous organic solvents by ion exchangers.
In one configuration, optional step 3 is performed with an organic cation exchanger. Strong acid cation exchangers that are based on cross-linked polystyrene (styrene- divinylbenzene co-polymer) with sulfonic acid groups are preferred. Examples of commercial strong acid cation exchange resins are DOWEX 50WX2, DOWEX 50WX4, DOWEX 50WX8, DOWEX HCR-S, DOWEX HCR-W2, DOWEX HGR, DOWEXHGR-W2, DOWEX M-31, DOWEX G26, Purolite C100H, Purolite C100X10MBH, Purolite C100X16MBH, Amberlite HPR1100, Amberlite HPR1300, Amberlite HPR650, Amberlite HPR2800 and Amberlite HPR2900. The main differences between these resins are the degree of crosslinking and the fact whether the resins are gels or macroporous (i.e. macroreticular). The degree of crosslinking has an effect on the selectivity of the resin. Gel-based resins have a higher capacity, but macroporous resins are preferred in more aggressive environments. The resins can be used in the hydrogen form (H+) or in the lithium form (Li+). The resin in the H+ form can be transformed into the Li+ form by reaction with lithium hydroxide solution. For this, the lithium hydroxide solution produced in step 5 of the process could be used. The general trend of the affinity of strongly acidic ion exchangers for metal ions is: Li+ < H+ < Na+ < K+ < Mg2+ < Ca2+. In alcohol solvent, this trend is the same but even more pronounced. Taking into account that the resin has the weakest affinity for Li+ ions, the impurities can be removed from the organic solution by the ion exchange resin. The advantage of using a cation exchanger in the Li+ form is that, upon uptake of the impurity ions, Li+ ions are released into the solution. The cation exchanger has as stronger affinity for the impurities than for H+ ions, so that a resin in the H+ form will take up the impurity ions and release protons to the organic solution. This release of protons will increase the acidity of the solution.
In another configuration, optional step 3 is performed with an inorganic cation exchanger. Instead of strongly acidic cation exchangers based on cross-linked polystyrene resins with a sulfonic acid group, inorganic cation exchange materials may be used. Of the inorganic ion exchange materials, hydrous antimony pentoxide (so-called antimonic acid or (poly)antimonic acid) is of interest. It exhibits a relatively high adsorption capacity with a reasonable rate of adsorption and desorption for alkali metal ions.
For purification of technical-grade lithium chloride with a high concentration of sulphate ions (for instance lithium chloride prepared from lithium sulphate with a metathesis reaction with sodium chloride or potassium chloride in an organic solvent), an extra non-aqueous ion exchange operation for sulphate removal can be performed in this step, using a strong base anion exchanger in the chloride form. The same step can be used to remove nitrate impurities, for instance when the lithium chloride is obtained from nitrate-rich brines. Typical strong base anion exchangers are cross-linked polystyrene resins with quaternary ammonium functional groups. Examples of such commercial resins are Purolite A300E, Purolite A850, Amberlite HPR4800, Amberlite FPA22, Dowex 1X2, Dowex 1X4, Dowex 1X8, Dowex SBR-C and Dowex Marathon 11. The anion exchange operation can be applied before or after the cation exchange operation.
The removal of metal impurities from the organic solutions by ion exchangers can be performed by columns filled with the ion exchange resin particles or particles of the inorganic ion exchanger. The organic solution is fed into the column at one end of the column, flows through the column and is collected at the other end of the column. To increase the performance of step 3 of the processes of the present invention, two or more columns can be coupled in series. Although these columns can contain the same type of ion exchanger, also different types of ion exchangers can be used in the different columns. Alternatively, multi-compartment ion exchange columns can be used to work with different ion exchangers. After exhausting of the column, the column needs to be regenerated. The cation exchanger resins can be regenerated by a hydrochloric acid solution. This will transform the cation exchanger back in the H+ form, and the cation exchanger in the H+ form can be converted back into the Li+ from by reaction with a lithium hydroxide solution. The weak base anion exchange resins can be regenerated by water.
In another configuration, the ion exchange operation is done batch-wise. The organic solution is stirred in a batch reactor (stirred-tank chemical reactor). The separation of the purified liquid from the resin(s) can be performed by filtration.
It is possible to operate ion exchange columns in a continuous operation. An example is the use of a Higgins continuous contactor [Higgins (1961) Industrial and Engineering Chemistry 53, 635-637].
The approach of non-aqueous ion exchange has the advantage of general applicability, because it can be used for all alcohols that are solvents for lithium chloride, both the lower alcohols that are miscible with water (methanol, ethanol, isopropanol) and those that are immiscible with water (amyl alcohol, 2-ethylhexanol, 1-octanol). The ion exchange with the water-miscible alcohols can be performed at room temperature or at elevated temperatures up to the boiling point of the alcohol solvent. On the other hand, the ion exchange with the water-immiscible alcohol solvents is preferably run at elevated temperatures because of the higher viscosity of these alcohol solvents. Working at enhanced temperatures will reduce the viscosity and, hence, accelerate the mass transfer of the ion exchange reaction.
Step 4: Oraanic-to-aaueous conversion
In process step 4, the purified lithium chloride is typically transferred from the organic solution to an aqueous solution. In the case of water-immiscible alcohol solvents such as amyl alcohol, 2-ethylhexanol or 1-octanol, the transfer can be performed by contacting the organic lithium chloride solution with distilled or de-ionized water. In the case of water-miscible alcohol solvents, such as methanol, ethanol, isopropanol, the solvent can be removed by evaporation or distillation in such a way that the solvent can be condensed and recycled for reuse. Lithium chloride is recovered in solid form and can be re-dissolved in water to form an aqueous solution that can be used as feed for process step 5. In another configuration, the solution of lithium chloride in the water-miscible alcohol solvent is diluted with deionized water (or distilled water) and the alcohol solvent is distilled off.
In specific embodiments of the invention wherein the conversion into LiOH is not needed or postponed, alcohol solvents, such as methanol, ethanol, isopropanol, are removed by evaporation or distillation. Lithium chloride is then recovered in solid form.
Step 5: Chloride-to-hvdroxide conversion by solvent extraction In process step 5, lithium chloride in solution is converted to lithium hydroxide by a two-step solvent extraction process. The input of step 5 is the aqueous LiOH solution obtained in step 4. It is known to a person skilled in the art that direct exchange of hydroxide ions for chloride ions by solvent extraction with a basic extractant is an in efficient process [Dehmlow et al. (1985) Tetrahedron 41, 2927-2932.]. Therefore, two coupled solvent extraction circuits are used. In the first solvent extraction circuit, a solvent comprising a chloride, an alcohol or phenol derivative and a diluent is contacted with an aqueous solution of sodium hydroxide. The alcohol or phenol reacts with the hydroxide ions in the organic phase and is transformed into an alkoxide or phenolate ion. Simultaneously, the chloride ion of the onium chloride is transferred to the aqueous phase, resulting in the formation of an onium salt dissolved in the organic phase. This organic phase enters a second solvent extraction circuit, where it is contacted with the aqueous lithium chloride solution. The alkoxide or phenolate ion is protonated by reaction with water, forming a hydroxide ion in the aqueous phase and yielding the alcohol or phenol in the organic phase. Simultaneously, a chloride ion is transferred from the aqueous phase to the organic phase. Thus, the starting mixture of the onium chloride and the alcohol or phenol is regenerated. The organic phase is transferred once more to the first solvent extraction circuit, where it can be recontacted with the concentrated sodium hydroxide solution, and the cycle is repeated.
The first solvent extraction step (SX1) can be represented by the following reaction: [Q+] [Cl ]0rg + ROHorg + NaOHaq = [Q+][OR1org + NaClaq + H20 (1)
The second solvent extraction step (SX2) can be represented by the following reaction:
[Q+] [OR ]0rg + LlClaq + H20 = [Q+][CI1org + ROH0rg + LIOHaq (2)
Here [Q+][CI ] is an onium chloride, ROH is a primary alcohol, secondary alcohol, a diol or a phenol, [Q+][OR ] is an onium alkoxide or phenolate, while subscript "org" means species in the organic phase, subscript "aq" means species in the aqueous phase, and "=" are double arrows (equilibrium reactions). Although there is only a single organic phase, the aqueous phases in reactions represented by equations (1) and (2) comprise different aqueous phases.
In the onium chloride, Q+ has the general formula (RI)(R2)(R3)(R4)X+, wherein X is N, P, As or Sb, but preferably X is N or P. Ri, R2, R3, R4 can be identical or different, and are a linear or branched alkyl chain, a cycloalkyl group or an aromatic group, wherein said alkyl chain, cycloalkyl or aromatic groups are optionally substituted with one or more substituents selected from the group consisting of halogen, hydroxyl, sulfhydryl, amino, cyano, carboxyl, amide, ester and nitro; and wherein said Ri, R2, R3, and/or R optionally comprise at least 1 heteroatom, each of said heteroatoms being independently selected from the group consisting of nitrogen, oxygen, sulfur, selenium and phosphorus. The total number of carbon atoms for Ri + R2 + R3 + R is at least 24 to minimize the solubility of the onium chloride in the organic phase. Examples of onium chlorides include tetrahexylammonium chloride, tetraheptylammonium chloride, tetraoctylammonium chloride, trioctylmethylammonium chloride, tricaprylmethylammonium chloride (Aliquat 336 , a commercial mixture of Cs and C10 quaternary ammonium chlorides with trioctylmethylammonium chloride as main component), Mextral 336 At (a commercial mixture with a composition similar to Aliquat 336) and trihexyl(tetradecyl)phosphonium chloride (Cyphos IL 101). The preferred onium chloride is tricaprylmethylammonium chloride (Aliquat 336).
The primary and secondary alcohols, and the diols are hydrophobic compounds, to minimize its solubility in the aqueous phase. Examples of primary alcohols are 1- pentanol, 1-hexanol, 2-ethylhexanol, 1-heptanol, 1-octanol, 1-nonanol, 8-methyl-l- nonanol (isodecanol), 1-decanol, 1-undecanol, 1-dodecanol, 1-tridecanol, isotridecanol, and 1-tetradecanol. Examples of secondary alcohols are 2-pentanol, 3- pentanol, 2-hexanol, 2-octanol, cyclohexanol,, l-penten-3-ol, 3-methyl-3-pentanol and 2,4-dimethyl-3-pentanol. Examples of diols are 2, 2-dimethyl-l, 3-propanediol (neopentyl glycol), 2-methyl-2,4-pentanediol, 2,3-dimethyl-2,3-butanediol (pinacol), 2,5-dimethyl-2,5-hexanediol, 2-methyl-2-propyl-l, 3-propanediol, 2,2,4- trimethyl-l,3-pentanediol, 2-ethyl-l,3-hexanediol, 1,2-octanediol, 1,2-pentanediol, 1,5-pentanediol and 1,2-hexanediol. Diols are preferred over primary and secondary alcohols. The preferred diol is 2,5-dimethyl-2,5-hexanediol.
The phenol derivative is a hydrophobic compound, to minimize its solubility in the aqueous phase, for instance an alkylphenol Alkylphenols include cresol (any isomer), ethylphenol (any isomer), and especially phenols containing an alkyl chain from 3 to 15 alkyl carbon atoms, for example propylphenol (any isomer), butylphenol (any isomer), pentylphenol (any isomer), hexylphenol (any isomer), heptylphenol (any isomer) octylphenol (any isomer), nonylphenol (any isomer), decylphenol (any isomer), undecylphenol (any isomer), dodecylphenol (any isomer), tridecylphenol (any isomer), tetradecylphenol (any isomer). The alkyl group may be linear or branched and the commercial products are often complex mixtures. Also dialkylphenols or polyalkylphenols can be used. Examples include 2-methylphenol (ort/70-cresol), 3-methylphenol (meta-cresol), 4-methylphenol {para- cresol), 2- ethylphenol, 3-ethylphenol, 4-ethylphenol, 2,6-di-methylphenol (2,6-xylenol), 3,5- dimethylphenol, 2-isopropylphenol, 3-isopropylphenol, 2-tert-butylphenol, 2-tert- butyl-4-methylphenol, 4-tert-butylphenol, 2,4-di-tert-butylphenol, 3,5-di-tert- butylphenol, 2,6-di-tert-butyl-4-methylphenol, 2,6-di-tert-butylphenol, 2,4,6-tri- tert-butylphenol, 2-sec-butylphenol, 4-sec-butylphenol, isobutylphenol, amylphenol, 4-tert-amylphenol, cyclohexylphenol, 4-tert-octylphenol, 4-tert-nonylphenol, 4-tert- dodecylphenol, 4-ocylphenol, isooctylphenol, branched octylphenol, branched 4- octylphenol, octylphenyl (isomeric mixture), sec-octylphenol, 4-nonylphenol, isononylphenol, branched nonylphenol, branched 4-nonylphenol, nonylphenol (isomeric mixture), isododecylphenol, dioctylphenol and dinonylphenol. The alkyl chain can also contain a phenyl group, for example 4-sec-butyl-2-(a-methylbenzyl)- phenol (BAMBP) and 4-tert-butyl-2-(a-methylbenzyl)-phenol (t-BAMBP). Phenols having an alkyl chain containing more than 15 carbon atoms are not preferred as they tend to cause emulsification. The preferred phenols are 4-tert-butylphenol, 2,6- di-tert-butylphenol and the mixed p-nonylphenols obtained by condensation of phenol and propylene trimer. Mixtures of different phenols can be used as well. The onium chloride and the phenol derivative are preferably dissolved in a diluent. The concentration of the onium salt in the diluent can be varied between 1, 5 or 10 up to 30, 35 or 40 vol.%, while the concentration of the phenol derivative in the diluent is in the range of 1, 5 or 10 up to 30, 35 or 40 vol.%. The concentration of the onium chloride and the phenol derivative in the diluent can be varied independently, but it is preferable that the concentration of the onium chloride matches the concentration of the phenol derivative as closely as possible on a mole basis, or that the mole ratio of the onium salt is higher than that of the alcohol/phenol. A higher than 1:1 molar ratio of alcohol/phenol to onium salt is avoided because it will lead to extraction of sodium and to contamination of the lithium solution.
The diluent can be an aliphatic diluent, a mixed aliphatic/aromatic diluent or an aromatic diluent. Examples of commercial aliphatic diluents are Shellsol D70, Shellsol D80, Shell GTL solvents GS190, GS215, GS250, GS270, GS310, Exxsol D80, Exxsol D90, Exxsol D100, Isopar L, Isopar, Escaid 110, Escaid 115, Escaid 120 and Orfom SX11. Examples of commercial aromatic diluents are Shellsol A100, Shellsol A150, toluene, xylene and 4-isopropyltoluene (para-cymene). Examples of mixed aliphatic- aromatic diluent are Shellsol 2046 and Shellsol 2325. No phase modifier needs to be added because the phenol derivative can act as a phase modifier. Preferred diluents are aliphatic diluents with cycloalkanes and branched alkanes. The preferred diluent is Shellsol D70.
The solvent for the lithium chloride to lithium hydroxide conversion can be prepared by making a homogeneous mixture of the onium salt, the phenol (or alcohol) and the diluent in the desired volume or mass ratios. Alternatively, a commercial solvent can be used. An example is LIX 7820 (from BASF), which is a mixture of Aliquat 336 and 4-nonylphenol, and was developed for extraction of gold and silver from alkaline cyanide leach liquors and copper from alkaline cyanide waste solutions.
The concentration of the onium salt is between 0.05 M and 2 M, and preferably between 0.5 and 0.75 M. The concentration of the phenol (or alcohol) is between 0.05 M and 2 M, and preferably between 0.5 and 0.75 M. The concentration of the onium salt and the phenol (or alcohol) can be selected independently, although a molar ratio close to 1: 1 is preferred and the ratio should be never less than 1: 1 (i.e. the onium salt must be in molar excess with respect to the phenol (or alcohol).
The combination of an alcohol/phenol and an onium salts is needed because neither compound alone can perform the chloride/hydroxide exchange well. An onium salt on its own is not efficient because it cannot extract hydroxide ions to the organic phase. An alcohol or phenol on their own is not efficient because it will extract sodium ions from the aqueous phase to the organic phase.
The sodium hydroxide solution has a concentration between 1 M and 20 M, but preferably less than 2.5 M. It is known to a person skilled in the art that onium salts are susceptible to decomposition by strong bases. The quaternary ammonium chlorides are more stable than quaternary phosphonium chlorides and, therefore, quaternary ammonium salts are preferred over quaternary phosphonium salts. The contact time between the NaOH solution and the organic phase should be kept as short as possible, and the onium alkoxide or phenolate formed in situ in the organic phase should be reacted as quickly as possible with the lithium chloride phase. The solvent extraction process is carried out at room temperature to minimize the decomposition of the onium salt by contact with strong bases.
The solvent extraction process of step 5 is preferably carried out in industrial contactors, such as mixer-settlers, tray columns, pulse columns or centrifugal contactors. Two circuits of contactors are typically used: a first circuit is used to transform the mixture of the onium chloride and the phenol (alcohol) into the onium phenolate (alkoxide) by contact with a concentrated aqueous sodium hydroxide solution (according to equation 1), while the second circuit is used to convert the lithium chloride into lithium hydroxide (according to equation 2). The aqueous and the organic phases are sent in countercurrent mode through the circuit of contactors, because this configuration is the most efficient one for full conversion of lithium chloride into lithium hydroxide, and it makes use of the sodium hydroxide solution in the most efficient way. A flowsheet for the chloride-to-hydroxide conversion by solvent extraction with mixer-settlers in counter-current mode is shown in Figure 2. The required number of mixer-settlers in the two extraction circuits depends on the specific process parameters. The number of mixer-settlers in the circuit that converts lithium chloride into lithium hydroxide also depends on the targeted purity of the final lithium hydroxide product. In order to minimize the conversion of lithium hydroxide to lithium carbonate by reaction of the lithium hydroxide with carbon dioxide in the air, mixer-settlers are kept closed or under a blanket of inert atmosphere, or the solvent extraction is carried out in columns or centrifugal contactors. The contactors can be run in different phase continuity, but it advantageous to run the first solvent extraction circuit in aqueous-continuous mode and the second circuit in organic- continuous mode.
The depleted sodium hydroxide solution that is contaminated by the sodium chloride formed by the anion exchange reaction may be used for precipitation of magnesium hydroxide from concentrated lithium brines used for the preparation of the technical- grade lithium chloride.
Part of the lithium hydroxide solution produced in step 5 may be used in step 3 for conversion of the cation exchangers from the hydrogen form into the lithium form. The lithium is not lost since it can be re-introduced in the flow sheet when the organic lithium chloride solution is purified by non-aqueous ion exchange.
In one configuration, an additional ion exchange step is used for final purification of the lithium hydroxide solution. The chloride content in the lithium hydroxide solution can also be reduced to any targeted low value; the required number of contactors will increase accordingly. For this reason, one can opt to limit the number of contactors by having an extra purification step with a strongly base anion exchanger in the hydroxide form to remove the last traces of chloride ions. Typical strong base anion exchangers are cross-linked polystyrene resins with quaternary ammonium functional groups. Examples of such commercial resins are Purolite A300E, Purolite A850, Amberlite HPR4800, Amberlite FPA22, Dowex 1X2, Dowex 1X4, Dowex 1X8, Dowex SBR-C and Dowex Marathon 11. The purification step with an anion exchanger cannot replace the solvent extraction step for chloride-to-hydroxide conversion because this would be inefficient due to the large anion exchange capacity needed. In case the sodium content after the solvent extraction step is too high (due to drag over of sodium hydroxide from one aqueous phase to the other via entrainment in the organic phase), these traces can be removed by a strong acid cation exchanger in the lithium form. Examples of commercial strong acid cation exchange resins are DOWEX 50WX2, DOWEX 50WX4, DOWEX 50WX8, DOWEX HCR-S, DOWEX HCR-W2, DOWEX HGR, DOWEXHGR-W2, DOWEX M-31, DOWEX G26, Purolite C100H, Purolite C100X10MBH, Purolite C100X16MBH, Amberlite HPR1100, Amberlite HPR1300, Amberlite HPR650, Amberlite HPR2800 and Amberlite HPR2900.
It is known to a person skilled in the art that organic impurities can be removed from the aqueous phase after solvent extraction by an activated carbon adsorbent. Thus, by flowing the aqueous lithium hydroxide solution over a bed of activated carbon particles, entrained organic phase can be removed before the aqueous solution is sent to the crystallization step.
Step 6: Crystallization
In process step 6, lithium hydroxide monohydrate is crystallized from the aqueous lithium hydroxide solution. The solution is concentrated by evaporation and water can be further removed by evaporation in vacuo. The solid lithium hydroxide monohydrate crystallizes out, is separated from the mother liquor by solid-liquid separation (filtration, centrifugation or spray drying) and dried. Because lithium hydroxide monohydrate has a strong tendency to absorb carbon dioxide from the atmosphere, precautions are taken during crystallization, handling and storing to exclude carbon dioxide.
In another embodiment, lithium hydroxide monohydrate is recovered from the aqueous solution by antisolvent precipitation (drowning-out crystallization) by addition of ethanol to the aqueous solution. It is known that lithium chloride is well soluble in water-ethanol mixtures, whereas lithium chloride monohydrate is not [Taboada et al. (2007) Chemical Engineering Research and Design 85, 1325-1330.]. This approach allows to obtain battery-grade lithium hydroxide monohydrate from the outlet solution of the lithium chloride to lithium hydroxide solvent extraction process, even if the conversion of chloride to hydroxide is less than 99%. After precipitation or crystallization, the lithium hydroxide monohydrate can be recovered by solid-liquid separation (filtration, centrifugation or spray drying) and the ethanol can be recovered from the water-ethanol solution by distillation. The distilled ethanol is ready to be reused in the process. The aqueous solution after ethanol removal contains mainly lithium chloride and some lithium hydroxide and this solution can be returned to step 4 of the process so that no lithium gets lost.
In yet another embodiment, lithium hydroxide monohydrate is recovered from the aqueous solution by antisolvent precipitation (drowning-out crystallization) by addition of isopropanol (isopropyl alcohol) to the aqueous solution. The advantage of using isopropanol instead of ethanol is that the solubility of LiOH in isopropanol (0 g/100 mL or O M at 20 °C) is much lower than that in ethanol (2.36 g/100 mL or 1 M at 20 °C). Therefore, isopropanol is the preferred solvent for antisolvent precipitation. When ethanol is being used as antisolvent, it is recommended to partially evaporate water to increase the LiOH concentration before adding ethanol.
High-puritv lithium chloride
The purified lithium chloride solution formed in step 4 of the invention can be concentrated by evaporation of water and can be totally dewatered by evaporation in vacuo. Drying of the hydrated lithium chloride at temperatures between 101 and 200 °C yields high-purity anhydrous lithium chloride that is suitable as a start product for preparation of lithium metal by molten salt electrolysis. Conversion of battery-grade lithium carbonate into lithium hydroxide monohvdrate The present invention can be used to prepare battery-grade lithium hydroxide monohydrate from battery-grade lithium carbonate by a simple modification of the process flow sheet: instead of the purified lithium chloride solution prepared in step 5, step 6 is performed by a lithium chloride solution prepared by dissolution of purified, battery-grade lithium carbonate in hydrochloric acid. The flowsheet is shown in Figure 3. This embodiment enables to transform purified lithium carbonate that is produced by the carbonation process from brines into a product with a higher added value, i.e. into lithium hydroxide monohydrate. It avoids the formation of calcium carbonate waste that is produced in the conventional process of converting lithium carbonate into lithium hydroxide by reaction with calcium hydroxide solution or slurry.
Conversion of technical-grade lithium chloride in battery-grade lithium carbonate The present invention can be used to prepare battery-grade lithium carbonate from technical-grade lithium chloride by a simple modification of the process flow sheet. Instead of crystallization (step 6), the lithium hydroxide solution produced in step 5 is reacted with carbon dioxide gas to precipitate lithium carbonate. The solid lithium carbonate crystallizes out, is separated from the solution by solid-liquid separation (filtration, centrifugation or spray drying) and dried.
Conversion of alkali chlorides and onium chlorides in the corresponding hydroxides The solvent extraction process described in process step 5 for the conversion of lithium chloride to lithium hydroxide can be used for the synthesis of alkali metal hydroxides from the corresponding alkali metal chlorides (alkali metal = potassium, rubidium, caesium) or for the synthesis of water-soluble onium chlorides (onium chlorides with show alkyl chains). For this, the lithium chloride solution is replaced by an alkali chloride or onium chloride solution. Examples of compounds that can be prepared via this method are potassium hydroxide (KOH), rubidium hydroxide (RbOH), caesium hydroxide (CsOH), tetramethylammonium hydroxide, tetraethylammonium hydroxide, tetra-r?-propylammonium hydroxide, tetra- isopropylammonium hydroxide and tetra-r?-butylammonium hydroxide. The solid alkali hydroxides and onium hydroxides can be obtained in solid form by crystallization, as described for step 6. EXAMPLES
Example 1: Selective dissolution of chloride salts in ethanol
The dissolution of three types of mixtures of solid alkali and alkali earth chloride salts was tested in absolute ethanol, 95 vol.% ethanol and 90 vol.% ethanol. The composition of the mixtures is given in Table 1. The mixtures were shaken with 10 mL of solvent with a speed of 300 rpm for 30 minutes at room temperature , followed by filtration through a syringe filter with a pore size of 45 micrometres. The samples were prepared in triplicate. The compositions of the filtrates are shown in Table 2-4. NaCI and KCI have a limited solubility in ethanol, which enabled a good separation from LiCI. The concentrations of sodium and potassium in solution increased with the increasing water content of the ethanol solvent. Ethanol with no more than 5 vol.% water (absolute ethanol or ethanol 95 vol.) is appropriate to keep the solubility of NaCI and KCI sufficiently low. This shows that the use of absolute ethanol is not necessary in order to get a good separation of LiCI from NaCI and KCI impurities. However, MgCh and CaCh did show a significant solubility in ethanol, and are co dissolved with LiCI.
Table 1. Composition of the solid chloride salt mixtures.
Mass of chloride salts (g)
Type of mixtures (feeds) Feed 1 Feed 2 Feed 3
NaCI 0.130 0.130 0.130
KCI 0.100 0.100 0.100
LiCI 1.866 1.866 1.866
MgCI2 0.007 0.036
CaCI2 0.008 0.039
Table 2. Dissolution of chloride salts in absolute ethanol (g L 1).
Mg Ca K Na Li
Feed 1 - - 0.023±0.007 0.022±0.013 29.4±1.4
Feed 2 0.115±0.004 0.238±0.007 0.024±0.006 0.024±0.013 27.3±0.5
Feed 3 0.511±0.010 0.979±0.077 0.029±0.005 0.024±0.014 25.7±1.1 Table 3. Dissolution of chloride salts in 95 vol.% ethanol (g L 1).
Mg Ca K Na Li
Feed 1 - - 0.071±0.022 0.024±0.011 27.9±1.1
Feed 2 0.135±0.010 0.231±0.009 0.058±0.017 0.024±0.013 26.4±1.0
Feed 3 0.683±0.026 1.028±0.032 0.096±0.023 0.031±0.01 26.7±0.1
Table 4. Dissolution of chloride salts in 90 vol.% ethanol (g L 1).
Mg Ca K Na U
Feed 1 : 0.167±0.042 0.045±0.021 27.1±0.5
Feed 2 0.137±0.006 0.221±0.010 0.130±0.026 0.042±0.022 27.0±1.1
Feed 3 0.672±0.006 1.053±0.033 0.175±0.042 0.044±0.013 27.3±1.3 Example 2: Dissolution of low-grade LiCI
The dissolution of mixtures of NaCI, KCI and LiCI was tested under similar conditions as described in Example 1 (10 mL sample), but with a smaller amount of LiCI (0.9333 g). 95 vol.% ethanol and (ultrapure) water were tested as solvents. Only trace amounts of sodium and potassium were found in the ethanolic solution after equilibration: 0.021±0.005 g L 1 K and 0.009±0.008 g L 1 Na. All the LiCI present in the solid mixtures did dissolve in the 95 vol.% ethanol. All the NaCI, KCI and LiCI dissolved well in water.
Example 3: Dissolution of LiCI from mixtures mimicking brine compositions Similar experimental conditions as those described in Example 1 were applied for the dissolution of a mixture of chloride salts with large amounts of MgCh and CaCh (brine
1) and with more moderate amounts (brine 2) in lithium-containing brines (Table ). 95 vol.% ethanol or water were used as solvents. The concentrations of sodium and potassium in the 95 vol.% ethanol solutions were very low (Table ), notwithstanding the fact that the solid feed mixtures did contain large amounts of NaCI and KCI compared to LiCI w (Table ). All salt mixtures did completely dissolve well in water. Table 5: Composition of solid salt mixtures mimicking brine compositions.
Mass of chloride salts (g)
Type of mixtures (feeds) Brine 1 _ Brine 2
NaCI 2.615 1.777
KCI 0.477 0.057
LiCI 0.095 0.095
MgCI2 1.209 0.117
CaCI2 1.080 0.083
Table 6. Concentrations in 95 vol.% ethanol (g L 1) after equilibration with solid salt solutions mimicking brine compositions. _
Mg Ca K Na Li
Brine 1 16.5±0.7 29.1±2.9 0.032±0.001 0.006±0.001 1.2±0.1
Brine 2 2.43±0.14 2.19±0.15 0.053±0.005 0.105±0.014 1.2±0.1
Example 4: One-step dissolution/precipitation
Solid chloride salts mixtures as specified in Table 1 were dissolved in 0.105 mol L 1 solution of LiOH-hhO in 95 vol. % ethanol or in (ultrapure) water. The experimental conditions were the same as described in Example 1. In this experiment, a simultaneous dissolution of LiCI, MgCh and CaCh and precipitation of magnesium and calcium as Mg(OH)2 and Ca(OH)2 by LiOH was performed in the 95 vol.% ethanol solution (Table ). Lithium chloride solutions with >99.5% purity (trace metal basis) were obtained. The precipitation of calcium by addition of the LiOH-hhO solution in water was not efficient.
Table 7. Concentrations of metal chlorides in 95 vol. % ethanol after addition of UOH-H2O solutions.
Figure imgf000035_0001
Figure imgf000036_0001
*ND - not detected
Example 5: One-step dissolution/precipitation of low-grade LiCI
Amounts of chloride salts as shown in Table were dissolved in 10 mL of a 0.105 mol L 1 solution of LiOH-hhO in 95 vol. % ethanol. The experimental conditions were the same as described in Example 1. A simultaneous dissolution of LiCI, MgCh and CaCh and precipitation of Mg and Ca as Mg(OH)2 and Ca(OH)2 by LiOH-hhO was efficiently performed from the low-grade LiCI ethanolic feed (Table ). Only trace amounts of NaCI and KCI salts dissolved along with LiCI.
Table 8. Concentration in 0.105 mol L 1 solution of UOH-H2O in 95 vol. % ethanol after addition of solid chloride mixture (10 mL sample). _
Mg Ca K Na Li
Mass of added chloride 0.007 0.008 0.130 0.100 0.933 salts (g)
Concentrations of ND ND 0.027±0.002 0.021 14.4±0.7 metals in filtrate
(9 L-1) _
*ND - not detected Example 6: One-step dissolution/precipitation by NaOH and LiOH from mixtures similar to brines
The amounts of chloride salts as shown in Example 3 (Table 5, brine 2), were dissolved in 10 mL of a 0.282 mol/L solution of NaOH in 10 mL of 95 vol.% ethanol or in water. After shaking the mixture for 30 minutes at 300 rpm and room temperature, 3 mL of a 0.210 mol/L LiOH-H20 in 95 vol. % ethanol was added to the ethanolic mixture, and 0.210 mol/L of aqueous LiOH-H20 was added to the aqueous mixture. The mixtures were shaken for additional 5 minutes, and filtered through a syringe filter with 0.45 micrometers pore size. Samples without the addition of UOH-H20 were prepared in the same way. From the ethanolic feed magnesium and the vast majority of calcium was precipitated by the NaOH in the feed (Table 9). In the reaction of NaOH with the chloride salts of magnesium and calcium, NaOH was converted to NaCI, which does not dissolve well in ethanol. Less than a stoichiometric amount of NaOH necessary for the precipitation of magnesium and calcium was added in order to avoid contamination of the feed with sodium ions from possibly unreacted NaOH. The calcium ions that remained in the ethanolic feed were precipitated with U0H-H20. The feed was purified from magnesium and calcium, and only traces of sodium and potassium ions were present along with lithium ions in the ethanolic feed. The conclusion is that the ethanolic solution of NaOH can be used to simultaneously precipitate large amounts of magnesium and calcium, while dissolving LiCI. The precipitation of calcium was inefficient from the aqueous feed, in both cases. Moreover, large amounts of sodium and potassium salts were dissolved in the aqueous feed.
Table 9. Concentration in NaOH or LiOH-HzO solution in 95 vol. % ethanol after addition of solid chloride mixture ( 10 ml. sample).
Lixiviants Mg Ca K Na Li
NaOH in 95 vol.% ethanol ND 1.21 0.060 0.154 1.45
NaOH and LiOH in 95 vol.% ethanol
ND ND 0.065 0.160 1.27
NaOH in water
ND 2.42 3.11 72.9 1.47
NaOH and LiOH in water
ND 1.55 2.47 56.3 1.49
Example 7: One-step dissolution/precipitation in a 1 L reactor
The combined dissolution/precipitation process described in Example 5 was performed in a 1 L glass reactor with overhead stirrer (HITEC Zang). The composition of the solid salt mixture was: 3.5 g MgCh, 3.5 g CaCh, 10 g KCI, 13 g NaCI and 186 g LiCI. The salts were dissolved and precipitated from 1 L of 0.128 mol L 1 UOH-H20 in 95 vol. % ethanol. The mixture was stirred at a speed of 400 rpm for 30 minutes at room temperature. Initially, the temperature of the reaction mixture rapidly increased spontaneously from room temperature to 60 °C, after which it continuously decreased. After 30 minutes, the mixture was filtered through a glass filter which was covered with Manchery-Nagel MN 615 cellulose filter paper, with 4-12 micrometers retention capacity. The concentration of potassium was 0.10±0.01 g L 1 and that of lithium was 27.2±1.3 g L 1. No impurities of magnesium, calcium or sodium could be detected. A LiCI solution with >99.5% purity (trace metal basis) was obtained. Example 8: Step-by-step dissolution and precipitation from concentrated magnesium feed
The precipitation of Mg and Ca from absolute ethanol, 95 vol.% ethanol, 90 vol.% ethanol and (ultrapure) water was tested. The mixtures of chloride salts (Table ) were dissolved in 50 mL of the solvents for 30 minutes, at room temperature and 300 rpm. The mixtures were filtered and to 5 mL of each filtrate, 0.250 mL of an aqueous LiOH-hhO solution was added. The concentrations of the added LiOH-hhO solutions are shown in Tables 11-13. The solutions were then shaken at a speed of 300 rpm for 15 minutes at room temperature, and filtered through a syringe filter with 0.45 micrometers pore size. The precipitation percentages of magnesium and calcium are shown in Tables 11-13. From the ethanolic solutions, both magnesium and calcium could be precipitated upon the addition of a 2-4 mol L 1 LiOH-l-hO solution, depending on the water content. This resulted in a LiCI solution with a high trace metal purity (>99.5%). Magnesium was also precipitated from the aqueous feed upon addition of an at least 2 mol L 1 LiOH-l-hO solution. Calcium was not efficiently precipitated in the aqueous solution (Table 14).
Table 10. Initial composition of the feed solutions.
Mg Ca Na K Li
Mass of added chloride salts (g) 0.201 0.039 0.650 0.500 9.33
Feed absolute ethanol (g L 1) 0.846 0.192 ND 0.007 23.1 Feed 95 vol.% ethanol (g L 1) 0.978 0.221 ND 0.048 27.3 Feed 90 vol.% ethanol (g L 1) 0.880 0.197 ND 0.116 25.1
Aqueous feed (g L 1) 0.772 0.178 5.52 5.80 24.5
*ND - not detected Table 11. Precipitation percentages of magnesium and calcium from absolute ethanol. _
Concentration of Li0H-H20 (mol Mg Ca Li purity (%),
L 1) added to the feed in precipitation precipitation trace metal absolute ethanol (%) (%) basis
0.5 31.4 0 96.6
1.0 74.2 3.4 98.0
2.0 100.0 99.2 >99.9
2.5 100.0 93.3 >99.9
3.0 100.0 92.2 >99.9
4.0 100.0 99.7 >99.9
Table 12. Precipitation percentages of magnesium and calcium from 95 vol.% ethanol. _
Concentration of LiOH-hhO Mg Ca Li purity (%),
(mol/L) added to the feed in 95 precipitation precipitation trace metal vol.% ethanol (%) (%) basis
0.5 47.0 26.9 96.9
1.0 74.7 30.7 98.0
2.0 100.0 29.1 99.2
2.5 100.0 57.9 99.5
3.0 100.0 100.0 99.9
4.0 100.0 100.0 99.8
Table 13. Precipitation percentages of magnesium and calcium from 90 vol.% ethanol. _
Concentration of Li0H-H20 (mol Mg Ca Li purity (%),
L 1) added to the feed in 90 precipitation precipitation trace metal vol.% ethanol % % basis
0.5 27.8 3.8 96.4
1.0 68.6 22.5 97.8
2.0 100.0 16.2 98.9
2.5 100.0 77.8 99.4
3.0 100.0 100.0 99.5
4.0 100.0 100.0 99.6
Table 14. Precipitation percentages of magnesium and calcium from water.
Concentration of UOH-hhO Mg Ca 1 Li purity (%),
(mol L ) added to the precipitation precipitation trace metal basis aqueous feed % %
0.5 0.5 0 66.8
1.0 49.8 4.0 67.6
2.0 100.0 0 67.1
2.5 100.0 0 69.8
3.0 100.0 10.8 69.0
4.0 100.0 18.5 69.9 Example 9: Step-by-step dissolution and precipitation from concentrated calcium feed
The experimental conditions were the same as in Example 8. The difference is that calcium was more concentrated than magnesium in the feed solutions (Table 15). The conclusion is that upon the addition of an at least 2 mol L 1 UOH- hO solution to the ethanolic feed solutions, Magnesium and calcium were also precipitated from the solutions in which calcium is more concentrated than magnesium (Table 16-17). The purity of the LiCI solution after precipitation was very high (>99.5%). Precipitation of magnesium and calcium from aqueous solutions was inefficient (Table 19. Precipitation percentages of magnesium and calcium from water.
)
Table 15. Initial composition of the feed solutions.
Mass of added Mg Ca Na K Li chloride salts (g)
Feed absolute ethanol 0.117 1.056 ND 0.028 25.0 (g L-1)
Feed 95 vol.% 0.163 1.199 0.018 0.089 29.1 ethanol (g L 1)
Feed 90 vol.% 0.166 1.188 0.008 0.154 27.4 ethanol (g L 1)
Aqueous feed (g L 1) 0.170 1.187 5.817 6.745 30.0
*ND - not detected
Table 16. Precipitation percentages of magnesium and calcium from absolute ethanol. _
Concentration of LiOH-l-hO (mol Mg Ca Li purity (%),
L 1) added to the feed in precipitation precipitation trace metal absolute ethanol % % basis
0.5 100.0 26.5 97.2
1.0 97.9 94.5 99.7
2.0 95.1 94.4 99.6
2.5 100.0 98.6 99.8
3.0 100.0 98.7 99.8 4.0 99.5 99.4 99.8
Table 17. Precipitation percentages of magnesium and calcium from 95 vol.% ethanol. _
Concentration of Li0H-H20 (mol Mg Ca Li purity (%),
L 1) added to the feed in 95 precipitation precipitation trace metal vol.% ethanol % % basis
0.5 100.0 25.6 96.6
1.0 100.0 93.2 99.4
2.0 100.0 99.1 99.6
2.5 100.0 99.1 99.6
3.0 100.0 98.6 99.6
4.0 100.0 99.2 99.6
Table 18. Precipitation percentages of magnesium and calcium from 90 vol.% ethanol. _
Concentration of LiOH-hhO (mol Mg Ca Li purity (%),
L 1) added to the feed in 90 precipitation precipitation trace metal vol.% ethanol % % basis
0.5 100.0 23.8 96.0
1.0 94.1 89.0 99.0
2.0 100.1 99.4 99.4
2.5 100.0 99.3 99.4
3.0 100.0 99.3 99.4
4.0 100.0 98.5 99.4 Table 19. Precipitation percentages of magnesium and calcium from water.
Concentration of U0H-H20 Mg Ca Lj .. ,0/ .
(mol L 1) added to the aqueous precipitation precipitation pun . . c . n// trace metal basis
0.5 96.4 0 68.2
1.0 92.6 0 69.2
2.0 62.8 0 68.7
2.5 67.3 0 67.2
3.0 63.8 0 68.0
4.0 92.2 22.2 69.2
Example 10: Precipitation of equimolar amounts of magnesium and calcium The experimental conditions were similar as in Example 8. The difference is that magnesium and calcium salts were added in approximately equimolar amounts. The precipitation was performed by the addition of a 4 mol L 1 aqueous solution LiOH-hhO solution to the 95 vol.% ethanolic or aqueous feed. The conclusion is that up to 37 mmol L 1 of magnesium and calcium each were precipitated from the 95 vol.% ethanolic feed by the LiOH-hhO solution (feed 4 table 20). High-purity LiCI solutions were obtained (>99.5%). The precipitation of calcium was not efficient from the aqueous feed solution (Table 21. Precipitation percentages of magnesium and calcium from water.
). Moreover, large amounts of NaCI and KCI dissolved in the aqueous solution. This resulted in an impure LiCI aqueous solution.
Table 20. Precipitation percentages of magnesium and calcium from 95 vol.% ethanol.
Figure imgf000044_0001
1 0.148 0.240 ND 0.064 27.5 98.4 100.0 100.0 99.8
2 0.492 0.787 ND 0.066 27.8 95.4 100.0 100.0 99.7
3 0.718 1.221 ND 0.053 25.7 92.8 99.8 100.0 99.8
4 0.894 1.470 ND 0.061 25.3 91.2 100.0 100.0 99.8
88.6 8.2
5 1.229 1.756 ND 0.045 25.2 89.3 94.2
47.9 0
6 2.570 2.091 ND 0.043 24.0 83.6 88.2
*ND - not detected Table 21. Precipitation percentages of magnesium and calcium from water.
Figure imgf000044_0002
1 0.157 0.231 5.37 5.12 28.3 72.3 61.5 4.0 73.4
2 0.463 0.753 5.34 5.17 27.2 69.9 84.1 14.7 71.7
3 0.841 1.377 5.97 5.70 27.5 66.5 99.9 26.9 70.4
4 0.990 1.570 5.69 5.45 26.4 65.8 100.0 21.0 69.6
5 1.322 2.010 4.38 4.67 29.1 70.1 99.1 9.7 74.1
6 2.880 2.503 4.28 4.42 28.5 66.9 76.7 13.0 71.8 Example 11: Precipitation of magnesium and calcium from ethanol-water mixtures
Precipitation of magnesium and calcium by a 4 mol L 1 aqueous LiOH-hhO solution from ethanol-water mixtures was tested. MgCb, CaCh and LiCI were dissolved in solvents consisted of 95 vol.% to 25 vol.% of ethanol. Precipitation of dilute solutions (feeds 1-5, Table 22) and concentrated solutions (feeds 6-10, Table 22) of magnesium and calcium was investigated. 0.250 mL of a 4 mol L 1 aqueous LiOH-l-hO solution was added to 5 mL of feed. The mixture was shaken at a speed of 300 rpm for 15 minutes at room temperature, and then filtered. The results in Table 21 show that precipitation of magnesium and calcium is feasible from solutions up to 50 vol.% of water. A LiCI solution with >99.5% metal purity was obtained. With increasing water content, the precipitation of calcium was less efficient.
Table 22. Precipitation percentages of magnesium and calcium by 4 mol L 1 aqueous ϋOH·H2q solution from ethanol-water mixtures.
Figure imgf000045_0001
1. 95 vol.% ethanol 0.260 0.168 27.8 98.5 100.0 100.0 100
2. 90 vol.% ethanol 0.252 0.175 28.6 98.5 96.1 99.2 99.9
3. 80 vol.% ethanol 0.245 0.156 26.6 98.5 98.9 99.0 99.9
4. 50 vol.% ethanol 0.233 0.174 26.9 98.5 96.2 100.0 99.9
5. 25 vol.% EtOH 0.236 0.184 28.4 98.5 59.2 100.0 99.7
6. 95 vol% ethanol 1.034 0.688 25.6 93.7 99.4 100.0 99.9
7. 90 vol.% ethanol 1.140 0.703 25.7 93.3 99.9 100.0 99.9
8. 80 vol.% ethanol 1.177 0.742 28.0 93.6 99.4 100.0 99.9
9. 50 vol.% ethanol 1.174 0.745 27.9 93.6 96.6 100.0 99.9
10. 25 vol.% ethanol 1.197 0.818 31.1 93.9 75.5 100.0 99.1 Example 12: Dissolution and precipitation of chloride salts in 2-ethyl hexanol
The mixtures of salts composed of 1.888 g of LiCI, 0.13 g of NaCI, 0.1 g of KCI, 0.036 g of MgCh and 0.039 g of CaCh were dissolved in 10 mL or in 50 mL of 2- ethylhexanol. The experimental conditions were similar as in Example 1. The difference is that after the shaking the samples were firstly centrifuged for 30 minutes at 4000 rpm, and then filtered through a syringe filter with 0.45 micrometres pore size. Additionally, the same mixtures of salts were prepared together in 10 mL of 2- ethylhexanol but with dissolution time of 24 hours. The concentrations of metals in the filtrates are shown in the Table 23 as g kg-1 and not as g L_1. The mass of samples, instead of the volume, was considered for the calculation of concentrations due to the better accuracy of mass measurements of relatively viscous samples. In 10 mL of 2-ethylhexanol up to 31% of LiCI dissolved, regardless of the dissolution time. In 50 mL of 2-ethylhexanol the dissolution of LiCI was somewhat improved (up to 43%). Chloride salts of sodium and potassium were almost insoluble. Chloride salts of magnesium and calcium dissolved quite well in 2-ethylhexanol, along with LiCI. The conclusion to be drawn is that under the same experimental conditions, 2- ethylhexanol is less suitable for the dissolution of impure concentrated LiCI than ethanol. 2-Ethylhexanol can be suitable for the separation of low-grade lithium chloride from sodium and potassium, due to the solubility differences of their chloride salts.
Table 23. Concentrations in 2-ethyl hexanol feeds (g kg 1).
Type of sample Mg Ca K Na Li
10 mL of 2- ethyl hexanol (30 0.273±0.063 0.617±0.015 <0.025 <0.025 8.58±0.39 minutes)
10 mL of 2- ethyl hexanol (24 Q.659 I.42 <0.025 <0.025 9.46 hours)
50 mL of 2- ethyl hexanol (30 0.199±0.005 0.179±0.021 <0.025 <0.025 2.62±0.21 minutes) Example 13: Removal of sodium and potassium by non-aqueous ion exchange
A mixture of chloride salts as in the feed 1 of Example 1 was prepared in 90 vol.% ethanol, and under the same conditions as in Example 1. The recovery of sodium and potassium from these solutions was tested batchwise with Amberlyst 15 dry ion exchange resin in H+ or Li+ form, and Chelex 100 resin in H+ form. A liquid-to-solid ratio of 20 was used. The reaction time was 18 hours at room temperature and shaking was performed at 300 rpm. Both resins could more efficiently recover potassium ions than sodium ions from the feed (Table 24). More efficient uptake of potassium was achieved with Amberlyst 15 resin, due to its higher ion exchange capacity.
Table 24. Removal of sodium and potassium by non-aqueous ion exchange.
_ . . , Concentration (g L 1) Lithium purity (%),
Feed solutions and resins Na K Li trace metal basis
Initial cone, in the feed with 90 n ri1 o 0.153 28.2 99.4 vol.% ethanol
Amberlyst 15 H+ form 0.014 0.065 25.2 99.7
Chelex 100 H+ form 0.018 0.104 27.1 99.6
Amberlyst 15 Li+ form 0.014 0.068 28.0 99.7
Example 14: Non-aqueous ion exchange of sodium and potassium from LiCI/LiOH feeds
The recovery of sodium and potassium by Amberlyst 15 in H+ form was investigated from a feed containing LiOH and LiCI. Firstly, NaCI, KCI, LiCI, MgCh and CaCh were dissolved in 50 mL of 50 vol.% ethanol to obtain the concentrations of the feed as shown in the Table 25. To 10 mL of this solution, 0.5 mL of an aqueous 4 mol L 1 LiOH- hO solution was added, shaken for 15 min at 300 rpm, and filtered. The filtrate still contained calcium ions (Table 25). When NaCI and KCI salts were present in the feed, the precipitation of calcium was less efficient from the 50 vol. % ethanol solution than in the case when these salts were not present, as in Example 11. Then, to 2.5 mL of the filtrate, 0.125 mL of a 4 mol L 1 LiOH-FhO solution was added and the procedure was repeated as in the first precipitation step. The second filtrate (1 mL) was added to 0.1 g of Amberlyst 15 resin and shaken at a rate of 300 rpm for 15 minutes at room temperature. The composition of each solution is shown in Table 25. The resin could recover about 43% of the sodium ions and 54% of the potassium ions from the feed after precipitation of magnesium and calcium with the UOH-H20 solution.
Table 25. Non-aqueous ion exchange of LiCI/LiOH feeds.
Concentration (g L 1) Mg Ca Na K Li
Before precipitation 0.405 0.712 0.225 0.472 33.4
After 1st precipitation ND 0.104 0.225 0.464 35.1
After 2nd precipitation ND ND 0.213 0.447 37.8
After NAIX ND ND 0.122 0.207 31.0
Example 15: pH dependence of sodium and potassium recovery by Amberlyst 15 First, NaCI, KCI and LiCI were dissolved in 95 vol.% ethanol, like in Example 1, feed
1. In order to assure measurable concentrations of sodium ions in the feed, the sample was diluted with water to achieve 80 vol.% of ethanol and an additional amount of NaCI was dissolved in the feed (Table 26). In 1.5 mL of the feed 0.075 microliters of one of the following aqueous solutions was added: a) 0.05 mol L_1 HCI, b) 0.05 mol L_1 LiOH-hhO, and c) 0.20 mol L_1 LiOH-hhO. Then, 1 mL of each feed was added to 0.1 g of Amberlyst 15 resin. The feed without any addition of HCI or UOH-H20 was also tested. The mixtures were shaken for 15 minutes at room temperature and shaking speed of 300 rpm. The recovery of potassium from the tested feeds was up to 40 %. The conclusion to be drawn is that when the concentration of potassium in the concentrated solution of LiCI was very low (e.g. <0.04 mg L_1), its recovery by the Amberlyst 15 resin was not significantly affected by the pH of the feed. When the concentration of sodium was very low (<0.03 mg L x) its recovery by the Amberlyst 15 resin within 15 minutes of reaction time was negligible. However, the solubility of NaCI in the concentrated ethanolic feeds of LiCI was typically very low (examples 1-3). Table 26. Concentrations (g L 1) in the fepcls and samples after non-aqueous ion-exchange with Amberlyst 15 resin.
Lithium
Type of feeds and
Figure imgf000049_0001
purity (%),
Na Li samples trace metal basis
Feed 0.05 mol L-1 HCI 0.036 0.026 28.6 99.78
Sample 0.05 mol L-1
HCI 0.022 0.026 24.7 99.81
Feed 0.05 mol L-1
LiOH-H20 0.037 0.019 27.4 99.80
Sample 0.05 mol L-1
LiOH-H20 0.022 0.023 23.0 99.80
Feed 0.2 mol L-1
LiOH-H20 0.038 0.021 27.6 99.78
Sample 0.2 mol L-1
LiOH-H20 0.023 0.024 23.4 99.80
Feed 80 vol. ethanol 0.041 0.018 30.0 99.80
Sample 80 vol. ethanol 0.026 0.018 25.9 99.83 Example 16: Screening process: selecting suitable composition of organic phase for double solvent extraction operation
The performance of various alcohols/phenols and diluents in the double solvent extraction operation as described in process step 5 is evaluated. The conversion rates for both the first (SX1) and the second (SX2) extraction step are reported for a range of different alcohols/phenols as well as different diluents, Shellsol A150, Shell GTL Solvent GS190 and ShellSol D70. The conversion of tricaprylmethylammonium chloride (Aliquat 336, [A336][CI]) to its phenolate form [A336][OR] was investigated as a function of the alcohols/phenols and the diluents. Equal masses of the organic phase (0.004 mol [A336][CI] and 2,6-di-tert-butylphenol (molar ratio = 1: 1) in a diluent) and the aqueous phase containing (10 M NaOH) were contacted in 20 mL vials at 20 °C for 30 min. After extraction, the two phases were separated, and the aqueous phase was titrated with 0.05 M AgNCh to determine the chloride concentration.
The composition and the purity of tricaprylmethylammonium chloride (Aliquat 336, [A336][CI]) was determined with LC-MS. The tricaprylmethylammonium chloride had a quaternary ammonium content of 90.6%. It was composed out of 30.7% of trioctylmethylammonium choride, 39.7% of dioctyldecylmethymammonium chloride, 23.5% of octyldidecylmethylammonium chloride and 6.0% of tridecylammonium chloride. Based on the established composition of tricaprylmethylammonium chloride, an average molecular weight of 432 g/mol was calculated.
The percentage conversion for the first solvent extraction step was defined as %E- SX1 = ([CI ]aq*(Vaq/Vorg)/[CI ]ini, org)*100, where [Cl ]aq and [Cl ]ini, org are the chloride concentration in the aqueous phase and in the initial organic phase; Vaq and Vorg are the volume of the aqueous and organic phase, respectively. The experiments were performed in triplicate and the data are presented in Table 27 as average values with a standard deviation.
No data are reported for the experiments carried out with Shellsol A150 due to the formation of a precipitate and a change in volume ratio when contacting Shellsol A150 with NaOH. Also the use of nonylphenol and 4-tert-octyl phenol in combination with the diluent Shell GTL Solvent GS190 and Shellsol D70 resulted in gel formation. For these phenols, no experimental data are reported. The conclusion to be drawn from Table 24 is that phenols lead to much higher [A336][OR] percentage conversion than the alcohols and diols, respectively. The order of increasing percentage conversion is alcohols < diols < phenols. There is little difference between the percentages conversion when comparing the diluents Shell GTL Solvent GS190 and Shellsol D70.
The organic phase containing [A336][OR] in a diluent was used for the second solvent extraction conversion of LiCI to LiOH. Equal masses of the organic phase and the aqueous phase containing (0.25 M LiCI) were contacted in 20 mL vials at 20 °C for 30 min. After extraction, the two phases were separated, and the aqueous phase was titrated with 0.1 M HNO3 to determine the hydroxide concentration.
LiOH is known to form U2CO3 when it absorbs carbon dioxide from the air. Some of the LiOH formed in the second solvent extraction step is transformed into U2CO3 by contact with air dissolved in the aqueous phase. This is clearly seen by the presence of two equivalence points in the titration curve. The content of hydroxides and carbonates are determined in an aqueous acid-base titration using nitric acid of 0.1 M as titrant. The total amount of produced LiOH in the aqueous phase was calculated by considering the amount of LiOH that was converted to U2CO3 and the detected amount of LiOH. The percentage conversion for the second solvent extraction step was defined as %E- SX2 = ([LiOH]eq/[LiCI]ini)*100, where [LiOH]eq is the LiOH concentration in the aqueous phase and [LiCI]™ is the initial concentration of LiCI in the aqueous feed solution. The experiments were performed in triplicate and the data are presented in Table 28 as average values with a standard deviation. The percentage conversion of the two solvent extraction steps combines is shown in Table 29 as average values with a standard deviation.
From Table 29 it can be concluded that the order of increasing percentage conversion of the double solvent extraction operation is alcohols < diols < phenols and that the most promising candidate is 2,6-di-tert-butylphenol. The total percentage conversion is also higher when diluent Shellsol D70 is used instead of Shell GTL Solvent GS190.
Table 27. Percentage conversion of the first solvent extraction step (SX1)§
%E-SX1 %E-SX1
Alcohols/phenols
Diluent: GS190 Diluent: D70
1-pentanol 18.7 ± 0.1 19.0 ± 0. 1
2-pentanol 14.0 ± 0.1 14.0 ± 0. 1
3-pentanol 13.2 ± 0.1 13.2 ± 0. 1 l-penten-3-ol 19.4 ± 0.2 19.7 ± 0. 1 3-methyl-3-pentanol 12.9 ± 0.3 12.3 ± 0. 2
2.4-dimethyl-3-pentanol 13.5 ± 0.1 12.9 ± 0. 2
1-hexanol 20.3 ± 0.1 20.1 ± 0. 1
2-ethyl-l-hexanol 16.4 ± 0.1 16.6 ± 0. 1 1-octanol 19.8 ± 0.2 20.2 ± 0. 1
1-decanol 20 ± 0.3 20.1 ± 0. 7
2-ethyl-l,3-hexanediol 36.7 ± 0.1 36.7 ± 0. 2 2-methyl-2,4-pentanediol 44.4 ± 0.3 45.1 ± 0. 7
2.5-dimethyl-2,5-hexanediol 25.8 ± 0.3 25.6 ± 0. 2 p-cresol 78.7 ± 1.1 88.9 ± 0. 7
Extraction mixture Extraction mixture
4-tert-butylphenol solidified solidified
2.6-di-tert-butylphenol 82.0 ± 0.9 94.8 ± 0.2
2.4.6-trimethylphenol 88.8 ± 1.2 86.1 ± 0.6
2.6-dimethylphenol 88.7 ± 1.9 87.2 ± 0.2
§Organic phase: 0.004 M [A336][CI] and phenol/alcohol (molar ratio = 1: 1) in Shell GTL Solvent GS190 or Shellsol D70; Aqueous phase: 10.0 M NaOH; mass ratio = 1: 1; 20 °C; 30 min. Table 28. Percentage conversion of the second solvent extraction step (SX2)§
%E- SX2 %E- SX2
Alcohols/phenols Diluent: GS190 Diluent: D70
1-pentanol 87.7 ± 5.4 86.6 ± 0.7
2-pentanol 81.4 ± 0.1 82.8 ± 1.4
3-pentanol 79.4 ± 0.8 81.6 ± 2.5 l-penten-3-ol 78.8 ± 1.9 79.1 ± 4.4 3-methyl-3-pentanol 48.4 ± 3.3 72.6 ± 3.0
2.4-dimethyl-3-pentanol 35.2 ± 1.4 70.8 ± 8.2
1-hexanol 56.7 ± 0.5 79.2 ± 6.5
2-ethyl-l-hexanol 84.1 ± 0.1 80.7 ± 7.7 1-octanol 76.1 ± 9.2 86.7 ± 6.7
1-decanol 71.2 ± 2.5 78.9 ± 6.6
2-ethyl-l,3-hexanediol 69.7 ± 5.0 58.2 ± 3.1 2-methyl-2,4-pentanediol 55.3 ± 2.8 50.6 ± 3.8
2.5-dimethyl-2,5-hexanediol 3-phase system 41.9 ± 0.7 p-cresol 31.1 ± 0.2 32.3 ± 2.1
2.6-di-tert-butylphenol 72.9 ± 1.2 78.10 ± 2.2
2.4.6-trimethylphenol 25.3 ± 0.6 27.4 ± 0.9
2.6-dimethylphenol 18.7 ± 0.8 21.2 ± 0.9 §Organic phase: [A336][OR] and phenol/alcohol (molar ratio = 1: 1) in Shell GTL Solvent GS190 or Shellsol D70; Aqueous phase: 0.25 M LiCI; mass ratio = 1: 1; 20 °C; 30 min.
Table 29. Percentage conversion of the double solvent extraction operation
% E-Tot %E-Tot
Alcohols/phenols
Diluent: GS190 Diluent: D70
1-pentanol 16.8 ± 0.9 16.8 ± 0.2
2-pentanol 11.7 ± 0.1 11.9 ± 0.2
3-pentanol 10.7 ± 0.1 11.1 ± 0.3 l-penten-3-ol 15.7 ± 0.3 16.0 ± 0.8 3-methyl-3-pentanol 6.3 ± 0.4 8.9 ± 0.2 2,4-dimethyl-3-pentanol 4.8 ± 0.2 9.4 ± 0.9
1-hexanol 11.5 ± 0.1 16.4 ± 1.3
2-ethyl-l-hexanol 14.2 ± 0.1 13.7 ± 1.2 1-octanol 15.1 ± 1.9 17.5 ± 1.3
1-decanol 14.1 ± 0.7 15.9 ± 0.8
2-ethyl-l,3-hexanediol 11.1 ± 0.8 9.2 ± 0.5 2-methyl-2,4-pentanediol 10.4 ± 0.4 9.8 ± 0.4
2.5-dimethyl-2,5-hexanediol / 3.0 ± 2.6 p-cresol 24.4 ± 0.2 32.7 ± 1.75
2.6-di-tert-butylphenol 59.7 ± 0.4 78.0 ± 1.4
2.4.6-trimethylphenol 22.4 ± 0.8 24.3 ± 0.9
2.6-dimethylphenol 16.6 ± 0.9 18.5 ± 0.8
Example 17: Effect of the onium chlorides on the double solvent extraction operation
A double solvent extraction operation as described in Example 16 was performed using different onium chlorides and using Shellsol D70 as a diluent. The tested onium chlorides are tricaprylmethylammonium chloride (Aliquat 336) and trihexyl(tetradecyl)phosphonium chloride (Cyphos IL 101). The experiments were performed in triplicate. The conversion rates for both the first (SX1) and the second (SX2) extraction step, as well as for both the extraction steps combined are reported in Table 30 as average values with a standard deviation. Only a small difference was observed between the conversion rates obtained for tricaprylmethylammonium chloride and trihexyl(tetradecyl)phosphonium chloride. Table 30. Effect of the onium chloride on the double solvent extraction operation
Onium
Phenol %E-SX1§ %E- SX2§§ %E-Tot
Cl
A336 88.92 ± 0.7 35.90 ± 1.9 32.7 ± 1.8 p-cresol C101 89.73 ± 0.1 49.27 ± 2.3 41.6 ± 1.9
2.6-di-tert- A336 94.83 ± 0.2 80.16 ± 1.3 78.0 ± 1.4 butylphenol C101 99.32 ±0.4 72.40 ± 0.6 67.7 ± 0.8
2.4.6- A336 86.51 ± 0.6 39.22 ± 5.5 34.8 ± 4.8 trimethylphenol ciOl 90.49 ± 0.3 49.29 ± 1.4 42.0 ± 1.2
§Organic phase: 0.004 mol [A336][CI] or Cyphos IL 101 and phenol (molar ratio = 1: 1) in Shellsol D70; Aqueous phase: 10.0 M NaOH; mass ratio = 1: 1; 20 °C; 30 min §§Organic phase: [A336][CI] or Cyphos IL 101 (molar ratio = 1: 1) in Shellsol D70; Aqueous phase: 0.25 M LiCI; mass ratio = 1: 1; 20 °C; 30 min. Example 18: Effect of the diluent on the double solvent extraction operation
A double solvent extraction operation as described in Examples 16 and 17 was performed using 2,6-di-tert-butylphenol and Aliquat 336 in different diluents as the organic feed. The conversion rates for both the first (SX1) and the second (SX2) extraction step, as well as for both the extraction steps combined are reported in Table 31 as average values with a standard deviation.
The composition of the diluent seems to have an effect on the percentage conversion. Shell GTL Solvent GS190 is composed entirely of paraffins, while Shellsol D70 has 50% paraffins and 50% naphthenes. Shellsol 2325 and Shellsol 2046 distinct themselves further from Shellsol D70 by also containing 20% of aromatics besides 40% paraffins and 40% naphthenes. p-Cymene, on the other hand, is for 100% composed of aromatics.
Table 31. Effect of the diluent on the double solvent extraction operation
Diluent %E-SX1 %E- SX2 %E-Tot
Shell GTL Solvent GS190 82.0 ± 0.9 72.9 ± 1.2 59.7 ± 0.4
Shellsol D70 94.8 ± 0.2 78.10 ± 2.2 78.0 ± 1.4 p-cymene 94.0 ± 0.2 88.15 ± 1.8 85.0 ± 1.9 Shellsol 2325 94.6 ± 0.9 81.2 ± 1.8 84.7 ± 0.8 Shellsol 2046 93.5 ± 0.8 77.1 ± 3.2 82.5 ± 2.7
§Organic phase: 0.004 mol [A336][CI] and 2,6-di-tert-butylphenol (molar ratio = 1: 1) in diluents; Aqueous phase: 10.0 M NaOH; mass ratio = 1: 1; 20 °C; 30 min §§Organic phase: [A336][OR] and 2,6-di-tert-butylphenol (molar ratio = 1: 1) in diluents; Aqueous phase: 0.25 M LiCI; mass ratio = 1: 1; 20 °C; 30 min.
Example 19: Effect of the base on the first solvent extraction step (SX1)
A double solvent extraction operation as described in Examples 16, 17 and 18 was performed using 2,6-di-tert-butylphenol, p-cresol or 2,4,6-trimethylphenol and Aliquat 336 in the diluent Shellsol D70 as the organic feed. The sodium hydroxide solution (10 M) employed in the first solvent extraction step of the double solvent extraction procedure was substituted with a 15 M ammonia solution. The conversion rates for the first extraction step are reported in Table 32 as average values with a standard deviation. From Table 32 it can be concluded that the conversion rates with the ammonia solution are much lower than when using a NaOH solution.
Table 32. Effect of the base on the first solvent extraction step (SX1)§
%E-SX1 %E-SX1
Phenol 10 M NaOH 15 M NH3
2.6-di-tert-butylphenol 94.8 ± 0.2 14.5 ± 0.1 p-cresol 88.9 ± 0.7 49.7 ± 0.1
2.4.6-trimethylphenol 86.5 ± 0.6 40.2 ± 0.3
§Organic phase: 0.004 mol [A336][CI] and phenol (molar ratio = 1: 1) in Shellsol D70; Aqueous phase: 10.0 M NaOH or 15 M NH3; mass ratio = 1: 1; 20 °C; 30 min
Example 20: Effect of the sodium hydroxide concentration on the first solvent extraction step (SX1)
The conversion of tricaprylmethylammonium chloride (Aliquat 336, [A336][CI]) to its phenolate form [A336][OR] was investigated as a function of the sodium hydroxide concentration. Equal volumes of the organic phase (0.65 M [A336][CI] and 2,6-di- tert-butylphenol (molar ratio = 1:1) in Shellsol D70) and the aqueous phase (0.5-20 M NaOH) were contacted in 20 mL vials at 20 °C for 30 min. After extraction, the two phases were separated, and the aqueous phase was titrated with 0.05 M AgN03 to determine the chloride concentration.
The percentage conversion for the first solvent extraction step was defined as %E- SX1 = ([CI ]aq*(Vaq/Vorg)/[CI ]ini, org)*100, where [Cl ]aq and [Cl ]ini, org are the chloride concentration in the aqueous phase and in the initial organic phase; Vaq and Vorg are the volume of the aqueous and organic phase, respectively. The experiments were performed in triplicate and the data are presented in Table 33 as average values with a standard deviation of less than 1%.
The conclusion to be drawn is that higher conversions to [A336][OR] were achieved with increasing the NaOH concentration up to 10 M. Further increase in sodium hydroxide concentration ( > 10 M NaOH) led to a decrease in conversion of [A336][OR] due to decomposition of the organic phase. Therefore, the optimum concentration of NaOH is preferably less than 2.5 M.
Table 33. Effect of the NaOH concentration on the first solvent extraction step (SX1)§
NaOH (M) %E-SX1 0.50 46.2 0.75 55.6 1.00 61.9
1.25 68.0
1.50 69.8 1.75 73.0 2.00 75.0
2.25 77.6
2.50 78.4 5.00 92.1 10.0 96.0
12.5 89.3 15.0 37.7
17.5 17.7 20.0 8.94
§Organic phase: 0.65 M (or 30% wt.)[A336][CI] and 2,6-di-tert-butylphenol (molar ratio = 1: 1) in Shellsol D70; Aqueous phase: 0.5-20 M NaOH; O/A = 1; 20 °C; 30 min. Example 21: Effect of O/A volume ratio on the first solvent extraction step (SX1)
In the same way as in Example 20, the influence of O/A volume ratio on the conversion of [A336][OR] was investigated from 1/5 to 5/1 (Table 34). The conclusion to be drawn is that the higher volume of aqueous phase used, the higher percentage extraction %E-SX1 was achieved. The use of 2.0 M NaOH led to a higher conversion of [A336][OR] than that of 1.0 M NaOH. Furthermore, the McCabe-Thiele diagram shown in Figure 4 predicts at least two counter-current stages are needed to convert >95% of [A336][CI] to [A336][OR] at O/A = 1/2.
Table 34. Effect of the O/A volume ratio on the first solvent extraction step (SX1)§
%E-SX1
O/A volume ratio 1.0 M NaOH 2.0 M NaOH
1/5 79.8 89.9
1/4 78.5 88.2
1/3 75.6 85.1
1/2 70.8 82.3
1/1 59.2 73.4
2/1 44.7 62.4
3/1 37.3 55.0
4/1 31.6 49.7
5/1 28.2 45.3
§Organic phase: 0.65 M [A336] [Cl] and 2,6-di-tert-butylphenol (molar ratio = 1: 1) in Shellsol D70; Aqueous phase: 1.0 or 2.0 M NaOH; O/A = 1/5 to 5/1; 20 °C; 30 min. Example 22: Effect of the contact time on the first solvent extraction step (SX1)
In the same way as in Examples 20 and 21, the effect of time was studied between 1-60 min for the conversion of [A336][CI] to [A336][OR] using 2.0 M NaOH at O/A = 1/2. Table 35 shows that the first solvent extraction step is fast and efficient. Equilibrium was achieved within 4 min of shaking. In fact, a short contact time (< 2 min) between the NaOH solution and the organic phase at 20 °C is preferred to minimize the decomposition of the organic phase. Table 35. Effect of the contact time on the first solvent extraction step (SX1)§
Time (min) %E-SX1
1 61.3
2 72.4
4 85.1
6 86.7
8 87.0
10 86.7
12 86.7
14 87.0
16 87.0
18 87.0
20 87.3
30 87.6
40 87.9
50 87.9
60 88.5
§Organic phase: 0.65 M [A336] [Cl] and 2,6-di-tert-butylphenol (molar ratio = 1: 1) in Shellsol D70; Aqueous phase: 2.0 M NaOH; O/A = 1/2; 20 °C; 1-60 min.
Example 23: Effect of the lithium chloride concentration on the second solvent extraction step (SX2)
The organic phase containing 0.59 M [A336][OR] in Shellsol D70 was used for the second solvent extraction conversion of LiCI to LiOH. The influence of LiCI concentration on the conversion of LiOH was determined from 0.14 to 2.36 M. Equal volumes of the organic phase (0.59 M [A336][OR] and 2,6-di-tert-butylphenol (molar ratio = 1: 1) in Shellsol D70) and the aqueous phase (0.14-2.36 M LiCI) were contacted in 20 mL vials at 20 °C for 30 min. After extraction, the two phases were separated, and the aqueous phase was titrated with 0.1 M HNO3 to determine hydroxide concentration.
The percentage conversion for the second solvent extraction step was defined as as %E- SX2 = ([LiOH]eq/[LiCI]ini)*100, where [LiOH]eq is the LiOH concentration in the aqueous phase and [LiCI]™ is the initial concentration of LiCI in the aqueous feed solution. The conclusion to be drawn from Table 36 is that increasing LiCI concentration in the aqueous feed reduced the percentage conversion of LiOH. Table 36. Effect of the LiCI concentration on the second solvent extraction step (SX2)§
LiCI (M) %£- SX2
0.14 64.3
0.19 61.5
0.24 57.7
0.28 52.3
0.33 50.6
0.38 48.0
0.42 48.1
0.47 45.6
0.52 44.5
0.57 42.6
1.63 27.8
2.36 21.2
§Organic phase: 0.59 M [A336][OR] in Shellsol D70; Aqueous phase: 0.14-2.36 M LiCI; O/A = 1/1; 20 °C; 30 min
Example 24: Effect of the phenolate concentration [A336][OR] on the second solvent extraction step (SX2)
In the same way as in Example 23, the conversion of LiOH was investigated as a function of phenolate concentration [A336][OR] in the organic phase. The conclusion to be drawn from Table 37 is that percentage conversion of LiOH increased with increasing [A336][OR]. More than 62% of LiCI was converted to LiOH with 0.59 M [A336][OR] in a single contact. It is noted that concentrations of [A336][OR] higher than 0.59 M negatively affected the phase separation and increased the extraction of Na+ as impurity (due to mutual solubility of the aqueous phase in the organic phase), and a decreased conversion efficiency of LiOH.
Table 37. Effect of the [A336][OR] concentration on the second solvent extraction step (SX2)§
[A336][OR] %£- SX2
0.29 20.3
0.32 26.9
0.36 30.7
0.40 37.8
0.47 48.8
0.53 51.5
0.55 53.7
0.57 58.3
0.59 62.9
§Organic phase: 0.29-0.59 M [A336][OR] in Shellsol D70; Aqueous phase: 0.2 M LiCI; O/A = 1/1; 20 °C; 30 min.
Example 25: Effect of the O/A volume ratio for the second solvent extraction step (SX2)
In the same way as in Examples 23 and 24, the dependence of LiOH conversion on the variation of O/A volume ratio was studied from 1/5 to 5/1 (Table 38). The conclusion to be drawn from Table 38 is that increasing organic volume resulted in an increase in LiOH conversion. At a fixed O/A ratio, the percentage conversion of LiOH slightly decreased with increasing LiCI in the feed solution. On the other hand, an increase in [A336][OR] concentration up to 1.0 M led to a slightly higher LiOH conversion, but caused problems to the phase separation. For these reasons, an O/A = 3/1 and 0.59 M [A336][OR] were selected for upscaling SX2. The McCabe-Thiele diagram suggested that 8 counter-current stages are required to achieve >92% LiOH conversion at O/A = 3/1, as shown in Figure 5.
Table 38. Effect of the O/A volume ratio on the second solvent extraction step (SX2)§
%E- SX2
O/A volume ratio
1.64 M LiCI 2.36 M LiCI 2.36 M LiCI
1/5 6.91a 4.88a 7.04b
1/4 8.44a 5.81a 8.78b
1/3 11.0a 7.70a 11.0b
1/2 15.8a 11.3a 15.2b
1/1 27.8a 21.2a 26.2b
2/1 45.4a 36.3a 40.9b
3/1 57.4a 47.0a 50.4b
4/1 66.7a 55.6a 57.3b
5/1 73.2a 62.3a 64.0b
§Organic phase: a0.59 M and b1.00 M [A336][OR] in Shellsol D70; Aqueous phase: 1.64 and 2.36 M LiCI; O/A = 1/5 to 5/1; 20 °C; 30 min. Example 26. Effect of the contact time for the second solvent extraction step (SX2)
In the same way as in Examples 23-25, the contact time required for the LiOH conversion was studied in the range 1.0-60 min (Table 39). The conclusion to be drawn from Table 39 is that the conversion of LiOH is fast and efficient. Equilibrium state was achieved within 1 min (%E- SX2 > 54%). A short contact time is preferred, particularly during operation of mixer-setters, to minimize the conversion of LiOH to U2C03.
Table 39. Effect of the contact time for the second solvent extraction step (SX2)§
Time (min) %E- SX2 1.0 54.4 2.0 55.4 4.0 55.7 6.0 56.8 8.0 57.3 10 57.3 20 57.4 30 57.4 60 57.5
§Organic phase: 0.59 M [A336][OR] in Shellsol D70; Aqueous phase: 1.64 M LiCI; O/A = 3/1; 20 °C; 1-60 min. Example 27: Dependence on phenol/[A336][CI] molar ratio for SX1 and SX2
In the same way as in examples 20-26, the influence of phenol/[A336][CI] molar ratio was examined in the range 0.5 to 1.5. The concentration of [A336][CI] was kept constant at 0.65 M, meanwhile the concentration of 2,6-di-tert-butylphenol was varied from 0.33 M to 1.3 M (Table 40).
The conclusion to be drawn from Table 40 is that the conversion of SX1 and SX2 increased with increasing the phenol/[A336][CI] molar ratio from 0.5 to 1.0. The conversion reached a plateau 88% for SX1 and 52% for SX2 at phenol/[A336][CI] molar ratio's > 1:1. Importantly, the phenol/[A336][CI] must be monitored closely as possible on a mole basis to avoid co-extraction of sodium. For instant, the molar ratio of phenol/[A336][CI] of 1:1 yielded the lowest co-extraction of 50 mg L 1 Na+ as impurity in the final lithium hydroxide product solution. On the other hand, if molar ratios of phenol/[A336][CI] higher than 1: 1 are used, the higher the lithium solution is contaminated with Na+ (e.g. 211 mg L 1 Na+ at phenol/[A336] = 3/2). Furthermore, the use of too low phenol/[A336][CI] resulted not only in low conversion efficiency of SX1 and SX2, but also in rather high co-extraction of Na+ (e.g. 85 mg/L Na+ at phenol/[A336][CI] = 1/2) (due to an increased in mutual solubility of aqueous phase in organic phase).
Table 40. Effect of phenol/[A336][CI] molar ratio on the first and second solvent extraction step (SX1) and (SX2) and on impurity extraction9
Phenol/rA336irCll %E-SX1 %E- SX2_ Na+ (mg L 1)
0.5 51.7 39.8 84.6
0.6 59.1 43.3 69.0
0.7 67.2 47.0 55.5
0.8 72.4 49.5 53.8
0.9 78.0 51.5 54.0
1.0 81.7 52.8 50.7
1.1 85.1 52.5 71.3
1.2 86.7 52.5 115.2
1.3 87.9 52.1 151.2
1.4 88.2 51.0 199.8
1.5 88.9 50.0 210.8
§Organic phase: 0.65 M [A336] [Cl] ; molar ratio 2,6-di-tert-butylphenol/[A336][CI] = 0.5-1.5 in Shellsol D70; Aqueous phase: 2.0 M NaOH (SX1; O/A = 1/2) and 1.65 M LiCI (SX2; O/A = 3/1); 20 °C; 30 min. Example 28: Simulation of 2-stage counter-current for the first solvent extraction step (SX1)
A batch simulation of two-stage counter-current conversion of [A336][CI] to [A336][OR] was performed to choose the optimized conditions that can be applied for mixer-settlers. For the first contact, the fresh organic phase (0.65 M [A336][CI] and 2,6-di-tert-butylphenol (molar ratio = 1: 1) in Shellsol D70) was mixed with the fresh aqueous solution of 2.0 M NaOH at O/A = 1/2 and 2000 rpm for 4 min. Afterwards, the obtained raffinate was used as the aqueous feed for the second stage where it was contacted with fresh organic phase. Subsequently, the loaded organic phase obtained in the second stage was contacted with fresh aqueous phase. The process was repeated until the steady state was achieved. The final raffinates had similar chloride concentration and resembled the streams that would exist in an actual continuous counter-current extraction. The conclusion to be drawn from Table 41 is that high percentage extraction of >97% is achieved within 2-stage counter- current simulation. The finding is in agreement with the constructed McCabe-Thiele diagram (Figure 4).
Table 41. Simulation of 2-stage counter-current for the first solvent extraction step (SX1)§ _
Figure imgf000063_0001
1 10.2
2 _ 9 5 _
§Organic phase: 0.65 M [A336] [Cl] and 2,6-di-tert-butylphenol (molar ratio = 1: 1) in Shellsol D70; Aqueous phase: 2.0 M NaOH; O/A = 1/2; 20 °C; 4 min
Example 29: Simulation of 6-stage counter-current for the second solvent extraction step (SX2)
In the same way as in Example 28, the simulation of 6-stage counter-current for LiOH conversion was determined using 1.64 M and 2.36 M LiCI in the feed solution at O/A = 3/1 (Table 39). The organic phase containing 0.59 M [A336][OR] in Shellsol D70 was obtained from Example 28. The conclusion to be drawn from Table 42 is that complete conversion of >99.5% is possible from the feed solution of 1.64 M LiCI. On the other hand, increasing LiCI caused a slightly lower conversion of >82% LiOH. The findings are somehow better than prediction of McCabe-Thiele diagram, as shown in Figure 5. The high conversion of >99.5% allows to produce high battery-grade LiOH. Table 42. Simulation of 6-stage counter-current for the second solvent extraction step (SX2)§
%£- SX2 _
Stage
1.64 M UCI 2.36 M LiCI
1 13.4 3.23
28.0 4.37
44.2 19.3
61.8 34.5
81.4 55.8
99.5 82.0
§Organic phase: 0.59 M [A336][OR] in Shellsol D70; Aqueous phase: 1.64 and 2.36 M LiCI; O/A = 3/1; 20 °C; 4 min
Example 30: Individual operation of 2-stage counter-current for the first solvent extraction step (SX1) in mixer-settlers
Continuous experiments were conducted in a PTFE lab-scale mixer-settler units (Rousselet Model UX 1.1) to evaluate a feasibility of upscaling the developed solvent extraction. A mixer-settler battery comprised of a mixer volume of 35 mL, a settler volume of 143 mL with a settler area of 49 cm2. In each settler, one baffle and two PTFE coalescence plates were present to accelerate phase disengagement. Peristaltic pumps (Masterflex L/S®) were used for pumping the organic phase and the aqueous phase (Figure 6). Solvent extraction conversion of [A33][OR] was operated in counter-current mode using two stages in aqueous continuous mode. The flowrates were adjusted at 12 mL min-1 for aqueous phase and 6 mL min-1 for organic phase. Samples were taken every 30 min. The chloride concentration was monitored by titration method with 0.05 M AgNCh. The conclusion to be drawn from Table 43 is that the equilibrium state was reached 68% within 2 hours. The system remained stable for 8 hours. Afterwards, the flowrate was reduced a half to 6 mL min-1 for aqueous phase and 3 mL min-1 for organic phase to increase the retention time. Accordingly, more than 90% conversion of [A336][OR] was achieved after two more hours of operation. Neither precipitation nor third-phase formation was observed. The loaded organic phase at equilibrium was collected and used for the second solvent extraction conversion of LiCI to LiOH in another block of mixer-settlers. Table 43. Effect of the equilibration time in the individual operation of the first solvent extraction step (SX1) in 2-stage counter-current mixer-settlers5
%E-SX1 _
Time (h) Stage 1_ Stage 2
0.3 28.5 38.4
0.5 28.8 58.3
0.7 22.6 63.6
0.8 23.6 63.3
1.0 26.4 66.7
2.0 25.4 68.2
3.0 26.0 68.2
4.0 27.9 66.4
8.0 29.5 65.7
9.0 27.3 65.1
10.0 26.7 77.5
10.5 29.1 90.8 11.0 20.5 91.5 12.0 20.5 91.2
12.5 22.9 90.8 13.0 22.9 90.5
§Organic phase: 0.65 M [A336] [Cl] and 2,6-di-tert-butylphenol (molar ratio = 1: 1) in Shellsol D70; Aqueous phase: 2.0 M NaOH; O/A = 1/2; 20°C; retention time = 2-4 min; 1000 rpm.
Example 31: Individual operation of 6-stage counter-current for the second solvent extraction step (SX2) in mixer-settlers
In the same way as in Example 30, the solvent extraction conversion of LiOH was operated counter-current mixer-settlers using six stages at O/A = 3/1 (Figure 7). The organic phase containing 0.58 M [A336][OR] in Shellsol D70 was obtained from Example 30. The flow rate was setup 2 mL min-1 for aqueous phase and 6 mL min-1 for organic phase. The mixer-settler was operated in organic-continuous mode. Aqueous samples were taken every 60 min and were titrated with 0.1 M HNO3 to determine the LiOH concentration. From Table 44 it can be concluded that the percentage conversion %E- SX2 increased with time. A steady state was achieved after 5 hours of operation with more than 81% LiOH conversion. An attempt to enhance the conversion efficiency of LiOH by reducing the flowrate or increasing the retention time from 4 to 9 min. However, the conversion achieved was slightly lower. Table 44. Effect of the equilibration time in the individual operation of the second solvent extraction step (SX2) in 6-stage counter-current mixer- settlers5
%E- SX2
Time (h) Stage 1 Stage 2 Stage 3 Stage 4 Stage 5 Stage 6
0.5 8.3 15.3 21.0 24.2 26.4 45.1
1.0 7.2 14.6 20.4 24.5 34.8 56.8
2.0 6.6 13.6 21.8 31.3 48.8 70.8
3.0 7.0 15.0 25.0 37.9 56.5 76.3
4.0 7.6 16.6 27.8 41.0 59.3 81.1
5.0 9.3 19.5 31.4 45.4 63.1 82.8
6.0 9.8 22.0 34.4 48.8 66.2 84.8
7.0 10.9 22.8 35.4 49.7 68.0 85.7
8.0 12.1 24.3 37.7 52.6 69.2 86.3
9.0 13.0 25.1 37.6 52.2 67.3 83.8
10 12.8 26.5 39.0 51.7 67.0 84.4
11 12.8 25.4 40.0 55.3 67.0 83.1
12 11.2 23.7 38.4 53.6 66.7 81.8
§Organic phase: 0.58 M [A336][OR] in Shellsol D70; Aqueous phase: 1.64 M LiCI; O/A = 3/1;
20 °C; retention time = 4-9 min; 1000 rpm Example 32: Simultaneous operation of 2-stage and 6-stage counter-current for the first and second solvent extraction step (SX1) and (SX2) in mixer- settlers
In the same way as in Example 30 and 31, the first and second solvent extraction steps (SX1) and (SX2) were carried out simultaneously in mixer-settlers (Table 42). Two blocks of mixer-settler were connected so that the organic phase obtained in
SX1 could enter immediately the SX2. The stirring speed was adjusted to 1500 rpm
(Figure 8). The conclusion to be drawn from Table 45 is that the percentage conversion %E-SX1 remained stable, while the %E- SX2 increased with the time up to 6 hours of operation. At equilibrium, more than 87.1% [A336][OR] was formed in SX1 and 98.5% LiOH was converted in SX2. Higher concentrations of LiOH were achieved because the organic phase [A336][OR] was fresh. The back-conversion of LiOH to U2CO3 was less than 7%. This result demonstrates that it is possible to produce a highly pure LiOH solution with the proposed method. Table 45. Contact time for the simultaneous operation of the first and second solvent extraction step (SX1) and (SX2) in mixer-settlers5
Time (h) %E-SX1 %E- SX2 1.0 87.3 26.1 2.0 87.9 59.7 3.0 87.3 75.6 4.0 89.0 84.4 5.0 87.6 89.1 6.0 86.8 93.3 7.0 86.5 95.7 8.0 86.8 95.9 9.0 87.1 98.5
§(SX1) Organic phase: 0.65 M [A336] [Cl] and 2,6-di-tert-butylphenol (molar ratio = 1: 1) in Shellsol D70; Aqueous phase: 2.0 M NaOH; O/A = 1/2; 20 °C; retention time in each mixer = 2 min; 1500 rpm; (SX2) Organic phase: 0.58 M [A336][OR] in Shellsol D70; Aqueous phase: 1.64 M LiCI; O/A = 3/1; 20°C; retention time in each mixer = 4 min; 1500 rpm
Example 33. Comparative study for the simultaneous operation of the first and second solvent extraction step (SX1) and (SX2) in mixer-settlers using fresh and regenerated organic phase
In the same way as in Example 32, the operation of mixer-setters was tested with regenerated organic phase (Table 46). In particular, the mixture of [A336][CI]/phenol obtained after SX2 (Example 32) was regenerated and was used as organic feed solution of SX1, where it was re-contacted with NaOH solution (Figure 9). The conclusion to be drawn from Table 46 is that it is possible to regenerate the organic phase during the operation of mixer-settlers with a similar performance of the fresh one. In both cases, the back-conversion of LiOH to U2CO3 was less than 8%.
Table 46. Comparative study for the simultaneous operation of the first and second solvent extraction step (SX1) and (SX2) in mixer-settlers with fresh and regenerated organic phase5
Organic phase %E- SX1 %E- SX2 Fresh 87.1 98.5
Recycled _ 89.2 _ 98.3 _
§(SX1) Organic phase: 0.65 M [A336] [Cl] and 2,6-di-tert-butylphenol (molar ratio = 1: 1) in Shellsol D70; Aqueous phase: 2.0 M NaOH; O/A = 1/2; 20°C; retention time in each mixer= 2 min; 1500 rpm; (SX2) Organic phase: 0.58 M [A336][OR] in Shellsol D70; Aqueous phase: 1.64 M LiCI; O/A = 3/1; 20°C; retention time in each mixer = 4 min; 1500 rpm Example 34: Crystallization of lithium hydroxide monohydrate from the final product solution
The crystallization of lithium hydroxide monohydrate from the final LiOH product solution (obtained in Examples 32 and 33) was compared using different methods: crystallization via evaporation and antisolvent precipitation. In the first method, water was evaporated at 90 °C until the solid lithium hydroxide crystallized out. In the second method, lithium hydroxide was recovered from the aqueous solution by adding a known amount ethanol or isopropanol. The volume ratio of alcohol/aqueous solution was varied in the range 0.5 to 7.0. In all cases, the crystals of UOH-H20 were filtered off using a Buchner funnel under reduced pressure on Whatman filter paper (0.45 m), washed with the corresponding solvents, and dried at 110 °C for 3 hours. Table 47 shows the mass recovery and the purity of lithium hydroxide monohydrate obtained by different methods. The conclusion to be drawn is that both the conventional evaporation-induced crystallization and antisolvent precipitation method can be used for obtaining UOH-H20 with high purity (>99.8%). More than 87% LiOH-H20 was crystallized by evaporation of water. The mass loss is largely caused by the washing step. Quantitative recovery of LiOH-H20 was achieved by addition of isopropanol, but not of ethanol, to the aqueous LiOH solution. The extremely low solubility of LiOH in isopropanol (<0.11 mg/ 100 mL) accelerates the LiOH-H20 precipitation in the solution. Furthermore, the mass recovery of LiOH is dependent on the volume ratio of Visopropanoi/Vgqueous, as shown in Table 48. In practice, a combination of partial evaporation of water and antisolvent precipitation is a preferred method to produce highly pure LiOH-H20 at a reasonable cost.
Table 47. Crystallization of lithium hydroxide monohydrate from the final product solution5 _
Evaporation Antisolvent precipitation _
Method _ Water _ Ethanol _ Isopropanol _
Recovery yield (%) 87.3 3.87 ^4.6
Purity (%) _ 99 3 _ 100 _ _
§Final product solution: 1.55 M LiOH and 0.025 M LiCI. Table 48. Effect of volume ratio Vaicohoi/Vaq on the mass recovery of lithium hydroxide (as the monohydrate)
Mass recovery (%)
Valcohol/Vaq
Ethanol Isopropanol
0.5 6.32 5.94
1.0 6.06 6.06
1.5 5.81 6.39
2.0 5.29 9.30
3.0 4.19 50.6
4.0 4.13 77.5
5.0 4.45 86.7
6.0 3.16 92.4
7.0 3.87 94.6

Claims

1. A process for refining lithium chloride comprising the steps of: 1) dissolving a solid composition comprising less than 98 % (w/w) lithium chloride in an organic solvent thereby selectively dissolving lithium chloride;
2) adding an alkali metal hydroxide solubilised in an organic solvent;
3) optionally further purifying lithium chloride in the organic solution by non- aqueous ion exchange; and 4) removing the organic solvent.
2. The process according to claim 1, further comprising the step of:
5) converting an aqueous solution comprising lithium chloride obtained in step c) into an aqueous solution comprising lithium hydroxide by solvent extraction with a mixture of a basic extractant in the chloride form and a weakly acid compound.
3. The method according to claim 2, further comprising the step of:
6) crystallizing lithium hydroxide monohydrate from the aqueous lithium hydroxide solution, or converting lithium hydroxide in the aqueous solution to lithium carbonate by reaction with carbon dioxide gas.
4. The process according to any one of claims 1 to 3, wherein step 4 comprises:
- recovering lithium chloride by evaporation of the solvent of the organic phase and crystallization of lithium chloride from the aqueous phase, or
- recovering lithium chloride aqueous phase by evaporation of the solvent of the organic phase and drying the aqueous phase comprising lithium chloride.
5. The process according to any one of claims 1 to 4, where the composition comprising lithium chloride of step 1) is obtained from salt lake brines, from geothermal brines, or from a lithium ore.
6. The process according to any one of claims 1 to 5, where the organic solvent in step 1) and or in step 2) is an alcohol.
7. The method according to claim 6, wherein the alcohol comprises between 1 and 5 % (v/v) water.
8. The method according to claim 6 or 7, wherein the alcohol is ethanol.
9. The process according to any one of claims 1 to 8, where the organic solvent in step 1) and in step 2 is the same organic solvent.
10. The process according to any one of claims 1 to 9, wherein the alkali metal hydroxide is step 2) is lithium hydroxide.
11. The process according to any one of claims 1 to 10, wherein non-soluble material is removed after step 2)
12. The process according to any one of claims 1 to 12 where said non-aqueous ion exchange in step 3) is performed with a cation exchanger.
13. The process according to any one of claim 1 to512, where said non-aqueous ion exchange in step 3) is performed with a strongly acidic ion exchange resin of cross-linked polystyrene with sulfonic acid functional groups, with an inorganic ion-exchange material, or with an ionic exchanger in the H+ form or in the Li+ form, or with a weakly basic ion exchange resin of cross-linked polystyrene with tertiary amine functional groups.
14. The process according to any one of claims 2 to 13, where the solvent extraction for the chloride-hydroxide anion exchange of step 5) is carried out by a mixture of a quaternary ammonium chloride and an alcohol or phenol or a phenol derivate in a diluent.
15. The process according to any one of claims 14, where the phenol derivative is selected from the group consisting of 4-tert-butylphenol, 2,4-di-tert- butylphenol, 2,6-di-tert-butyl-4-methylphenol, 2,6-di-tert-butylphenol, 2,6- dimethylphenol, 3,5-dimethylphenol, 2,4,6-tri-tert-butylphenol, 4- nonylphenol, nonylphenol (isomeric mixture) and cresol.
16. The process according to claim 14 or 15, where the solvent extraction process for the chloride-hydroxide anion exchange is carried out by a mixture of tricaprylmethylammonium chloride and 2,5-dimethyl-2,5-hexanediol in a diluent.
17. The process according to claim 14 or 15 where the solvent extraction for the chloride-hydroxide anion exchange of step 5) is carried out by a mixture of tricaprylmethylammonium chloride and 4-tert-butylphenol in a diluent, a mixture of tricaprylmethylammonium chloride and 2,6-di-tert-butylphenol in a diluent or by a mixture of tricaprylmethylammonium chloride and nonylphenol in a diluent.
18. The method according to claim 2, further comprising the step of recovering of lithium hydroxide monohydrate from aqueous solution by antisolvent precipitation with an alcohol.
19. The method according to claim 18, wherein the alcohol is isopropanol.
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Citations (22)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US2561439A (en) 1947-07-08 1951-07-24 Union Carbide & Carbon Corp Method of treating lithiferous ores to recover lithium as lithium chloride
US3278260A (en) 1963-02-08 1966-10-11 Kerr Mc Gee Oil Ind Inc Preparation and purification of lithium chloride
US3410653A (en) 1966-02-07 1968-11-12 Corning Glass Works Alkali metal salt separation
US4036713A (en) 1976-03-04 1977-07-19 Foote Mineral Company Process for the production of high purity lithium hydroxide
US4207297A (en) 1978-03-27 1980-06-10 Foote Mineral Company Process for producing high purity lithium carbonate
US4274834A (en) 1979-04-11 1981-06-23 Foote Mineral Company Process for purification of lithium chloride
US4980136A (en) 1989-07-14 1990-12-25 Cyprus Foote Mineral Company Production of lithium metal grade lithium chloride from lithium-containing brine
IL133622A (en) 1999-12-20 2005-11-20 Yissum Res Dev Co Process for the preparation of metal and quaternary onium hydroxides
US7329354B2 (en) 1998-06-09 2008-02-12 Ppt Technologies, Llc Purification of organic solvent fluids
US20110044882A1 (en) 2008-04-22 2011-02-24 David Buckley Method of making high purity lithium hydroxide and hydrochloric acid
EP2487136A1 (en) * 2011-01-20 2012-08-15 Rockwood Lithium Inc. Production of high purity lithium compounds directly from lithium containing brines
WO2014078908A1 (en) * 2012-11-23 2014-05-30 Ady Resources Limited Process for recovering lithium from a brine with reagent regeneration and low cost process for purifying lithium
US20170197204A1 (en) 2014-06-03 2017-07-13 Arkema France Method for eliminating metal ions from a viscous organic solution
WO2017121663A1 (en) 2016-01-12 2017-07-20 Umicore Lithium-rich metallurgical slag
CN108004420A (en) * 2017-12-08 2018-05-08 中国科学院青海盐湖研究所 The technique that lithium is extracted from the bittern of alkalescence containing lithium based on centrifugal extractor
CN109574047A (en) * 2018-10-22 2019-04-05 天齐锂业(江苏)有限公司 The method of high-purity Sustiva and lithium chloride is recycled from medicine waste liquid containing lithium
US20190169038A1 (en) 2016-08-08 2019-06-06 Lithium Extraction Technologies (Australia) Pty Lt D Process For The Conversion Of Lithium Phosphate Into A Low Phosphate Lithium Solution Suitable As Feedstock For The Production Of Saleable Lithium Products And For The Recovery Of Phosphorous For Re-Use In The Production Of Lithium Phosphate
US20190248667A1 (en) * 2017-06-15 2019-08-15 Energysource Minerals Llc System and process for recovery of lithium from a geothermal brine
US20190256368A1 (en) * 2017-06-15 2019-08-22 Energysource Minerals Llc Process for selective adsorption and recovery of lithium from natural and synthetic brines
WO2019220003A1 (en) 2018-05-18 2019-11-21 Outotec (Finland) Oy Method for recovering lithium hydroxide
CN110669938A (en) * 2019-11-08 2020-01-10 湘潭大学 Extraction system for separating magnesium from magnesium-containing brine by using secondary amide/alkyl ketone composite solvent and extracting lithium, extraction method and application thereof
WO2020010464A1 (en) 2018-07-12 2020-01-16 Element Ai Inc. Automated generation of documents and labels for use with machine learning systems

Patent Citations (22)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US2561439A (en) 1947-07-08 1951-07-24 Union Carbide & Carbon Corp Method of treating lithiferous ores to recover lithium as lithium chloride
US3278260A (en) 1963-02-08 1966-10-11 Kerr Mc Gee Oil Ind Inc Preparation and purification of lithium chloride
US3410653A (en) 1966-02-07 1968-11-12 Corning Glass Works Alkali metal salt separation
US4036713A (en) 1976-03-04 1977-07-19 Foote Mineral Company Process for the production of high purity lithium hydroxide
US4207297A (en) 1978-03-27 1980-06-10 Foote Mineral Company Process for producing high purity lithium carbonate
US4274834A (en) 1979-04-11 1981-06-23 Foote Mineral Company Process for purification of lithium chloride
US4980136A (en) 1989-07-14 1990-12-25 Cyprus Foote Mineral Company Production of lithium metal grade lithium chloride from lithium-containing brine
US7329354B2 (en) 1998-06-09 2008-02-12 Ppt Technologies, Llc Purification of organic solvent fluids
IL133622A (en) 1999-12-20 2005-11-20 Yissum Res Dev Co Process for the preparation of metal and quaternary onium hydroxides
US20110044882A1 (en) 2008-04-22 2011-02-24 David Buckley Method of making high purity lithium hydroxide and hydrochloric acid
EP2487136A1 (en) * 2011-01-20 2012-08-15 Rockwood Lithium Inc. Production of high purity lithium compounds directly from lithium containing brines
WO2014078908A1 (en) * 2012-11-23 2014-05-30 Ady Resources Limited Process for recovering lithium from a brine with reagent regeneration and low cost process for purifying lithium
US20170197204A1 (en) 2014-06-03 2017-07-13 Arkema France Method for eliminating metal ions from a viscous organic solution
WO2017121663A1 (en) 2016-01-12 2017-07-20 Umicore Lithium-rich metallurgical slag
US20190169038A1 (en) 2016-08-08 2019-06-06 Lithium Extraction Technologies (Australia) Pty Lt D Process For The Conversion Of Lithium Phosphate Into A Low Phosphate Lithium Solution Suitable As Feedstock For The Production Of Saleable Lithium Products And For The Recovery Of Phosphorous For Re-Use In The Production Of Lithium Phosphate
US20190248667A1 (en) * 2017-06-15 2019-08-15 Energysource Minerals Llc System and process for recovery of lithium from a geothermal brine
US20190256368A1 (en) * 2017-06-15 2019-08-22 Energysource Minerals Llc Process for selective adsorption and recovery of lithium from natural and synthetic brines
CN108004420A (en) * 2017-12-08 2018-05-08 中国科学院青海盐湖研究所 The technique that lithium is extracted from the bittern of alkalescence containing lithium based on centrifugal extractor
WO2019220003A1 (en) 2018-05-18 2019-11-21 Outotec (Finland) Oy Method for recovering lithium hydroxide
WO2020010464A1 (en) 2018-07-12 2020-01-16 Element Ai Inc. Automated generation of documents and labels for use with machine learning systems
CN109574047A (en) * 2018-10-22 2019-04-05 天齐锂业(江苏)有限公司 The method of high-purity Sustiva and lithium chloride is recycled from medicine waste liquid containing lithium
CN110669938A (en) * 2019-11-08 2020-01-10 湘潭大学 Extraction system for separating magnesium from magnesium-containing brine by using secondary amide/alkyl ketone composite solvent and extracting lithium, extraction method and application thereof

Non-Patent Citations (10)

* Cited by examiner, † Cited by third party
Title
BUKOWSKY H ET AL: "THE SEPARATION OF CALCIUM AND MAGNESIUM FROM LITHIUM CHLORIDE BY LIQUID-LIQUID EXTRACTION WITH DI(2-ETHYLHEXYL)PHOSPHORIC ACID", HYDROMETALLURGY, ELSEVIER SCIENTIFIC PUBLISHING CY. AMSTERDAM, NL, vol. 28, no. 3, 1 April 1992 (1992-04-01), pages 323 - 329, XP000274419, ISSN: 0304-386X, DOI: 10.1016/0304-386X(92)90037-Z *
DANG ET AL., SEPARATION AND PURIFICATION TECHNOLOGY, vol. 233, 2020
DEHMLOW ET AL., TETRAHEDRON, vol. 41, 1985, pages 2927 - 2932
FLEMINGMONHEMIUS, HYDROMETALLURGY, vol. 4, 1979, pages 159 - 167
H. BUKOWSKY ET.AL.: "The recovery of pure lithium chloride from "brines" containing higher contents of calcium chloride and magnesium chloride", HYDROMETALLURGY, vol. 27, 1 January 1991 (1991-01-01), pages 317 - 325, XP055739628 *
HIGGINS, INDUSTRIAL AND ENGINEERING CHEMISTRY, vol. 53, 1961, pages 635 - 637
LAVI ET AL., INDUSTRIAL & ENGINEERING CHEMISTRY RESEARCH, vol. 40, 2001, pages 6045 - 6050
MOODYTHOMAS, THE ANALYST, vol. 93, 1968, pages 557 - 588
ROTHENBERG ET AL., CHEMICAL COMMUNICATIONS, 2000, pages 1293 - 1294
TABOADA ET AL., CHEMICAL ENGINEERING RESEARCH AND DESIGN, vol. 85, 2007, pages 1325 - 1330

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