Method for selectively extracting salts to be extracted from salt water or brine

Abstract

Method for selectively extracting salts to be extracted from brine The present invention relates to a method for selectively extracting a salt to be extracted from a brine to be treated containing the salt, wherein the brine to be treated (S1) is sent to an initial extraction agitator-mixer (1a) in which a liquid hydrophobic organic phase is also introduced, the brine and the liquid hydrophobic organic phase are mixed in the initial extraction agitator-mixer (1a), and the mixture is sent to an initial decanter-separator (1d, 1dc) to obtain a brine freed of the salt to be extracted and an organic phase loaded with the salt to be extracted, and the brine freed of the salt to be extracted is recovered and used for extraction. The organic phase loaded with salts to be extracted is sent to an initial wash agitator-mixer-decanter-centrifuge (5a, 5dc or 7a, 7dc) to obtain a purified organic phase loaded with salts to be extracted, which is then sent to the so-called "organic" heat exchanger (1e) where it is heated before being sent to the so-called initial "regeneration" column (1c; 2c) where it undergoes countercurrent exchange with hot water at a temperature higher than that of the initial brine to obtain water containing the salts to be extracted and a regenerated organic phase depleted of the salts to be extracted, which is recycled to the initial extraction agitator-mixer (1a).

Description

[Technical Field]

[0001] The present invention relates to a method for selectively extracting a salt to be extracted from a salt water or brine containing the salt to be extracted and other salts.

[0002] Upgrading of alkali, alkaline earth, transition metal or rare earth metal salts solubilized in natural or leach brines is still a relatively complex and expensive operation, has a significant environmental impact and there is still significant room for improvement.

[0003] Historically, this has been based on the separation of various salts by crystallization as a function of temperature, for example at 50°C, in conjunction with differences in aqueous solubility according to the following increasing order of saturation molar concentration: NaF <NaHCO3≒BaCl2<Na2SO4<MgSO4<ZnSO4<KCl<MgCl2<NaCl<CdCl2<MnCl2<KNO3<NH4Cl<CaCl2<CsCl<LiCl<ZnCl2<NH4NO3…

[0004] Naturally, for each brine to be treated, the number of ions and salts present is generally much smaller, which results in sufficiently large differences in solubility to improve their separation capabilities. Indeed, for example, with regard to the separation of alkali and alkaline earth salts, the implementation of solar evaporation tanks or thermal crystallization systems allows for the continuous precipitation of different salts depending on their initial concentration and their solubility in water. However, this creates a high degree of variability in the crystallization process due to the variability of the mineral composition of the source, their variability over time, and the degree of concentration of the brine to be treated. This approach requires very good knowledge of the crystallization and co-crystallization conditions of the salts involved. Despite this, the mineral salts resulting from this thermal crystallization are either pure, with low to moderate yields, or impure, requiring one or more additional downstream purification steps. Furthermore, the high energy cost of water evaporation (682 kWh at 20°C) th / ton), even if modern evaporation processes incorporating mechanical vapor compression and / or multiple-effect distillation were implemented industrially, this approach would require only 20-60 kWh of distilled water in terms of energy, i.e. elect / ton region, or for crystallized salt, if, for example, 75-150 tons of water need to be evaporated per ton of NaCl crystallized, this is still very costly, in the region of 100 times higher energy. As a result, except for applications involving salts with high economic value, solar evaporation ponds are used instead, which have the advantage of offering zero OpenX in terms of water evaporation energy, at the expense of significant local water consumption, since over 90% of the water in the brine to be treated is evaporated. This approach is acceptable for water evaporation by sea, but much less for operations in deserts, arid regions, or landlocked areas.

[0005] Depending on the chemical composition of the metals in the water and solution to be recovered and upgraded, other separation strategies, especially hydrometallurgy, acid (HSO 4) Alternatively, liquid route metal treatment processes have been developed which include a stage in which the metal is solubilized for its purification using oxidizing agents (Cl2, H2O2...). This leaching or dissolution is used in particular for some transition metals such as zinc, copper, nickel, cobalt, manganese, rare earths and uranium.

[0006] Once solubilized, some of the metals may precipitate as insoluble hydroxides, carbonates, or sulfides. This is a precipitation-based approach. These precipitates essentially depend on the pH of the solvent, the solubility products, and the redox potential of the medium. Most frequently, the pH of the solution is raised to form the precipitated hydroxides. The metals are then recovered as solids by simple decantation, within the limits of their solubility product, often resulting in concentrations at the process outlet that exceed environmental discharge standards. In addition to the potential need for further downstream processing, this approach may require large amounts of reagents (neutralizing bases and acids, or Na2CO3, etc.), create operational difficulties due to scaling issues, or require upstream pretreatment to allow precipitation of products of acceptable purity. All of these make this approach unattractive from an economic standpoint, while potentially generating large amounts of waste, which are often stored in residual impoundments.

[0007] Another commonly used approach is ion exchange, in which an aqueous solution is mixed with the metal to be extracted, usually with protons H + or sodium Na + One drawback of this approach is that the ion exchange is + When ion exchange is carried out in H, the ion exchange resin, whether solid or liquid, undergoes an initial regeneration with acid to desorb the metal ions, and is then neutralized with NaOH before it can be used again for extraction. +When performed directly with ionic metals, the pH of the water in contact with them drops, necessitating the use of reagents for extraction, washing, and regeneration of the substances (neutralizing acids and / or bases), which can be large volumes. In both of these cases, this results in significant operating costs and potentially the installation of local chlorine and caustic soda plants to produce these acids and bases. Another drawback can arise from the regeneration or acid washing of the ion exchange material, which, depending on the initial composition of the selected ion exchange material, can also undergo its own decomposition over time, if it is inorganic, with the potential for leaching and therefore the release of transition metals and / or heavy metals Mn, Ti, Sn, Cu, Al, V, or Sb into the environment. This is all the more problematic because the ecotoxicity thresholds of these metals are very low and difficult to completely capture.

[0008] Finally, we note that there is no direct extraction of neutral salts, as well as the systematic addition of salts / waste products such as NaSO resulting from the overall combination of NaOH and HSO to the system and / or its environment.

[0009] Another approach is salt adsorption, which involves physically or electrochemically immobilizing metal salts on the surface or inside the pores of an adsorbent material. Very few applications have been developed to date, but examples include the selective adsorption of lithium for the desalination of low-salinity brackish water and capacitive deionization. Economically, adsorption is limited to the removal of small amounts of ions due to the very large porous adsorbent surfaces required, i.e., less than 5 g NaCl / L of solution for capacitive deionization (CDI) and less than 6.5 g LiCl / L for selective extraction by adsorption. Summary of the Invention

[0010] The object of the present invention is to obtain a technical and economic solution that allows the selective extraction of the salt to be extracted over a wide concentration range, with a high extraction yield, an energy consumption close to the theoretical minimum (associated with the extraction of salt from water), the ability to produce this extracted salt with high purity, a water consumption also close to the theoretical minimum, and zero consumption of reagents.

[0011] To this end, the present invention relates to a method for selectively extracting a salt to be extracted, which comprises at least one cation and at least one anion, from a salt water or brine to be treated, which comprises the salt to be extracted and salts other than the salt to be extracted, The brine to be treated is sent to a so-called initial extraction agitator-mixer in which a liquid hydrophobic organic phase is also introduced, said organic phase at least a first hydrophobic protic organic compound that solvates the anion of the salt to be extracted and has a pKa in water at 25°C of at least 9; at least a second hydrophobic organic compound that selectively extracts the cations of the salt to be extracted and has a complexation constant for the cation to be extracted, the log K value of which in methanol at 25°C is greater than 1, preferably greater than 2; Including, - the brine and the liquid hydrophobic organic phase are mixed in the extractor agitator-mixer, and the mixture is the brine from which the salt to be extracted has been removed; an organic phase charged with the salt to be extracted; is sent to a so-called initial extractor decanter-separator to obtain - the brine from which the salt to be extracted has been removed is recovered, The organic phase leaving the initial decanter-separator is sent to a new agitator-mixer, the so-called initial wash agitator-mixer, where it is mixed with wash water and the mixture is a purified organic phase charged with the salt to be extracted, which is sent to a so-called "organic" heat exchanger and heated there; Initial extraction agitator - Washing water filled with impurities sent to the mixer, is sent to a decanter-centrifuge, the so-called initial wash decanter-centrifuge, to obtain the organic phase, which is filled with the salt to be extracted and heated in a so-called "organic" heat exchanger, water containing the salt to be extracted, a regenerated organic phase depleted of the salts to be extracted; into a so-called initial regeneration column in which a countercurrent exchange is carried out with hot water at a temperature higher than that of the initial brine, in order to obtain the organic phase loaded with the salt to be extracted is heated in a heat exchanger by the organic phase leaving the regeneration column, - Characterized in that the regenerated organic phase is recycled to the initial extraction agitator-mixer.

[0012] In particular, the salts to be extracted can be extracted at low temperatures, i.e. at temperatures between −35° C. and +45° C. Regeneration in a regeneration column is then carried out at a higher temperature.

[0013] In particular, the organic phase may be washed at a temperature between 5°C and 45°C.

[0014] The temperature difference between the extraction step and the regeneration step may be at least 30° C., preferably above 40° C., even more preferably above 50° C. The regeneration step may be carried out at a temperature between 60° C. and 150° C.

[0015] In particular, the organic phase exiting the regeneration column can be heated before entering the "organic" heat exchanger.

[0016] In particular, the first organic compound that solvates the anion of the salt to be extracted can be selected from N-(3,4-dichlorophenyl)octanamide, N-(3,5-dichlorophenyl)octanamide and N-(3,4,5-trifluorophenyl)octanamide.

[0017] In certain embodiments, the second organic compound has the formula: [ka] and a lithium cation extractant compound selected from the group consisting of R1 and R2, which may be identical or different, are straight-chain or branched C1-C, regardless of their position on the nitrogen atom. 12 independently selected from alkyl, aryl, C4-C8 cycloalkyl, or - R1 and R2 together with the nitrogen atom carrying them form a 5-, 6-, 7- or 8-membered ring; -R3 is selected from hydrogen, linear or branched C1-C8 alkyl, C4-C8 cycloalkyl, C2-C6 alkoxyalkyl and alkoxyalkylaryl; -R4 is selected from hydrogen, linear or branched C1-C3 alkyl; -R5 is selected from hydrogen, linear or branched C1-C3 alkyl; -R6 is selected from hydrogen, straight or branched C1-C3 alkyl.

[0018] In certain embodiments, R1 and R2, whatever their position on the nitrogen atom, may be independently selected from methyl, ethyl, n-propyl, iso-propyl, n-butyl, iso-butyl, sec-butyl, tert-butyl, n-pentyl, iso-pentyl, 2-methylbutyl, 2-ethylpropyl, n-hexyl, n-heptyl, n-octyl, n-nonyl, n-decyl, 2-ethylhexyl, phenyl, cyclobutyl, cyclopentyl, cyclohexyl, cycloheptyl, cyclooctyl, or - R1 and R2 together with the nitrogen atom that carries them can form a pyrrolidine, piperidine, azepane or azocane ring; -R3 can be selected from hydrogen, methyl, ethyl, propyl, isopropyl, n-butyl, isobutyl, sec-butyl, tert-butyl, n-pentyl, isopentyl, 2-methylbutyl, 2-ethylpropyl, n-hexyl, cyclohexyl, methoxymethyl, methoxyethyl, methoxypropyl, methoxybutyl and -CH2-O-CH2-phenyl, and R4, R5 and R6 can be hydrogen or methyl; R1 and R2 are advantageously selected from butyl, pentyl, hexyl, heptyl, octyl, nonyl, decyl or phenyl, if the brine to be treated has a calcium concentration of more than 10 g / L, and / or Li + / Ca 2+ Selectivity is preferred, and R1 and R2 are advantageously selected from isopropyl, isobutyl, sec-butyl, tert-butyl, isopentyl, 2-methylbutyl, 2-ethylpropyl, cyclobutyl, cyclopentyl, cyclohexyl, cycloheptyl, cyclooctyl or R1 and R2, together with the nitrogen atom that carries them, when the brine to be treated has a calcium concentration of less than 10 g / L and / or Li + / Na + If selectivity is preferred, a pyrrolidine, piperidine, azepane or azocane ring is formed.

[0019] The lithium cation extractant compound can be selected from, but is not limited to: [Table A-1] [Table A-2] [Table A-3] [Table A-4]

[0020] In particular, the second organic compound extracting the cation of the salt to be extracted can be selected from 4-tert-butyl-calix[4]arenoic acid tetraethyl ester, 4-tert-butyl-calix[6]arenoic acid hexaethyl ester, 2-[2,2-bis[[2-(dicyclohexylamino)-2-oxoethoxy]methyl]butoxy]-N,N-dicyclohexyl-acetamide and N,N,N',N',N",N"-hexacyclohexyl-4,4',4"-propylidinetris(3-oxabutyramide).

[0021] The organic phase may also contain a hydrophobic polar organic diluent, preferably an aromatic diluent which may be selected from polar aromatic compounds such as 1-chloro-2-bromo-benzene, 1,2-dibromobenzene, 2-bromotoluene or 3,4-dibromotoluene.

[0022] The first organic compound that solvates the anions of the salt to be extracted may be present in the liquid hydrophobic organic phase at a concentration of at least 0.1 mol / L, the relative molar ratio of the solvating molecules of the complementary anionic species to the extracting molecules of the cationic species being greater than 1.

[0023] The second organic compound that extracts the cation of the salt to be extracted may be present in the liquid hydrophobic organic phase at a concentration of at least 0.1 mol / L.

[0024] Before recovering the brine freed of the salts to be extracted, the brine may be sent to another extraction agitator-mixer, followed by an extraction decanter-separator receiving a regenerated organic phase, after decantation, in order to obtain a brine with even less salt to be extracted, which is recovered or sent again to another extraction agitator-mixer and then sent to an extraction decanter-separator in order to obtain a brine with even less salt to be extracted than before, the operation of extracting the salts to be extracted from the brine may be repeated at least once, characterized in that the organic phase loaded with the salts to be extracted is returned from a given decanter-separator to a low-level agitator-mixer in order to be recovered at the outlet of the initial decanter-separator, and the organic phase loaded with salt is sent to a so-called "organic" heat exchanger.

[0025] In particular, the multiple extractor agitator-mixer-decanter-separators can operate in countercurrent flow: the aqueous phase entering the extractor agitator-mixer is brought into contact with the organic phase coming from the upper extractor decanter-separator.

[0026] There may be at least one other regeneration column connected in parallel to the initial regeneration column, which is fed with part of the flow from the so-called "organic" heat exchanger.

[0027] In particular, the regeneration column may be a stirred column or a pulsed column.

[0028] The water that will serve as regeneration water can be sent to a so-called "water" heat exchanger, heated by water loaded with the salts to be extracted, leaving the regeneration column or columns to obtain heated regeneration water that can be further heated before being introduced into the regeneration column or columns, if several are provided, and cooling water loaded with the salts to be extracted is obtained from said heat exchanger.

[0029] Before being sent to the initial extraction agitator-mixer, the impurity-laden wash water is Initial Extraction Agitator - Retained or concentrated or raffinate brine sent to the mixer; and Initial cleaning agitator - purified water fed back to the mixer can be sent to a purification unit to obtain

[0030] In particular, it is characterized in that before the washed organic phase leaving the so-called washing decanter-separator is sent to the so-called "organic" heat exchanger, it is sent after decantation to another so-called washing agitator-mixer in order to obtain an organic phase that has undergone a further washing step which is sent to the so-called "organic" heat exchanger, and subsequently to a so-called washing decanter-centrifuge which receives washing water, the operation of washing the organic phase being able to be repeated at least once, and the washing water loaded with impurities, or the retentate brine or concentrated brine or raffinate brine, is returned to the initial washing agitator-mixer after purification.

[0031] In particular, the multiple agitator-mixer-washing decanter-centrifuges may operate in countercurrent: the aqueous phase entering the washing agitator-mixer contacts the organic phase coming from the upper decanter-centrifuge.

[0032] The salt-laden water to be extracted, coming either from the regeneration column or from the so-called "water" heat exchanger, - concentrated and purified water loaded with salts to be extracted, -Purified water sent to either a regeneration column or a so-called "water" heat exchanger can be sent to a purification unit to obtain

[0033] The purification unit may be a reverse osmosis unit or a membrane distillation unit or an evaporation / condensation unit or a unit for absorption of water by a temperature regenerable organic phase, or a combination of at least two of these units.

[0034] -distilled water, -Freshwater, -Demineralized water, and - at least a portion of the concentrated and purified water charged with the salt to be extracted from the purification unit (2f), optionally after cooling, At least one of the above may be used as the wash water.

[0035] The initial decanter-separator and / or the wash decanter-separator, if present, may be a decanter-centrifuge and allow the removal of the phase composed of solid waste or the aqueous washout of the organic phase, or the separation of a liquid two-phase medium, the aqueous phase / organic phase ratio being less than 1.0 for the initial decanter-separator and less than 0.2 for the wash decanter-separator.

[0036] The salt to be selectively extracted can be selected from sodium salts such as NaCl, NaCN, NaNO3, NaBr, NaI, potassium salts such as KCl, KCN, KNO3, KBr, KI, lithium salts such as LiCl, LiCN, LiNO3, LiBr, LiI, or other salts that can be selectively extracted in the presence of other salts.

[0037] The salts that are preferentially extracted are the so-called diionic salts, i.e., Li + Cation, K + Cation or Na + cations with a single positive charge, such as the cation Cl, and the monatomic anion Cl - or polyatomic anion NO3 - and a monoatomic or polyatomic anion having a single negative charge, such as

[0038] The volume ratio of the washing water and the organic phase introduced into the so-called washing agitator-mixer can be between 0.01 and 0.1, preferably between 0.03 and 0.06.

[0039] The volume ratio of hot water to the organic phase introduced into the so-called regeneration column can be 0.05 to 0.2, preferably 0.08 to 0.15. [Brief explanation of the drawings]

[0040] [Figure 1]The installation shown in Figure 1 is designed to treat a saltwater or brine effluent S containing n dissolved salts composed of cations and anions, from which the salts to be extracted are at least extractable by a dedicated organic formulation effluent O circulating in a closed loop. [Figure 2] The installation shown in Figure 2 is identical to that shown in Figure 1, except that the decanter-separators 2d and 3d are replaced by centrifugal decanters 2dc and 3dc, and the reheat unit 1b, necessary to compensate for the heat losses at the cold ends of the exchangers 1e and 2e, is placed on the organic effluent O6 instead of the aqueous effluent A2. [Figure 3] The installation shown in Figure 3 is designed to treat a saltwater or brine effluent S containing n dissolved salts composed of cations and anions, from which the salts to be extracted are at least extractable by a dedicated organic formulation effluent O circulating in a closed loop. [Figure 4] The installation shown in Figure 4 is identical to that shown in Figure 3 with the addition of two further agitator-mixers 6a and 7a along with decanter-centrifuges 6dc and 7dc. [Figure 5] The installation of Figure 5 is identical to that of Figure 4, with the water recovery unit 2f moved from the low temperature zone to the high temperature zone of the secondary "water" heat exchanger 2e and the water recovery unit 1f removed. [Figure 6] By contacting this brine with the organic phase at volume ratios used for aqueous brine (O / A) extraction of 0.05, 0.2, 0.5, 1, and 4, respectively, at room temperature (20°C), a high relative extraction of sodium in the form of NaCl versus potassium or KCl was obtained, as shown in the graph in Figure 6. [Figure 7] This sodium extraction is selective without changing the potassium concentration, as shown in FIG. [Figure 8] Figure 8 shows the ion concentrations of various compounds in the organic phase after extraction. NaCl is predominant in the organic phase when the (O / A) extraction ratio is less than 2, and even more so when the (O / A) extraction ratio is low. [Figure 9]To demonstrate the very poor retention of KCl within the organic phase, formulations loaded with Na, K, and Cl were examined after extraction with an (O / A) extraction ratio of 0.5 and water washes at room temperature (20 °C) with very small amounts of water ranging from 1% to 20% of the volume of the organic phase, also referred to as the (A / O) wash ratio. The results are shown in Figure 9. [Figure 10] High temperature regeneration of the formulation obtained by passing NaCl through the aqueous phase and washing with 4% of the volume of water by contacting with hot water at 80°C in volumes of 2%, 5%, 10%, 20% and 50% of the volume of the organic phase was used to regenerate the organic phase, this ratio being denoted (A / O) regeneration. The results obtained are shown in Figure 10. [Figure 11] Figure 11 shows the ion concentrations of the brine obtained after regeneration at 80°C. After a wash step with an (A / O) wash ratio of 4%, the brine obtained from high temperature regeneration with these various (A / O) regeneration ratios consisted of 93% to 99% NaCl by mass, which allows for the purification of NaCl with minimal NaCl loss. [Figure 12] The results of cycle 4-1-3 are shown in FIG. [Figure 13] FIG. 13 shows the effect of adding a second countercurrent scrubbing stage at room temperature. [Figure 14] By contacting this brine at room temperature (20°C) with the organic phase at volume ratios used for aqueous brine (O / A) extraction of 0.1, 0.4, 0.8, 1.5 and 4, good relative extraction of potassium in the form of KCl relative to sodium or NaCl was obtained, as shown in the graph in Figure 14. [Figure 15] This extraction of potassium is selective without altering the sodium concentration, as shown in FIG. [Figure 16] FIG. 16 shows the ion concentrations of various compounds in the organic phase after extraction at 20°C. [Figure 17]To demonstrate very low retention of NaCl within the organic phase, formulations loaded with Na, K, and Cl were examined after extraction with an (O / A) extraction ratio of 0.8 and scrubbing at room temperature (20°C) with very small amounts of water ranging from 1% to 20% of the volume of the organic phase, also referred to as the (A / O) scrubbing ratio. The results are shown in Figure 17. [Figure 18] High temperature regeneration of the formulation obtained by washing with 4% of the volume of water was used to pass the KCl through the aqueous phase and regenerate the organic phase by contacting with hot water at volumes of 2%, 5%, 10%, 20%, and 50% of the volume of the organic phase at 80° C., this ratio being designated (A / O) regeneration. The results obtained are shown in Figure 18, which show that with 20% volume of water, 83.7% of the KCl was desorbed from the organic phase (compared to 14.8% previously at 20° C.). [Figure 19] FIG. 19 shows the ion concentrations of the brine obtained after high temperature regeneration at 80°C. [Figure 20] The results of cycle 4-1-3 are shown in FIG. [Figure 21] FIG. 21 shows the effect of adding a second countercurrent scrubbing stage at room temperature. [Figure 22] Figure 22 shows that the lithium yield can be as high as 97% and the purity of the LiCl produced depends on three complementary parameters: the (O / A) extraction ratio, the choice of cycle, and the (A / O) wash ratio. [Figure 23] FIG. 23 shows that the lithium yield can be as high as 97% and the purity of the LiCl produced depends on three complementary parameters: the (O / A) extraction ratio, the choice of cycle, and the (A / O) wash ratio. [Figure 24] Figure 24 provides a summary of the optimum dimensions adopted for selective lithium extraction of Maricunga brine for cycles with 4 extraction stages and 3 theoretical regeneration stages for 0, 1, 2 and 3 wash stages, respectively. [Figure 25] Figure 25 presents a parametric study of possible performance for cycle 4-3-3.

[0041] The following examples illustrate the invention without limiting its scope.

[0042] Example 1 The installation shown in Figure 1 is designed to treat a saltwater or brine effluent S containing n dissolved salts composed of cations and anions, from which the salts to be extracted are at least extractable by a dedicated organic formulation effluent O circulating in a closed loop.

[0043] The ionic composition of the saltwater or brine S, the ionic composition of the organic formulation O, and the ionic composition of the aqueous phase A vary throughout the process and are sequentially designated S1-S4, O1-O6, and A1 and A2, respectively.

[0044] The desired product is in the form of a hot aqueous phase P1 loaded with at least the salts to be extracted, which is recovered after cooling as production effluent P2.

[0045] The equipment shown in Figure 1 three extraction agitator-mixers 1a, 2a, 3a (each of these extraction agitator-mixers ensures thorough mixing of the aqueous and organic phases fed to them); three decanter-phase separators 1d, 2d and 3d associated with the extraction agitator-mixers 1a, 2a, 3a respectively (these decanter-phase separators, the first of which may be centrifugal and in that case designated 1dc), allow the separation of the organic and aqueous phases of the effluent after it has passed through the respective extraction agitator-mixer 1a, 2a, 3a); two heat exchangers 1e, 2e, each allowing the heating and cooling, respectively, of the organic and aqueous phases fed to them, exchanger 1e being the main exchanger; two differential contactors 1c, 2c, here static or stirred or pulsed columns, each allowing high-temperature desorption or de-extraction of the supplied organic phase of at least the salts to be extracted previously absorbed by the organic phase of the extraction agitator-mixers 1a-3a, heating unit 1b (whose role is described below); Includes:

[0046] The saltwater or brine S1 to be treated is mixed with an organic phase O3 in stirred reactor 1a to produce an organic phase dispersed in a continuous aqueous phase (or vice versa). The two-phase mixture is then transferred to a centrifugal decanter 1dc for separation of the two liquid phases, producing an effluent S2 sent to stirred reactor 2a and an effluent O4 sent to main heat exchanger 1e. This can also allow for the removal of aqueous washout in organic phase O4 (not shown in Figure 1) or the removal of a third phase composed of waste, especially solid waste, before effluent O4 is sent to heat exchanger 1e.

[0047] The salt water or brine S2 is then mixed with the organic phase O2 in stirred reactor 2a to produce an organic phase dispersed in a continuous aqueous phase (or vice versa). The two-phase mixture is then transferred to decanter 2d for gravity separation of the two liquid phases to produce effluent S3 which is sent to stirred reactor 3a, and effluent O3 which is sent to stirred reactor 1a.

[0048] The salt water or brine S3 is then mixed with the organic phase O1 in stirred reactor 3a to produce an organic phase dispersed in a continuous aqueous phase (or vice versa). The two-phase mixture is then transferred to decanter 3d for gravity separation of the two liquid phases to produce an effluent S4 (raffinate) that is treated, and an effluent O2 that is sent to stirred reactor 2a.

[0049] The organic phase O4 loaded with at least the salts to be extracted is then pumped to the main heat exchanger 1e, where it is reheated to obtain a hot effluent O5 loaded with the salts to be extracted, which is then sent to the top of columns 1c and 2c operating in parallel (where the density of O5 is considered to be greater than that of P1) for the production at the bottom of these columns of two organic effluents, these two organic effluents are then combined to form O6, and the hot salt-free organic phase is regenerated, pumped to the main heat exchanger 1e to be cooled and then recovered in the form of organic effluent O1.

[0050] The aqueous phase A1, which is generally fresh water, demineralized or deionized water, is pumped to a secondary heat exchanger 2e where it is heated to obtain a hot effluent A2, which is then sent to the bottom injection of columns 1c and 2c operating in parallel to produce at the top of the two columns two aqueous effluents charged with at least the salts to be extracted, which are then combined to form product P1, and the hot aqueous phase, charged with at least the salts to be extracted, is pumped to a secondary heat exchanger 2e where it is cooled and then recovered as production effluent P2.

[0051] Heating unit 1b compensates for heat losses due to temperature differences in the effluents at the cold ends of heat exchangers 1e and 2e.

[0052] Columns 1c and 2c are preferably used with a successively descending organic phase and an ascending dispersed aqueous phase having a high organic / aqueous (O / A) flow rate ratio during regeneration (hot water desorption) of greater than 5, preferably greater than 10, potentially greater than 15, or even greater than 20. This minimizes freshwater consumption for desorption of extracted salts while increasing reconcentration of extracted salts between the water to be treated and the regenerated or desorbed water produced. Increasing regeneration temperatures promote higher O / A.

[0053] No organic phase droplets are seen moving in the aqueous effluent P1, just as no gaseous organic headspace is seen at the top of columns 1c and 2c, because a system is installed in the aqueous phase that allows any organic phase coalescence associated with an adequate residence time for the organic material to settle. Note that this design is also relevant because there is no organic gas phase at the top of the column, even at high O / A flow ratios during regeneration.

[0054] The aqueous effluents S4, P1 or P2 can be subjected to removal of organic traces by the installation of an oil-water separation unit, which can be a decanter and / or a coalescer combined with an adsorption unit using adsorbents which can be activated carbon, silica gel, diatomite, and / or another similar approach can also be considered that allows the organic components to be recycled in the process and / or transported together with the adsorption medium to a disposal plant by incineration.

[0055] This cycle is called 3-0-Z, where 3 is the number of unit operations to extract the salt to be extracted, 0 is the number of washing operations, and Z is the number of theoretical plates for regeneration-desorption using hot water.

[0056] Example 2 The installation shown in Figure 2 is identical to that shown in Figure 1, except that the decanter-separators 2d and 3d are replaced by centrifugal decanters 2dc and 3dc, and the reheat unit 1b, necessary to compensate for the heat losses at the cold ends of the exchangers 1e and 2e, is placed on the organic effluent O6 instead of the aqueous effluent A2.

[0057] The saltwater or brine S used in this example is difficult to separate by gravity separation and requires centrifugation, and the economics of the process may favor the use of a decanter-centrifuge instead of a gravity decanter-separator.

[0058] Example 3 The installation shown in Figure 3 is designed to treat a saltwater or brine effluent S containing n dissolved salts composed of cations and anions, from which the salts to be extracted are at least extractable by a dedicated organic formulation effluent O circulating in a closed loop.

[0059] The ionic compositions of the saltwater or brine S and the organic compound O vary throughout the process and are sequentially designated S1-S6 and O1-O8, respectively.

[0060] The desired product is in the form of a hot aqueous phase P1 loaded with at least the salts to be extracted, which after cooling is recovered as production effluent P2 and, after reconcentration, is recovered as production effluent P3.

[0061] The equipment shown in Figure 3 is five agitator-mixers 1a, 2a, 3a, 4a, 5a (each of these agitator-mixers ensures thorough mixing of the aqueous and organic phases fed to them, agitator-mixers 1a, 2a, 3a and 4a being extraction agitator-mixers and agitator-mixer 5a being a washing agitator-mixer), five decanter-phase separators 1d, 2d, 3d, 4d and 5dc associated with the agitator-mixers 1a, 2a, 3a, 4a and 5a respectively (these decanter-phase separators (the last of which may be centrifugal and is designated 5dc) allow the separation of the organic and aqueous phases of the effluent after it has passed through the respective agitator-mixer 1a, 2a, 3a, 4a and 5a); two heat exchangers 1e, 2e, each allowing the heating and cooling, respectively, of the organic and aqueous phases fed to them, exchanger 1e being the main exchanger; two differential contactors 1c, 2c, here static or stirred or pulsed columns, each allowing high-temperature desorption or de-extraction of the supplied organic phase of at least the salts to be extracted, previously absorbed by the organic phase of the agitator-mixers 1a to 4a, and a heating unit 1b (the function of which is described below); - two units for reverse osmosis and / or membrane distillation and / or evaporation / condensation and / or temperature absorption of water by renewable organic phases 1f, 2f (allowing separation of the retentate or concentrate or raffinate from the permeate or condensate or de-extraction), Includes:

[0062] The salt water or brine S1 and aqueous retentate / concentrate / raffinate S8 to be treated are mixed with organic phase O4 in stirred reactor 1a to produce an organic phase dispersed in a continuous aqueous phase (or vice versa). The two-phase mixture is then transferred to decanter 1d for separation of the two liquid phases to produce effluent S3, which is sent to stirred reactor 2a, and effluent O5, which is sent to stirred reactor 5a.

[0063] The salt water or brine S3 is then mixed with the organic phase O3 in stirred reactor 2a to produce an organic phase dispersed in a continuous aqueous phase (or vice versa). The two-phase mixture is then transferred to decanter 2d for separation of the two liquid phases to produce effluent S4 which is sent to stirred reactor 3a, and effluent O4 which is sent to stirred reactor 1a.

[0064] The salt water or brine S4 is then mixed with the organic phase O2 in stirred reactor 3a to produce an organic phase dispersed in a continuous aqueous phase (or vice versa). The two-phase mixture is then transferred to decanter 3d for separation of the two liquid phases to produce effluent S5 which is sent to stirred reactor 4a, and effluent O3 which is sent to stirred reactor 2a.

[0065] The salt water or brine S5 is then mixed with the organic phase O1 in stirred reactor 4a to produce an organic phase dispersed in a continuous aqueous phase (or vice versa). The two-phase mixture is then transferred to decanter 4d for separation of the two liquid phases to produce effluent S6 (raffinate) to be treated, and effluent O2 which is sent to stirred reactor 3a.

[0066] The organic phase O5, loaded with at least the salt to be extracted, is then mixed with the aqueous phase A7 in the stirred reactor 5a to produce an aqueous phase dispersed in the continuous organic phase (or vice versa). The two-phase mixture is then transferred to a centrifugal decanter 5dc for separation of the two liquid phases. The resulting effluent S7 is sent to a water recovery unit 1f, such as a reverse osmosis unit, a membrane distillation unit, and / or an evaporation / condensation unit and / or a unit for water absorption by a temperature-regenerable organic phase, while the effluent O6 is sent to the main heat exchanger 1e. The purpose of this step is to wash the organic phase O5 with a very small amount of water stream A7 to remove salt impurities in order to improve the purity of the salt to be extracted contained in the organic effluent O6. The centrifugal decanter 5dc can also be used to remove the aqueous washout of the organic phase O5 while removing the organic washout of the aqueous phase S7, or to remove a third phase consisting of solid waste or other materials, before the effluent O6 is sent to the heat exchanger 1e and the effluent S7 is sent to the water recovery unit 1f.

[0067] The aqueous phase A7, which is generally composed of distilled, fresh or demineralized water and recycled water A6, is used at a small flow rate compared to the organic stream, with a flow rate ratio A / O of the water wash (cold water wash with a neutral pH, which does not affect the wash performance) of less than 0.25, preferably less than 0.1, potentially less than 0.05, or even around 0.01. This minimizes the consumption of fresh water for washing, while increasing the purity of the salt to be extracted from the product P1, while minimizing losses of the salt to be extracted by washing and / or losses in the overall extraction yield of the salt to be extracted.

[0068] The aqueous phase S7, pumped by the temperature-regenerable organic phase 1f to the reverse osmosis unit, and / or membrane distillation unit and / or evaporation / condensation unit and / or water absorption unit, is then separated into brine S8 (retentate or concentrate or raffinate) which is sent to stirred reactor 1a and aqueous effluent A6 (permeate or condensate or de-extractate) which is recycled to the inlet of stirred reactor 5a.

[0069] The organic phase O6 loaded with the salts to be extracted is then pumped to the main heat exchanger 1e, where it is heated to obtain a hot organic effluent O7 loaded with the salts to be extracted, which is then sent to the heads of columns 1c and 2c operating in parallel for the production at the bottom of the columns of two organic effluents, the two organic effluents then being combined to form effluent O8, a hot salt-free organic phase, which is regenerated and pumped to the main heat exchanger 1e to be cooled, and then recovered in the form of a cold organic effluent O1 (here it is assumed that the density of O7 is greater than that of P1).

[0070] The aqueous phase A1, typically fresh water, demineralized water or deionized water, which can be combined with the permeate-condensate-removed extract A2 from the water recovery unit 2f, the reverse osmosis unit, the membrane distillation unit and / or the evaporation / condensation unit and / or the unit for absorbing water by a temperature-regenerable organic phase, is pumped as aqueous effluent A3 to the secondary heat exchanger 2e, where it is heated to obtain a high-temperature aqueous effluent A4, which is then sent to the bottom injection of columns 1c and 2c operating in parallel, producing two aqueous effluents loaded with at least the salts to be extracted at the top of the two columns, which are then combined to form product P1, and the high-temperature aqueous phase loaded with at least the salts to be extracted is pumped to the secondary heat exchanger 2e, where it is cooled and then recovered as production effluent P2, which can be re-concentrated via the water recovery unit 2f, the reverse osmosis unit, the membrane distillation unit and / or the evaporation / condensation unit and / or the unit for absorbing water by a temperature-regenerable organic phase to produce the permeate-condensate-removed extract A2 and the concentrated production effluent P3.

[0071] Heating unit 1b compensates for heat losses due to temperature differences in the effluents at the cold ends of heat exchangers 1e and 2e.

[0072] Columns 1c and 2c are preferably used with a successively descending organic phase and an ascending dispersed aqueous phase having a high O / A flow ratio during regeneration (hot water desorption) of greater than 5, preferably greater than 10, potentially greater than 15 or even greater than 20. This allows for increased reconcentration of the extracted salts between the water to be treated and the regeneration or desorption water produced, while minimizing fresh water consumption for desorption of the extracted salts. The higher the regeneration temperature of the organic phase in 1c and 2c, the higher the O / A ratio.

[0073] At the top of columns 1c and 2c, a system is installed in the aqueous phase that allows for any organic phase coalescence associated with an adequate residence time for the organic material to settle, so that no organic phase droplets are seen to move in the aqueous effluent P1, just as no gaseous organic vapors are seen at the top of columns 1c and 2c. Note that this design is also relevant because there is no organic vapor phase at the top of the column, even at high flow ratios O / A during regeneration.

[0074] The devices 5a and 5dc can be combined to form a single device commonly known as a centrifugal extractor. This device can be doubled, tripled, or quadrupled by series connection, preferably countercurrently, to improve the washing of O5, in order to produce a higher purity product of the salt to be extracted as effluent P1. This device can be multiplexed and used in parallel to be able to process the O5 effluent at increased flow rates.

[0075] The aqueous effluents S6, S7, P1, P2 or P3 can be subjected to removal of organic traces by the installation of an oil-water separation unit, which can be a decanter and / or a coalescer, combined with an adsorption unit using an adsorbent which can be activated carbon, silica gel, diatomaceous earth, and / or another similar approach can be considered that allows the organic components to be recycled to the process and / or transported together with the adsorption medium to a disposal facility by incineration.

[0076] All effluents to the left of heat exchanger 1e and below heat exchanger 2e preferably operate at temperatures close to ambient temperature (5-35°C), below the boiling points of the water and brine effluents A4 and P1 at the operating pressures of 1c, 2c, and 2e, although the other effluents (top right of Figure 3) operate at higher temperatures.

[0077] This alternative configuration can be implemented when the saline or brine aqueous phase to be treated is such that sufficient gravity settling is not economically viable due to its density, viscosity and / or other inherent properties. In this case, instead of the extractor mixer-gravity decanters 1d, 2d, 3d and 4d, mixer-centrifugal decanters 1dc, 2dc, 3dc and 4dc are preferentially used, or potentially centrifugal extractors with increased internal mixing times compared to the current state of the art.

[0078] This cycle is called 4-1-Z, where 4 is the number of unit operations for extracting the salt to be extracted, 1 is the number of washing operations, and Z is the number of theoretical plates for regeneration-desorption of the organic phase using hot water.

[0079] Example 4 When treating salt water or brines where the relative concentration of the salt to be extracted is low compared to other salts present, it may be necessary to use several low temperature washing / scrubbing stages in series to improve washing and removal of impurities from the organic phase.

[0080] The installation shown in Figure 4 is therefore identical to that shown in Figure 3, with the addition of two further agitator-mixers 6a and 7a together with decanter-centrifuges 6dc and 7dc.

[0081] The organic phase O5 loaded with the salt to be extracted is then mixed with brine S10 in a stirred reactor 7a to produce an aqueous phase dispersed in a continuous organic phase (or vice versa). The two-phase mixture is then transferred to a centrifugal decanter 7dc for separation of the two liquid phases, the production of effluent S7 is sent to a water recovery unit 1f such as a reverse osmosis unit, a membrane distillation unit and / or an evaporation / condensation unit and / or a unit for absorption of water by a temperature-regenerable organic phase, and the effluent O6 is sent to an agitator-mixer 6a.

[0082] The aqueous phase S7, pumped by the temperature-regenerable organic phase 1f to the reverse osmosis unit, membrane distillation unit and / or evaporation / condensation unit and / or water absorption unit, is then separated into brine S8 (retentate or concentrate or raffinate) which is sent to stirred reactor 1a and aqueous effluent A6 (permeate or condensate or de-extractate) which is recycled to the inlet of stirred reactor 5a.

[0083] The organic phase O6, loaded with the salt to be extracted, is then mixed in stirred reactor 6a to produce an aqueous phase dispersed in a continuous organic phase (or vice versa). The two-phase mixture is then transferred to centrifugal decanter 6dc to separate the two liquid phases, producing effluent S10, which is sent to stirred reactor 7a, and effluent O7, which is sent to agitator-mixer 5a.

[0084] The organic phase O7 loaded with the salt to be extracted is then mixed in stirred reactor 5a to produce an aqueous phase dispersed in a continuous organic phase (or vice versa). The two-phase mixture is then transferred to centrifugal decanter 5dc to separate the two liquid phases, producing effluent S9 sent to stirred reactor 6a and effluent O8 sent to heat exchanger 1e.

[0085] The addition of one, two or three stirrer-mixers for cold washing of the organic phase improves the purity and selectivity of the salts to be extracted.

[0086] This cycle is called 4-3-Z, where 4 is the number of unit operations for extracting the salt to be extracted, 3 is the number of washing operations, and Z is the number of theoretical plates for regeneration-desorption of the organic phase using hot water.

[0087] The installation of Figure 5 is identical to that of Figure 4, with the water recovery unit 2f moved from the low temperature zone to the high temperature zone of the secondary "water" heat exchanger 2e and the water recovery unit 1f removed.

[0088] The water recovery unit 2f can then be a membrane distillation unit and / or an evaporation / condensation unit and / or a unit for absorbing water using a temperature-regenerable organic phase.

[0089] Example 5 In this example, the salt to be selectively extracted is sodium chloride NaCl.

[0090] The organic phase O is -Na + As an extractant for ions, Na in methanol at a concentration of 0.275 M (mol / L) at 25°C was used. + 4-tert-butyl-calix[4]arenoic acid tetraethyl ester (C 60 H 80 O 12 Sodium ionophore X, also known as CAS number: 97600-39-0 - N-(3,5-dichlorophenyl)-octanamide (C) as an anionic solvating agent with a concentration of 0.825 M and a pKa in water of 13.83 + / - 0.70 14 H 19 Cl2NO, CAS number 20398-46-3), and 1-chloro-2-bromo-benzene as a diluent Includes:

[0091] The organic phase has a density of 1.48 g / L at 20°C.

[0092] The brine to be treated contains sodium chloride and potassium chloride to be extracted. The brine contains 2.26 g / L of Na+ and 107 g / L of K + It has a salinity of 209.81 g / L.

[0093] 0.05, 0.2, 0.5, 1 and 4 aqueous brines (O / A) respectively 抽出 Contacting this brine with the organic phase in the volume ratio used for the above at room temperature (20°C) resulted in a high relative extraction of sodium in the form of NaCl relative to potassium or KCl, as shown in the graph of Figure 6.

[0094] This sodium extraction is selective without changing the potassium concentration, as shown in Figure 7. The potassium concentration is (O / A) 抽出 In tests with ratios of 0.05, 0.2, 0.5, and 1, the residual sodium concentration remained stable at about 107 g / L, (O / A) 抽出 168 mg / L in brine treated with a ratio of 1.

[0095] Figure 8 shows the ion concentrations of various compounds in the organic phase after extraction. NaCl is (O / A) 抽出 If the ratio is less than 2, the organic phase is the majority, (O / A) 抽出 It is even more common when the ratio is low.

[0096] Then, to demonstrate the very poor retention of KCl in the organic phase, an (O / A) of 0.5 抽出 Extraction ratio (A / O) 洗浄 Na, K and Cl loaded formulations were studied after performing water washes at room temperature (20°C) with very small amounts of water ranging from 1% to 20% of the volume of the organic phase, also referred to as ratio.

[0097] The results obtained are shown in Figure 9. As soon as the organic phase comes into contact with 1% by volume of water, 80.8% of the KCl absorbed in the organic phase appears to return to the water, compared to only 1.9% for NaCl. Also, with 4% volume of water, 94.8% of the KCl appears to desorb / transfer to the water in one washing step, compared to 6.4% of NaCl.

[0098] It can also be seen that NaCl is difficult to desorb at room temperature, and when the volume of water is 20% of the organic phase, only 21.1% of the NaCl is transferred to the water.

[0099] High temperature regeneration of the formulation obtained by passing NaCl through the aqueous phase and washing with 4% of the volume of water by contacting with hot water at 80°C in volumes of 2%, 5%, 10%, 20% and 50% of the volume of the organic phase was used to regenerate the organic phase, in this ratio (A / O) 再生 It is shown as follows.

[0100] The results obtained are shown in Figure 10. With 20% water by volume, 67.3% of the NaCl appears to have desorbed from the organic phase at 80°C (compared to 21.1% previously at 20°C).

[0101] Figure 11 shows the ion concentration of the brine obtained after regeneration at 80°C. 4% (A / O) 洗浄 After the washing step with the ratio of these various (A / O) 再生 The brine obtained from the high temperature regeneration of the ratio consists of 93% to 99% NaCl by mass, which allows for the purification of NaCl with minimal NaCl loss.

[0102] Parametric studies were carried out on cycles 4-1-3 (the cycle of Example 3 with three theoretical stages for hot water regeneration-desorption) and 4-2-3 (a cycle with two cold water wash operations and three theoretical stages for hot water regeneration-desorption).

[0103] The results of cycle 4-1-3 are shown in FIG.

[0104] tending to the left of the curve, i.e., towards 0 (A / O) 洗浄 By extending the ratio, the performance of the 4-0-3 cycle without water wash can be obtained, where (A / O) 洗浄 It is clear that this washing, carried out by increasing the ratio, makes it possible to increase the purity of the NaCl produced at the outlet of the regeneration column to a high level of about 94% by mass NaCl (excluding water). Moreover, a sufficiently high (O / A)抽出 In this case, the NaCl extraction yield at 20°C can be maintained above 90% while increasing the purity of the de-extracted NaCl by washing with water at room temperature (here 20°C). For this cycle 4-1-3, the selected sizing point is a 93.2% NaCl extraction yield at a 91.8% NaCl purity by mass, (O / A) 抽出 The ratio is 1.1 and (A / O) 洗浄 The ratio is 5% (circled point).

[0105] Figure 13 shows the effect of adding a second countercurrent scrubbing stage at room temperature. By switching from cycle 4-1-3 to cycle 4-2-3, adding a second scrubbing stage at room temperature makes it possible to further improve the performance of the process, obtaining both a high sodium extraction yield and even higher purity (high levels at 98% by mass NaCl) of the NaCl produced at the outlet (strip) of the regeneration stage.

[0106] In this 4-2-3 cycle, the selected sizing point resulted in a NaCl extraction yield of 92.8% with a NaCl purity of 97.4% by mass, (O / A) 抽出 The ratio is 1.1 (A / O) 洗浄 The ratio is 5% (circled point at the intersection of the two curves).

[0107] Table 1 shows the (O / A) 抽出 Ratio is 1.1, (A / O) 洗浄 Ratio is 5%, (A / O) 再生 The mass balance for cycle 4-1-3 with a ratio of 10% is shown. [Table 1]

[0108] This table shows that raffinate S6 now contains only 0.13 wt. % NaCl to obtain a purity of 99.87 wt. % (excluding water) of KCl and K2SO4, of which K + It is shown that the amount of ions is practically unchanged compared to the amount present in the brine S1 to be treated.

[0109] Table 2 shows (O / A) 抽出 Ratio is 1.1, (A / O) 洗浄 Ratio is 5%, (A / O) 再生 The mass balance for cycle 4-2-3 with a ratio of 10% is shown. [Table 2]

[0110] Comparing with Table 1, it can be seen that the P3 concentrate obtained with two washing steps contains more NaCl (97.34% by mass) than the P3 concentrate obtained with a single washing step (91.81% by mass).

[0111] Example 6 In this example, the salt to be selectively extracted is potassium chloride, KCl.

[0112] The organic phase O is -K + As an extractant for ions, K in methanol was used at a concentration of 0.175 M (mol / L). + 4-tert-butyl-calix[6]arene acid hexaethyl ester (C 90 H 120 O 18 , CAS number: 92003-62-8), - As an anionic solvating agent, N-(3,4-dichlorophenyl)-octanamide (C) with a concentration of 0.525 M and a pKa of 13.63 + / - 0.70 in water 14 H 19 Cl2NO, CAS number: 730-25-6), and 1-chloro-2-bromo-benzene as a diluent Includes:

[0113] The organic phase has a density of 1.43 g / L at 20°C.

[0114] The brine to be treated contains potassium chloride to be extracted and sodium chloride. The brine contains 62.7 g / L of Na + and 4.53 g / L K +It has a salinity of 168g / L.

[0115] The brine was diluted at room temperature (20°C) to 0.1, 0.4, 0.8, 1.5 and 4 aqueous brines (O / A). 抽出 By contacting the organic phase with the volume ratio used for the soluble potassium iodide, a good relative extraction of potassium in the form of KCl relative to sodium or NaCl was obtained, as shown in the graph of FIG.

[0116] This extraction of potassium is selective without changing the sodium concentration, as shown in Figure 15. The sodium concentration is (O / A) 抽出 In tests with ratios of 0.1, 0.4, 0.8, and 1.5, the potassium concentration remained stable at about 62 g / L, (O / A) 抽出 2.8 g / L in brine treated with a ratio of 1.5.

[0117] Figure 16 shows the ion concentrations of various compounds in the organic phase after extraction at 20°C. KCl has an (O / A) of less than 2. 抽出 It accounts for the majority of the organic phase.

[0118] Then, to demonstrate very low retention of NaCl in the organic phase, an (O / A) of 0.8 was used. 抽出 Extraction ratio (A / O) スクラビング Formulations loaded with Na, K, and Cl were studied after scrubbing at room temperature (20°C) with very small amounts of water ranging from 1% to 20% of the volume of the organic phase, also referred to as the ratio.

[0119] The results are shown in Figure 17. As soon as the organic phase comes into contact with 1% water by volume, 48.3% of the NaCl absorbed into the organic phase appears to be returned to water, compared to only 2.1% of the absorbed KCl.

[0120] Also, at 4% water by volume, 76.4% of the NaCl appears to desorb / transfer to water in one wash step compared to 4.9% KCl (90.3% of the NaCl was desorbed at 10% water compared to 8.9% KCl desorbed).

[0121] It can also be seen that KCl is difficult to desorb at room temperature, and when the volume of water is 20% of the organic phase, only 14.8% of the KCl is transferred to the water.

[0122] High temperature regeneration of the formulation obtained by passing KCl through the aqueous phase and washing with 4% of the volume of water by contacting with hot water at 80°C in volumes of 2%, 5%, 10%, 20% and 50% of the volume of the organic phase was used to regenerate the organic phase, in this ratio (A / O) 再生 It is shown as follows.

[0123] The results obtained are shown in Figure 18 and show that with 20% by volume water, 83.7% of the KCl was desorbed from the organic phase (compared to 14.8% previously at 20°C).

[0124] Figure 19 shows the ion concentrations of brine obtained after high temperature regeneration at 80°C. 4% (A / O) 洗浄 After the ratio washing step, the brine obtained from high-temperature regeneration consists of 78% to 90% by mass of KCl, which allows for the purification of KCl for quality improvement and the production of KCl.

[0125] A parametric study was carried out on these cycles 4-1-3 (the cycle of Example 3 with three theoretical stages for hot water regeneration-desorption) and 4-2-3 (a cycle with two cold water wash operations and three theoretical stages for hot water regeneration-desorption).

[0126] The results of cycle 4-1-3 are shown in FIG.

[0127] tending to the left of the curve, i.e., towards 0 (A / O) 洗浄 By extending the ratio, the performance of the 4-0-3 cycle without water wash can be obtained, where (A / O) 洗浄 It is clear that this washing, carried out by increasing the ratio, makes it possible to increase the purity of the KCl produced at the outlet of the regeneration column to a high level of about 91% by mass KCl (excluding water). Moreover, a sufficiently high (O / A) 抽出In this case, the KCl extraction yield can be maintained above 90% while increasing the purity of the de-extracted KCl by washing with water at room temperature (here, 20°C). For this cycle 4-1-3, the selected sizing point is a KCl extraction yield of 85.0% at a KCl purity of 78.5% by mass, (O / A) 抽出 The ratio is 2.2 and (A / O) 洗浄 The ratio is 5% (circled point).

[0128] Figure 21 shows the effect of adding a second countercurrent scrubbing stage at room temperature. By switching from cycle 4-1-3 to cycle 4-2-3, adding a second scrubbing stage at room temperature can further improve the performance of the process, resulting in both a higher potassium extraction yield and even higher purity (as high as 98% by mass of KCl) of the KCl produced at the outlet (strip) of the regeneration stage.

[0129] For this cycle 4-2-3, the sizing point chosen was a KCl extraction yield of 85.6% at a KCl purity of 91.5% by mass, (O / A) 抽出 The ratio is 2.2 and (A / O) 洗浄 The ratio is 5% (circled point).

[0130] Table 3 shows (O / A) 抽出 Ratio is 2.2, (A / O) 洗浄 Ratio is 5%, (A / O) 再生 The mass balance for cycle 4-1-3 with a ratio of 10% is shown. [Table 3]

[0131] This table shows that the raffinate S6 now contains only 0.76% by mass of KCl, of which Na + It is shown that the amount of ions is practically unchanged from that present in the brine S1 to be treated.

[0132] Table 4 shows (O / A) 抽出 Ratio is 2.2, (A / O) 洗浄 Ratio is 5%, (A / O)再生 The mass balance for cycle 4-2-3 with a ratio of 10% is shown. [Table 4]

[0133] Comparison with Table 3 shows that concentrate P3 obtained after two washing steps contains more KCl (91.46% by mass) than concentrate P3 obtained after a single washing step (78.46% by mass).

[0134] Examples 5 and 6 demonstrate that the organic selective extraction formulations and processes used in accordance with the present invention are capable of obtaining relatively pure effluents with NaCl and / or KCl.

[0135] Example 7 In this example, the salt to be selectively extracted is lithium chloride LiCl.

[0136] The organic phase O is -Li-ion + As an extractant for C, N,N,N',N',N",N"-hexacyclohexyl-4,4',4"-propylidinetris(3-oxabutyramide) (C) with a concentration of 0.25 M (mol / L) and a log K of 2.2 in methanol was used. 48 H 83 Lithium ionophore VIII, also known as N3O6, CAS number 133338-85-9 - As an anionic solvating agent, N-(3,5-dichlorophenyl)-octanamide (C) with a concentration of 0.75 M and a pKa of 13.83 + / - 0.70 in water 14 H 19 Cl2NO, CAS number: 20198-46-3), and 1-chloro-2-bromo-benzene as a diluent Includes:

[0137] The organic phase has a density of 1.43 g / L at 20°C.

[0138] The brine to be treated contains lithium chloride. The brine has a Li concentration of 2.26 g / L. + , 2.26g / L Na + and 107 g / L of K + Contains, with a salinity of 209.81 g / L, Li + , Na + , K. + , Mg 2+ , Ca 2+ , Boron, SO4 2- and Cl - It is a brine containing

[0139] Since most lithium-rich brines have a Ca / Li molar ratio below 0.5 (exceptions are Maricunga, Chile: Ca / Li = 2; Tres Quebradas, Argentina: Ca / Li = 7.4), and the Na / Li molar ratio can vary from 1 to 150, it appears that the key to directing lithium extraction is to have very good selectivity between lithium and sodium.

[0140] As an illustration, the average brine composition of the Maricunga Blanco project was examined with representative Na / Li and Ca / Li ratios (22.8 and 2.1, respectively). Table 5 shows the minimum, average, and maximum brine concentrations for the Maricunga Blanco project. [Table 5]

[0141] Laboratory tests combined with ion balance reconstructions for cycles 4-0-3 (for Example 1), 4-1-3 (for Example 3), and 4-2-3, all operated at 20°C for lithium extraction and washing and 80°C for desorption, when lithium extraction and washing are performed, show from Figures 22 and 23 that lithium yields can be as high as 97% and the purity of the LiCl produced is a function of three complementary parameters: (O / A) 抽出 Ratio, cycle selection, and (A / O)洗浄 It shows that it depends on the ratio.

[0142] The purity of the LiCl product (wt %, excluding water) varies from a maximum of 65% in cycle 4-0-3 to 90% in cycle 4-1-3 and 97.5% in cycle 4-2-3.

[0143] As can be seen in these two graphs, (O / A) 抽出 Increasing the ratio increases the lithium yield while decreasing the purity of the LiCl product due to co-extraction of NaCl and CaCl.

[0144] The surprising thing is (A / O) 洗浄 If the ratio remains below 10% (0.100) in cycle 4-1-3 and below 5% (0.050) in cycle 4-2-3, the use of pure water in the wash steps (one or two wash steps in series) results in a very significant improvement in the purity of the LiCl product with very limited loss in lithium yield.

[0145] Figure 23 and Table 6 show the significant benefit of washing with water to improve the purity of the LiCl product. The brine composition is provided at the outlet of the desorption column (P1). [Table 6]

[0146] While lithium is maintained, sodium and calcium salts are removed from the LiCl product. The high purity of the LiCl product and the low extraction rates of NaCl and CaCl allow for the efficient reuse of water through the use of reverse osmosis membrane units, membrane distillation units and / or multi-effect evaporation-condensation units and / or units for water absorption by a temperature-regenerable organic phase and / or a combination of at least two of these units on the output stream of the washing and regeneration-desorption stages, so that the application of this process offers a high level of protection for local water resources, both the water contained in the lithium-rich brine (since the water of this brine is not evaporated for the purpose of reconcentration if the lithium content of the lithium-rich brine to be treated is more than 100 mg / kg of water, preferably more than 150 mg / kg of water) or the fresh water required to operate the process for the washing and regeneration-desorption stages.

[0147] Tables 7 and 8 show detailed mass balances for cycles 4-1-3 and 4-2-3 when operated at a temperature of 20°C for the extraction and wash operations and at a temperature of 80°C in the desorption column for regeneration of the organic phase. In these two cycles, the freshwater reuse is 84% and 87%, respectively, resulting in very low freshwater requirements of 6.03 tons water / ton LiCl and 5.61 tons water / ton LiCl, respectively (and 37.7 and 43.2 tons water / ton LiCl without freshwater reuse). + , Mg 2+ and SO4 2- Note that no LiCl appears in the production of effluent P3, and no boron B is present, 100% of which remains in raffinate B6. Here, the LiCl produced is the same as in effluent P3 when water is excluded, cycle 4-1-3 (O / A) 再生=10 Purity of 85.9% by mass and Cycle 4-2-3 (O / A) 再生=10 The purity is 95.5% by mass.

[0148] O / A ratio 再生Note that decreasing the reclaimed water flow rate by increasing from 10 to 15 produces LiCl of approximately the same purity as produced, but at a cost of about 10 percentage points in lithium extraction yield. On the other hand, the LiCl produced is more concentrated, with an increase of up to 50% (61 g LiCl / kg water compared to 40 g LiCl / kg). [Table 7] [Table 8]

[0149] In addition to Figure 25, which presents a parametric study of possible performance for cycle 4-3-3, Table 9 and Figure 24 provide a summary of the optimal dimensions adopted for selective lithium extraction from Maricunga brine for cycles with 4 extraction stages and 3 theoretical regeneration stages for 0, 1, 2, and 3 wash stages, respectively.

[0150] The selectivity of the chosen organic formulation appears to make it possible to increase the LiCl content in the brine produced at the outlet (strip) of the regeneration process from 2.26 to 51.6% by mass.

[0151] Also, the use of an increasing number of countercurrent washing stages allows for the purification of the produced LiCl to be completed, increasing from 51.6% by mass without washing to 99.1% by mass with three washing stages (scrubbing), while maintaining the overall lithium extraction yield above 90%. [Table 9]

[0152] The total salinity (TDS) appears to decrease with increasing purity of LiCl (up to 99.09% of dry weight), which also allows the osmotic pressure of this effluent to be reduced to 47 bar, making it easier to separate the LiCl from the water for reuse. Compounds of the formula: [ka]

[0153] The compound of interest is synthesized in three successive steps. Step 1: Synthesis of secondary amines [ka]

[0154] Place the ketone (10 mmol, 1 equiv.), solvent (17 vol.), amine (45 mmol, 4.5 equiv.), and reagents in a clean, dry flask. Heat if necessary, depending on the reagents. Follow conversion by TLC, which follows the disappearance of the initial ketone. Evaporate the solvent (and sometimes the residual amine) under reduced pressure. If necessary, filter through Fontainebleau sand, then add methanol (12 vol.). If necessary, add a portion of NaBH4 (30 mmol, 3 equiv.) and stir at room temperature for 1-15 hours.

[0155] Purify on silica gel if necessary (DCM to DCM / ethyl acetate). Yield: 20-77%. Step 2: Synthesis of chloroacetamide [ka]

[0156] Place the previously formed amine (10 mmol, 1 equiv.), dichloromethane (3 volumes, 15 equiv.), and triethylamine (30 mmol, 3 equiv.) in a flask. Add chloroacetyl chloride (20-25 mmol, 2-2.5 equiv.) under argon atmosphere and cool. Shake at room temperature for 5-24 hours. Add 2 volumes of water and back-extract the aqueous phase twice with 2 volumes of DCM. Concentrate the organic phase on a rotary evaporator.

[0157] Purification of the product on a silica gel column (100% DCM eluent). Yield: 30-70%. Step 3: Synthesis of the desired compound [ka]

[0158] In triol, NaH (3.5-4 equiv.) is added to 10 volumes of anhydrous THF. The solvent is heated to reflux under argon. Hot introduction of the triol solubilized in 10 volumes of THF is followed by hot introduction of the synthetic chloroacetamide solubilized in 15-20 volumes of THF. The solvent is stirred under reflux for 1-24 hours. The solvent is neutralized by adding 10 volumes of water. Two back-extractions of the aqueous phase are performed with 5 volumes of DCM. The organic phase is then washed once or twice with 5 volumes of water. The organic phase is concentrated on a rotary evaporator.

[0159] The crude product is purified by silica gel chromatography (eluent: heptane / ethyl acetate). The yield is generally 50-70%.

[0160] Table 10 shows the formulas of the synthesized compounds, with Li VIII renamed CE00. [Table 10-1] [Table 10-2] [Table 10-3] [Table 10-4]

[0161] Tables 11 and 12 show the properties of the synthesized compounds. [Table 11] [Table 12]

Claims

1. 1. A method for selectively extracting a salt to be extracted, which comprises at least one cation and at least one anion, from a saltwater or brine to be treated, which comprises the salt to be extracted and salts other than the salt to be extracted, comprising: The brine to be treated (S1) is sent to a so-called initial extraction agitator-mixer (1a) in which a liquid hydrophobic organic phase is also introduced, said organic phase at least a first hydrophobic protic organic compound that solvates the anion of the salt to be extracted and has a pKa in water at 25°C of at least 9; at least a second hydrophobic organic compound which selectively extracts the cations of the salt to be extracted and which has a complexation constant for the cation to be extracted, the log K value of which in methanol at 25°C is greater than 1, preferably greater than 2; Including, - the brine and the liquid hydrophobic organic phase are mixed in the extraction agitator-mixer (1a), and the mixture is - brine from which salts to be extracted have been removed; an organic phase loaded with the salt to be extracted; into the so-called initial extractor decanter-separator (1d; 1dc) in order to obtain - the brine from which the salt to be extracted has been removed is recovered; The organic phase leaving the initial decanter-separator (1d; 1dc) is sent to a new agitator-mixer, the so-called initial wash agitator-mixer (5a; 7a), where it is mixed with wash water and the mixture is a purified organic phase loaded with the salt to be extracted, which is sent to a so-called "organic" heat exchanger (1e) and heated there; - impurity-loaded wash water sent to the initial extraction agitator-mixer (1a); is sent to a decanter-centrifuge (5dc; 7dc), the so-called initial wash decanter-centrifuge, in order to obtain the organic phase, charged with the salt to be extracted and heated in the so-called "organic" heat exchanger (1e), water containing the salt to be extracted; the regenerated organic phase depleted of the salt to be extracted; into the so-called initial regeneration column (1c; 2c), in which countercurrent exchange is carried out with hot water at a temperature higher than that of the initial brine, in order to obtain The organic phase loaded with the salt to be extracted is heated in the heat exchanger (1e) by the organic phase leaving the regeneration column (1c; 2c), - a process characterized in that said regenerated organic phase is recycled to said initial extraction agitator-mixer (1a).

2. 2. The method according to claim 1, characterized in that before recovering the brine freed of the salts to be extracted, the brine is sent to another extraction agitator-mixer (2a), and subsequently, after decantation, to an extraction decanter-separator (2d) which receives the regenerated organic phase in order to obtain a brine with even less salts to be extracted, which is recovered or sent again to another extraction agitator-mixer (3a) and then to an extraction decanter-separator (3d) in order to obtain a brine with even less salts to be extracted than before, the operation of extracting the salts to be extracted from the brine can be repeated at least once, and the organic phase loaded with the salts to be extracted is returned from a given decanter-separator to a low-level agitator-mixer in order to be recovered at the outlet of the initial decanter-separator, and the organic phase loaded with salts is sent to the so-called "organic" heat exchanger.

3. 3. The method according to claim 1 or 2, characterized in that at least another regeneration column (2c) is provided, connected in parallel to the initial regeneration column (1c) and fed with part of the stream coming from the so-called "organic" heat exchanger (1e).

4. The method according to any one of claims 1 to 3, characterized in that the regeneration column is a stirred column or a pulsed column.

5. 5. The method according to claim 1, wherein the water to serve as regeneration water is sent to a so-called "water" heat exchanger (2e) and heated by the water charged with the salt to be extracted leaving the regeneration column (1c; 2c) or regeneration columns to obtain heated regeneration water which can be further heated before being introduced into the regeneration column (1c; 2c) or into the regeneration columns, if several are provided, and cooling water charged with the salt to be extracted is obtained from the heat exchanger (2e).

6. Before being sent to the initial extraction agitator-mixer (1a), the impurity-laden wash water is - retentate brine or concentrated brine or raffinate brine sent to said initial extraction agitator-mixer (1a); Purified water fed back to the initial wash agitator-mixer (5a; 7a), 6. The method according to claim 1, wherein the sulphuric acid is passed to a purification unit (1f) to obtain

7. 7. The method according to claims 1 to 6, characterized in that before sending the washed organic phase leaving the so-called washing decanter-separator (7dc) to the so-called "organic" heat exchanger (1e), the washed organic phase is sent, after decantation, to another so-called washing agitator-mixer (6a) in order to obtain an organic phase that has undergone a further washing step, which is sent to the so-called "organic" heat exchanger (1e), and subsequently to a so-called washing decanter-centrifuge (6dc) which receives washing water, the operation of washing the organic phase being repeated at least once, and the impurity-laden washing water or the retentate or concentrated brine or raffinate brine being returned to the initial washing agitator-mixer (7a) after purification.

8. The water loaded with salts to be extracted, coming either from the regeneration column (1c; 2c) or from the so-called "water" heat exchanger (2e), - concentrated and purified water charged with the salt to be extracted; - purified water sent either to said regeneration column (1c; 2c) or to said so-called "water" heat exchanger (2e), 8. The method according to claim 1, wherein the sulphur dioxide is passed to a purification unit (2f) to obtain

9. 9. The method according to claim 1, wherein a reverse osmosis unit or a membrane distillation unit or an evaporation / condensation unit or a unit for absorption of water by a temperature-regenerable organic phase or a combination of at least two of these units is used as the purification unit (1f; 2f).

10. - distilled water, -Fresh water, - demineralized water, and - at least a portion of the water charged with the concentrated and purified salt to be extracted from the purification unit (2f), optionally after cooling; 10. The method according to claim 1, wherein at least one of the following is used as the washing water:

11. 11. A method according to any one of claims 1 to 10, characterized in that the initial decanter-separator and / or the washing decanter-separator, if present, are decanter-centrifuges (1dc, 5dc, 6dc, 7dc) and allow the removal of a phase constituted by solid waste, or the aqueous washout of the organic phase, or the separation of a liquid two-phase solvent with an aqueous phase / organic phase ratio of less than 1.0 for the initial decanter-separator and less than 0.2 for the washing decanter-separator.

12. 12. The method according to any one of claims 1 to 11, characterized in that the temperature difference between the extraction stage and the regeneration stage is at least 30°C, preferably more than 40°C, more preferably more than 50°C.

13. The salts to be selectively extracted are NaCl, NaCN, NaNO ３ , sodium salts such as NaBr and NaI, KCl, KCN, KNO ３ , potassium salts such as KBr and KI, LiCl, LiCN, and LiNO ３ 13. The method according to any one of claims 1 to 12, characterized in that the salt is selected from lithium salts such as LiBr, LiI, or other salts that can be selectively extracted in the presence of other salts.

14. 14. The method according to any one of claims 1 to 13, characterized in that the volume ratio between the wash water and the organic phase introduced into the so-called wash agitator-mixer (5a; 6a; 7a) is between 0.01 and 0.1, preferably between 0.03 and 0.

06.

15. 15. The process according to any one of claims 1 to 14, characterized in that the volume ratio between the hot water and the organic phase introduced into the so-called regeneration column (1c; 2c) is between 0.05 and 0.2, preferably between 0.08 and 0.15.

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