Method for selectively extracting a salt to be extracted from saltwater or brine

Abstract

The present invention relates to a method for selectively extracting a salt to be extracted from a brine to be treated containing the salt to be extracted, characterized in that: the brine to be treated (S1) is sent to what is referred to as a first extraction mixing agitator (1a) to which a liquid hydrophobic organic phase is also added; the brine and the liquid hydrophobic organic phase are mixed in the extraction mixing agitator (1a); and the mixture is sent to what is referred to as a first separating decanter (1d, 1dc), to obtain: a brine from which salt to be extracted has been removed; and an organic phase loaded with salt to be extracted; the brine from which the salt to be extracted has been removed is recovered; the organic phase loaded with salt to be extracted is sent to an initial washing mixing agitator / separating decanter centrifuge (5a, 5dc or 7a, 7dc), to obtain a purified organic phase loaded with salt to be extracted which is sent to what is referred to as an "organic" heat exchanger (1e) in which it is heated before being sent to what is referred to as an initial regeneration column (1c; 2c) in which countercurrent exchange with hot water is carried out, this hot water being at a temperature higher than that of the starting brine, to obtain: water containing the salt to be extracted; and a regenerated organic phase, which has lost the salt to be extracted; the regenerated organic phase is recirculated to the initial extraction mixing agitator (1a).

Description

[0001] Description

[0002] Title: Process for the selective extraction of a salt to be extracted from salt water or brine

[0003] The present invention relates to a method for selectively extracting a salt to be extracted from salt water or brine containing the salt to be extracted as well as other salts.

[0004] The recovery of alkali, alkaline earth salts, transition metals or rare earth metals dissolved in natural or leaching brine remains today a relatively complex, costly operation with a high environmental impact and therefore largely improvable.

[0005] It is historically based on a separation of the different salts by crystallization according to the temperature evolution of their respective solubilities coupled with their differences in water solubility which, for example at 50°C follows the increasing order of the following saturation molalities: NaF < NaHCO3~ BaC12 < Na2SO4< MgSO4< ZnSO4< KC1 < MgCl2< NaCl < CdCl2< MnCl2< KNO3< NH4C1 < CaCl2< CsCl < LiCl < ZnCl2< NH4NO3...

[0006] Of course, for each brine to be treated, the number of ions and salts present is generally much lower, which can allow for sufficiently large solubility differences to improve their separation capacity. Indeed, for example, when it comes to separating alkali and alkaline-earth salts, the use of solar evaporation basins or thermal crystallization systems allows for successive precipitation of the different salts according to their initial concentration and their solubility in water. But this generates a high variability in the crystallization process, due to the variability of the mineral composition of the sources, 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 cocrystallization conditions of the salts present.Despite this, the mineral salts resulting from this thermal crystallization are produced pure with low to medium yields or are impure and require one or more additional downstream purification steps. In addition, due to the high energy cost of water evaporation (682 kWh. th / tonne at 20°C), even if modern evaporation processes integrating mechanical vapor compression and / or multiple-effect distillation are implemented industrially, this approach remains very energy-intensive, i.e. in the order of 20 to 60 kWh élect / ton of distilled water or, reduced to crystallized salt, an energy of the order of 100 times higher due to the need, for example, to evaporate 75 to 150 tons of water per ton of crystallized NaCl. Consequently, except for applications with salts of high economic value, solar evaporation ponds are rather used which have the advantage of ensuring a zero Opex in water evaporation energy to the detriment of considerable local water consumption since more than 90% of the water in the brine to be treated is evaporated. This approach, acceptable for water evaporation at the seaside, is much less so when it comes to operations in desert, arid or landlocked locations.Depending on the chemical composition of the water and the metals in solution to be recovered, other separation strategies have been developed, in particular by hydrometallurgy, a process for treating metals by liquid means integrating a step where the metal is solubilized for its purification by the use of an acid (H2SO4) or an oxidant (CI2, H2O2, etc.). This leaching or dissolution is particularly applied for a number of transition metals such as zinc, copper, nickel, cobalt, manganese, for rare earths, uranium and others.

[0007] Once solubilized, some of them can be precipitated as insoluble hydroxide, carbonate or sulfide compounds. This is a precipitation approach. These precipitates depend essentially on the pH of the medium, the solubility products and the redox potentials of the medium. Most frequently, the pH of the solution is increased to form hydroxides which precipitate. The metal is then recovered as a solid by simple decantation, within the limit of its solubility product, which often results in process outlet concentrations that are often higher than environmental discharge standards. Beyond this potential need for additional downstream treatments, this approach can require significant quantities of reagents (neutralizing bases and acids, or Na2CO3„.), or generate operational difficulties due to scaling problems, or require upstream pre-treatment to allow the precipitation of a product of acceptable purity, all of which can make this approach unattractive from an economic point of view while generating significant volumes of waste often stored in tailings retention dams.

[0008] Another approach often implemented is ion exchange. This involves contacting the aqueous solution with a material or an organic formulation capable of exchanging the metal to be extracted, most often with a proton H + or a sodium Na + . A disadvantage of this approach is that if the ion exchange is done with a Na +, the ion exchange resin, whether solid or liquid, undergoes an initial regeneration with acid to desorb the metal ion, then neutralization with sodium hydroxide NaOH before it can be used again in extraction. If the ion exchange is done directly in H + / ionic metal, the water in contact sees its pH drop and it is also necessary to implement reagents in extraction, washing and regeneration of the material (acids and / or neutralization bases) which can be in significant quantities. In both cases, this generates significant operating costs and potentially the local installation of a chlorine and caustic soda plant for the production of these acids and bases. Another disadvantage may be due to the regeneration or acid washing of the ion exchange material which, if it is inorganic, can also undergo leaching and therefore its own degradation over time with potential emissions into the environment of transition and / or heavy metals Mn, Ti, Sn, Cu, Al, V or Sb, depending on the initial composition of the chosen ion exchange material. This case is all the more problematic since the ecotoxicity thresholds for these metals are very low and they are difficult to completely trap.Finally, note the absence of direct extraction of a neutral salt and the systematic addition of a salt / waste to the system and / or its environment such as Na2SC>4 resulting from the combination, overall, of 2 NaOH and H2SO4.

[0009] Another approach is salt adsorption. This involves physically or electrochemically attaching the metal salt to the surface or inside the pores of an adsorbent material. Very few applications have emerged, but include the selective adsorption of lithium and capacitive deionization for the desalination of low-salinity brackish water. Adsorption is naturally economically limited to stopping small quantities of ions due to the very large porous adsorption surfaces required, i.e. less than 5 g NaCl / L of solution in capacitive deionization (GDI) and less than 6.5 g LiCl / L of solution in selective extraction by Adsorption.

[0010] The present invention aims to obtain a technical and economic solution allowing the selective extraction of a salt to be extracted for a wide concentration range, with a high extraction yield, an energy consumption approaching the theoretical energy minimum (associated with the extraction of a salt from water), a capacity to produce this extracted salt in 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 the selective extraction of a salt to be extracted comprising at least one cation and at least one anion from salt water or brine to be treated containing said salt to be extracted and salts other than the salt to be extracted, characterized in that:

[0012] - the brine to be treated is sent to an initial extraction mixer-agitator into which a liquid hydrophobic organic phase is also introduced, said organic phase comprising: o at least one first organic compound solvating the anion of the salt to be extracted, protic and hydrophobic, whose pKa in water at 25°C is at least 9; and o at least one second hydrophobic organic compound selectively extracting the cation of the salt to be extracted, having a complexation constant of the cation to be extracted whose log K value, in methanol at 25°C, is greater than 1, preferably greater than 2,

[0013] - said brine and said liquid hydrophobic organic phase are mixed in said extraction agitator-mixer (1a), and the mixture is sent to an initial extraction decanter-separator, to obtain: o a brine from which the salt to be extracted has been removed; and o an organic phase loaded with salt to be extracted;

[0014] - the brine is recovered from which the salt to be extracted has been removed;

[0015] - the organic phase leaving the initial decanter-separator is sent to a new agitator-mixer called the initial wash agitator in which it is mixed with wash water, and the mixture is sent to a decanter-separator-centrifuge called the initial wash agitator to obtain: o an organic phase loaded with purified salt to be extracted, which is sent to a heat exchanger called "organic" in which it is reheated; and o wash water loaded with impurities which is sent to the initial extraction agitator-mixer;

[0016] - the organic phase loaded with salt to be extracted, heated in the so-called "organic" heat exchanger, is sent to 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 starting brine, to obtain: o water containing the salt to be extracted; and o a regenerated organic phase, having lost the salt to be extracted; said organic phase loaded with salt to be extracted being heated in the heat exchanger by the organic phase leaving the regeneration column,

[0017] - the regenerated organic phase is recycled into the initial extraction mixer-agitator.

[0018] In particular, the extraction of the salt to be extracted can be carried out cold, i.e. at a temperature between -35°C and +45°C. Regeneration in the regeneration column then takes place at a higher temperature.

[0019] In particular, the washing of the organic phase may be carried out at a temperature between 5°C and 45°C. The temperature difference between the extraction and regeneration steps may be at least 30°C, preferably more than 40°C and more preferably more than 50°C. The regeneration steps may be carried out at a temperature of 60°C to 150°C.

[0020] In particular, the organic phase leaving the regeneration column can be reheated before entering the so-called "organic" heat exchanger.

[0021] In particular, the first organic compound solvating the anion of the salt to be extracted may be chosen from N-(3,4-dichlorophenyl)octanamide, N-(3,5-dichlorophenyl)octanamide and N-(3,4,5-trifluorophenyl)octanamide.

[0022] In a particular embodiment, the second organic compound is a lithium cation extracting compound, chosen from the compounds of formula: in which:

[0023] R1 and R2, identical or different, are, whatever their position on a nitrogen atom, independently chosen from linear or branched C1-C12 alkyl, C4-cycloalkyl aryl

[0024] Cg; or - RI and R2, taken together with the nitrogen atom which carries them, form a five-, six-, seven- or eight-membered cycle;

[0025] - R3 is chosen from hydrogen, linear or branched C1-C8 alkyl, C4-C8 cycloalkyl, C2-C6 alkoxyalkyl and alkoxyalkylaryl;

[0026] - R4 is chosen from hydrogen, linear or branched C1-C3 alkyl;

[0027] - R5 is chosen from hydrogen, linear or branched C1-C3 alkyl;

[0028] - R6 is chosen from hydrogen, linear or branched C1-C3 alkyl;

[0029] In a particular embodiment:

[0030] - RI and R2 may be, whatever their position on a nitrogen atom, independently chosen 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,

[0031] - RI and R2, taken together with the nitrogen atom which carries them, can form a pyrrolidine, piperidine, azepane or azocane ring;

[0032] - R3 may be chosen from hydrogen, methyl, ethyl, propyl, iso-propyl, n-butyl, iso-butyl, sec-butyl, tert-butyl, n-pentyl, iso-pentyl, 2-methylbutyl, 2-ethylpropyl, n-hexyl, cyclohexyl, methoxymethyl, methoxyethyl, methoxypropyl, methoxybutyl and -Cib-O-Cib-Phenyl; and

[0033] - R4, R5 and R6 may be hydrogen or methyl, RI and R2 being advantageously chosen from butyl, pentyl, hexyl, heptyl, octyl, nonyl, decyl or phenyl in the case where the brine to be treated has a calcium concentration greater than 10 g / L and / or the selectivity Li + / That 2+ is privileged; and

[0034] RI and R2 being advantageously chosen from iso-propyl, iso-butyl, sec-butyl, tert-butyl, iso-pentyl, 2-methylbutyl, 2-ethylpropyl, cyclobutyl, cyclopentyl, cyclohexyl, cycloheptyl, cyclooctyl or RI and R2, taken together with the nitrogen atom which carries them, form a pyrrolidine, piperidine, azepane or azocane ring in the case where the brine to be treated has a calcium concentration of less than 10 g / L and / or the selectivity Li + / N / A + is privileged.

[0035] The lithium cation extracting compound may be chosen from, but not limited to:

[0036]

[0037] In particular, the second organic compound extracting the cation from the salt to be extracted may be chosen from 4-tert-butyl-Calix

[0004] arene acid tetraethyl ester, 4-tert-butyl-Calix

[0006] arene acid hexaethyl ester, 2- [ 2 , 2 -bis [ [ 2 - ( dicyclohexyl amino ) -2-oxoethoxy ] methyl ] butoxy] - N, N-dicyclohexyl-acetamide and N, N, N ' , N ' , N " , N " - Hexacyclohexyl-4 , 4 ' , 4 " -propylidynetris ( 3-oxabutyramide ) .

[0038] The organic phase may also comprise a hydrophobic polar organic diluent. Preferably, it is an aromatic diluent which may be chosen from polar aromatic compounds such as 1-chloro-2-bromobenzene, 1,2-dibromobenzene, 2-bromotoluene or 3,4-dibromotoluene.

[0039] The first organic compound solvating the anion 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 and the relative molar ratio of said solvating molecule of said complementary anionic species to said extracting molecule of said cationic species is greater than 1. The second organic compound extracting 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.

[0040] Before recovering the brine from which the salt to be extracted has been removed, said brine can be sent to another extraction agitator-mixer followed by an extraction decanter-separator, which receives the regenerated organic phase, to obtain, after decantation, a brine having even less salt to be extracted which is recovered or which is sent again to another extraction agitator-mixer followed by an extraction decanter-separator to obtain a brine having even less salt to be extracted than previously, said operation of extracting salt to be extracted from the brine can still be repeated at least once, and the organic phase loaded with salt to be extracted is returned from a given decanter-separator to the lower level agitator-mixer in order to recover at the outlet of the initial decanter-separator, the organic phase loaded with salt which is sent to the so-called "organic" heat exchanger.

[0041] The plurality of extraction agitator-mixer-decanter-separator can in particular operate countercurrently. The aqueous phase entering an extraction agitator-mixer is brought into contact with the organic phase coming from an upper extraction decanter-separator.

[0042] At least one other regeneration column may be provided, mounted in parallel with said initial regeneration column, which is supplied by part of the flow coming from the so-called heat exchanger.

[0043] "organic".

[0044] In particular, the regeneration column(s) may be stirred columns or pulsed columns.

[0045] The water which will serve as regeneration water can be sent to a heat exchanger called "water" in which it is heated by the water loaded with salt to be extracted leaving the regeneration column(s) if several are provided, to obtain heated regeneration water which can be heated again before being introduced into the regeneration column(s), and from said heat exchanger, cooled water loaded with salt to be extracted is obtained.

[0046] Before being sent to the initial extraction mixer-agitator, the wash water loaded with impurities is sent to a purification unit to obtain: o a retentate brine or concentrate or raffinate which is sent to the initial extraction mixer-agitator; and o purified water which is readdressed to the new initial washing mixer-agitator.

[0047] In particular, before sending the washed organic phase leaving the decanter-separator called the washing decanter to the heat exchanger called the "organic" decanter, said washed organic phase is sent to another agitator-mixer called the washing decanter followed by a decanter-separator-centrifuge called the washing decanter which receives washing water, to obtain, after decantation, an organic phase which has undergone a new washing step which is sent to the heat exchanger called the organic decanter, said 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 concentrate or raffinate, is returned, after purification, to the initial washing agitator-mixer.

[0048] The plurality of agitator-mixer-decanter-washing centrifugal separator can in particular operate countercurrently. The aqueous phase entering a washing agitator-mixer is brought into contact with the organic phase coming from an upper washing centrifugal separator-decanter.

[0049] The salt-laden water to be extracted, which comes either from the regeneration column(s) or from the so-called "water" heat exchanger, can be sent to a purification unit to obtain:

[0050] - water loaded with salt to be extracted, concentrated and purified;

[0051] - purified water which is sent either to the regeneration columns or to the so-called “water” heat exchanger.

[0052] A reverse osmosis unit or membrane distillation unit or an evaporation / condensation unit or a water absorption unit using a temperature-regenerable organic phase or a combination of at least two of these units can be used as a purification unit.

[0053] At least one of the following can be used as wash water:

[0054] - distilled water;

[0055] - fresh water;

[0056] - desalinated water; and

[0057] - at least part of the water loaded with salt to be extracted concentrated and purified coming from the purification unit, possibly after cooling. the initial decanter-separator and / or the washing decanter(s)-separator(s) when present may be centrifugal decanter-separators, allowing the elimination of a phase composed of solid waste or aqueous carryover of the organic phase or the separation of a liquid two-phase medium having 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(s).

[0058] The salt to be selectively extracted may be chosen from sodium salts, such as NaCl, NaCN, NaNOs, NaBr, NaI, potassium salts, such as KCl, KCN, KNO3, KBr, KI, lithium salts, such as LiCl, LiCN, LiNOs, LiBr, LiI, or other salts selectively extractable in the presence of other salts.

[0059] The salts preferentially extracted are so-called di-ionic salts, that is to say a salt comprising a monoatomic cation carrying a single positive charge, such as for example the Li cation + , the cation K + or the Na cation + , as well as a mono- or polyatomic anion carrying a single negative charge, such as for example the monoatomic anion Cl~ or the polyatomic anion NO3- . The volume ratio between the washing water and the organic phase introduced into the so-called washing agitator-mixer(s) may be between 0.01 and 0.1, preferably between 0.03 and 0.06.

[0060] The volume ratio between the hot water and the organic phase introduced into the so-called regeneration column(s) may be between 0.05 and 0.2, preferably between 0.08 and 0.15.

[0061] The following Examples illustrate the present invention without, however, limiting its scope.

[0062] Example 1

[0063] The installation of Figure 1 is intended to treat a salt water or brine effluent S, which contains n dissolved salts, composed of cations and anions, of which the salt to be extracted is at least extractable by an effluent of dedicated organic formulation O circulating in a closed loop.

[0064] The ionic composition of the salt water or brine S, that of the organic formulation 0 and that of the aqueous phase A vary throughout the process, receiving the successive notations respectively SI to S4, 01 to 06 and Al and A2.

[0065] The desired product is in the form of a hot aqueous phase PI loaded with at least the salt to be extracted, which, after cooling, is recovered as production effluent P2. The installation in Figure 1 comprises:

[0066] - three extraction mixers 1a, 2a, 3a; these extraction mixers each ensure intimate mixing of the aqueous and organic phases sent to them;

[0067] - three decanter-phase separators Id, 2d and 3d, associated respectively with the extraction agitator-mixers 1a, 2a, 3a; these decanter-phase separators - the first of which may be of the centrifugal type, then being noted Idc - allow the separation of the organic and aqueous phases of an effluent after passage of the latter through the respective extraction agitator-mixer 1a, 2a, 3a;

[0068] - two heat exchangers 1e, 2e, each allowing the heating and cooling respectively of the organic phase and the aqueous phase addressed to them, the exchanger 1e being the main exchanger;

[0069] - two differential contactors 1c, 2c, which are here static or agitated or pulsed columns, each allowing the hot desorption or de-extraction of the organic phase sent to it of at least the salt to be extracted previously absorbed by the organic phase in the extraction agitators-mixers 1a to 3a;

[0070] - a heating unit 1b, the role of which is indicated below.

[0071] The salt water or brine to be treated SI is mixed in the stirred reactor la with the organic phase O3 to produce an organic phase dispersed in a continuous aqueous phase (or vice versa). The two-phase mixture is then transferred to the centrifugal decanter Idc for the separation of the two liquid phases, and the generation of the effluent S2 sent to the stirred reactor 2a and the effluent O4 sent to the main heat exchanger le. It can also allow the elimination of an aqueous carryover in the organic phase O4 (not shown in Figure 1) or the elimination of a third phase composed of waste, in particular solid waste, before the effluent O4 is sent to the heat exchanger le.

[0072] The salt water or brine S2 is then mixed in the stirred reactor 2a with the organic phase O2 to produce an organic phase dispersed in a continuous aqueous phase (or vice versa). The two-phase mixture is then transferred to the decanter 2d for the gravity separation of the two liquid phases, and the generation of the effluent S3 sent to the stirred reactor 3a and the effluent O3 sent to the stirred reactor 1a.

[0073] The salt water or brine S3 is then mixed in the stirred reactor 3a with the organic phase O1 to produce an organic phase dispersed in a continuous aqueous phase (or vice versa). The two-phase mixture is then transferred to the decanter 3d for the gravity separation of the two liquid phases, and the generation of the treated effluent S4 (the raffinate) and the effluent O2 sent to the stirred reactor 2a.

[0074] The organic phase 04, loaded with at least the salt to be extracted, is then pumped to the main heat exchanger 1c where it is reheated to obtain the hot effluent 05 loaded with the salt to be extracted and then sent to the top (here it is considered that the density of 05 is greater than the density of PI) of the columns 1c and 2c, operating in parallel, for the production at the bottom of these columns of two organic effluents which are then combined for the formation of 06, a hot organic phase without salt, regenerated and pumped to the main heat exchanger 1c to be cooled and then recovered in the form of organic effluent 01.

[0075] The aqueous phase Al, generally fresh water, desalinated or deionized, is pumped to the secondary heat exchanger 2e where it is reheated to obtain the hot effluent A2 which is then sent to the bottom injection of the columns 1c and 2c, operating in parallel, for the production at the top of the two columns of two aqueous effluents, loaded with at least the salt to be extracted, which are then combined for the formation of the production PI, a hot aqueous phase loaded with at least the salt to be extracted and pumped to the secondary heat exchanger 2e to be cooled there and then recovered as production effluent P2.

[0076] Heating unit 1b allows compensation of heat losses generated by the temperature difference of the effluents at the cold end of heat exchangers 1c and 2e.

[0077] Columns 1c and 2c are preferably used with a continuous descending organic phase and an ascending dispersed aqueous phase, with a high organic phase / aqueous phase (O / A) flow rate ratio in regeneration (hot water desorption), greater than 5, preferably greater than 10, potentially greater than 15, or even greater than 20. This makes it possible to minimize the consumption of fresh water for the desorption of the salt to be extracted while increasing the reconcentration of the salt to be extracted between the water to be treated and the regeneration or desorption water to be produced. An increase in the regeneration temperature promotes higher O / A.

[0078] The aqueous effluent PI does not see any transfer of organic phase droplets, just as the top of columns 1c and 2c does not see any organic gaseous overhead due to the installation at the top of these columns, in the aqueous phase, of systems allowing the coalescence of any organic phase associated with an appropriate residence time so that the organic matter can settle. It should be noted that this design is also relevant due to the absence at the top of the column of organic gaseous phase, even if the ratio of 0 / A flows in regeneration is high.

[0079] It may also be considered that the aqueous effluents S4, PI or P2 may be subjected to the elimination of organic traces by the installation of an oil-water separation unit which may be a decanter and / or a coalescer combined with an adsorption unit using adsorbents which may be activated carbon, silica gel, diatomite earth and / or another similar approach allowing the recycling of the organic components to the process and / or their transport, with the adsorption media, to a disposal facility by incineration.

[0080] We denote this cycle 3-0-Z, 3 being the number of unit operations for the extraction of the salt to be extracted, 0 being the number of washing operations and Z being the number of theoretical stages for regeneration-desorption with hot water.

[0081] Example 2

[0082] The installation of Figure 2 is identical to that of Figure 1 except that the decanter-separators 2d and 3d are replaced by centrifugal decanters 2dc and 3dc and that the reheating unit 1b, necessary to compensate for the heat loss at the cold ends of the exchangers 1c and 2e, is placed on the organic effluent 06 in replacement of the aqueous effluent A2.

[0083] The salt water or brine S used in this example is difficult to separate by gravity separation and centrifugal separation is necessary. Also, the economic conditions of the process may be favorable to the implementation of centrifugal separators-decanters instead of gravity separators-decanters.

[0084] Example 3

[0085] The installation of Figure 3 is intended to treat a salt water or brine effluent S, which contains n dissolved salts, composed of cations and anions, of which the salt to be extracted is at least extractable by a dedicated organic formulation effluent 0 circulating in a closed loop.

[0086] The ionic composition of the salt water or brine S and that of the organic formulation 0 vary throughout the process, receiving the successive notations respectively SI to S6, and 01 to 08.

[0087] The desired product is in the form of a hot aqueous phase PI loaded with at least the salt to be extracted, which, after cooling, is recovered as production effluent P2, and after reconcentration, is recovered as production effluent P3.

[0088] The installation in Figure 3 includes:

[0089] - five mixer-agitators 1a, 2a, 3a, 4a, 5a; these mixer-agitators each ensure intimate mixing of the aqueous and organic phases sent to them, the mixer-agitators 1a, 2a, 3a and 4a being extraction mixer-agitators and the mixer-agitator 5a being a washing mixer-agitator;

[0090] - five decanter-phase separators Id, 2d, 3d, 4d and 5dc, associated respectively with the agitator-mixers la, 2a, 3a, 4a and 5a, these decanter-phase separators - the last of which can be of the centrifugal type, being noted 5dc - allow the separation of the organic and aqueous phases of an effluent after passage of the latter through the respective agitator-mixer la, 2a, 3a, 4a and 5a;

[0091] - two heat exchangers 1e, 2e, each allowing the heating and cooling respectively of the organic phase and the aqueous phase addressed to them, exchanger 1e being the main exchanger;

[0092] - two differential contactors 1c, 2c which are here static or agitated or pulsed columns, each allowing the desorption or hot de-extraction of the organic phase sent to it of at least the salt to be extracted previously absorbed by the organic phase in the agitator-mixers 1a to 4a;

[0093] - a heating unit 1b, the role of which is indicated below;

[0094] - two reverse osmosis units and / or membrane distillation and / or evaporation / condensation and / or water absorption by an organic phase regenerable at temperature I f , 2 f allowing the separation of a retentate or concentrate or final raffinate and a permeate or condensate or back-extract.

[0095] The salt water or brine to be treated SI and the aqueous retentate / concentrate / refined S 8 are mixed in the stirred reactor 1a with the organic phase 04 to produce a dispersed organic phase in a continuous aqueous phase (or vice versa). The two-phase mixture is then transferred to the decanter 1d for the separation of the two liquid phases, and the generation of the effluents S3 sent to the stirred reactor 2a and 05 sent to the stirred reactor 5a.

[0096] The salt water or brine S3 is then mixed in the stirred reactor 2a with the organic phase O3 to produce a dispersed organic phase in a continuous aqueous phase (or vice versa). The two-phase mixture is then transferred to the decanter 2d for the separation of the two liquid phases, and the generation of the effluents S4 sent to the stirred reactor 3a and O4 sent to the stirred reactor 1a.

[0097] The salt water or brine S4 is then mixed in the stirred reactor 3a with the organic phase O2 to produce a dispersed organic phase in a continuous aqueous phase (or vice versa). The two-phase mixture is then transferred to the decanter 3d for the separation of the two liquid phases, and the generation of the effluents S5 sent to the stirred reactor 4a and O3 sent to the stirred reactor 2a.

[0098] The salt water or brine S5 is then mixed in the stirred reactor 4a with the organic phase O1 to produce a dispersed organic phase in a continuous aqueous phase (or vice versa). The two-phase mixture is then transferred to the decanter 4d for the separation of the two liquid phases, and the generation of the treated effluents S6 (the raffinate) and O2 sent to the stirred reactor 3a.

[0099] The organic phase 05, loaded with at least salt to be extracted, is then mixed in the stirred reactor 5a with the aqueous phase A7 for the production of an aqueous phase dispersed in a continuous organic phase (or vice versa). The two-phase mixture is then transferred to the centrifugal decanter 5dc for the separation of the two liquid phases, and the generation of the effluents S7 sent to a water recovery unit If such as a reverse osmosis unit, a membrane distillation unit and / or an evaporation / condensation unit and / or a water absorption unit by a temperature-regenerable organic phase, and 06 sent to the main heat exchanger 1e. This step aims to wash the organic phase 05 with a very small flow of water A7 to remove salt impurities in order to improve the purity of the salt to be extracted contained in the organic effluent 06.The centrifugal decanter 5dc can also allow the removal of an aqueous carryover in the organic phase 05 while removing an organic carryover in the aqueous phase S7 or allow the removal of a third phase composed of solid or other waste, before the effluent 06 is sent to the heat exchanger 1e and before the effluent S7 is sent to the water recovery unit If.

[0100] The aqueous phase A7, generally composed of the aqueous phase A5 which is distilled water, fresh water or desalinated water and recycle water A6 is used as a minor flow compared to the organic flow, with a flow rate ratio A / 0 for the water wash (cold water wash, neutral pH, the pH having no impact on the washing performance), less than 0.25, preferably less than 0.1, potentially less than 0.05, or as low as 0.01. This allows both to minimize the consumption of fresh water for washing while increasing the purity of salt to be extracted from the PI production while minimizing the loss of salt to be extracted by washing and / or the loss of overall extraction efficiency of the salt to be extracted.

[0101] The aqueous phase S7 pumped to a reverse osmosis unit and / or a membrane distillation unit and / or an evaporation / condensation unit and / or a water absorption unit by an organic phase regenerable at temperature If is then separated into a brine S8 (retentate or concentrate or raffinate) sent to the stirred reactor 1a and into an aqueous effluent A6 (permeate or condensate or back-extract) recycled to the inlet of the stirred reactor 5a.

[0102] The organic phase 06, loaded with salt to be extracted, is then pumped to the main heat exchanger 1c where it is heated to obtain the hot organic effluent 07 loaded with salt to be extracted and then sent to the top of the columns 1c and 2c, operating in parallel, for the production at the bottom of the columns of two organic effluents which are then combined for the formation of the effluent 08, a hot organic phase without salt, regenerated and pumped to the main heat exchanger 1c to be cooled and then recovered in the form of cold organic effluent 01 (here it is considered that the density of 07 is greater than the density of PI).

[0103] The aqueous phase Al, generally fresh water, desalinated or deionized water, which can be combined with the permeate-condensate-back-extract A2 of a water recovery unit 2f, a reverse osmosis unit, a membrane distillation unit and / or an evaporation / condensation unit and / or a water absorption unit by a temperature-regenerable organic phase, is pumped as aqueous effluent A3 to the secondary heat exchanger 2e where it is reheated to obtain the hot aqueous effluent A4 which is then sent to the bottom injection of the columns 1c and 2c, operating in parallel, for the production at the top of the two columns of two aqueous effluents, loaded with at least salt to be extracted which are then combined for the formation of the production PI,a hot aqueous phase loaded with at least salt to be extracted and pumped to the secondary heat exchanger 2e to be cooled there and then recovered as production effluent P2 which can be reconcentrated via the water recovery unit 2f, a reverse osmosis unit, a membrane distillation unit and / or an evaporation / condensation unit and / or a water absorption unit by a temperature-regenerable organic phase, to produce a permeate-condensate-back-extract flow A2 and a concentrated production effluent P3. The heating unit 1b allows compensation for the heat losses generated by the temperature difference of the effluents at the cold end of the heat exchangers 1c and 2e.,

[0104] Columns 1c and 2c are preferably used with a continuous descending organic phase and an ascending dispersed aqueous phase, this with a high 0 / A flow rate ratio in regeneration (hot water desorption), greater than 5, preferably greater than 10, potentially greater than 15, or even greater than 20. This makes it possible to minimize the consumption of fresh water for the desorption of the salt to be extracted while making it possible to increase the reconcentration of the salt to be extracted between the water to be treated and the regeneration or desorption water to be produced. This 0 / A ratio can be all the higher as the regeneration temperature of the organic phase in 1c and 2c is high.

[0105] The aqueous effluent PI does not see any transfer of organic phase droplets, just as the top of columns 1c and 2c does not see any gaseous organic overhead due to the installation at the top of these columns, in the aqueous phase, of systems allowing the coalescence of any organic phase associated with an appropriate residence time so that the organic matter can settle. It should be noted that this design is also relevant due to the absence at the top of the column of organic gas phase, even if the flow rate ratio 0 / A in regeneration is high. Equipment 5a and 5dc can be combined into a single equipment commonly called a centrifugal extractor. This equipment can be doubled, tripled or quadrupled by placing in series, preferably countercurrent, to improve the washing of the 05 in order to generate as effluent PI, a product of greater purity in salt to be extracted.This equipment can be multiplied and used in parallel in order to be able to treat 05 effluents of increasing flow rates.

[0106] It may also be considered that the aqueous effluents S6, S7, Pl, P2 or P3 may be subjected to the elimination of organic traces by the installation of an oil-water separation unit which may be a decanter and / or a coalescer combined with an adsorption unit using adsorbents which may be activated carbon, silica gel, diatomite earth and / or another similar approach allowing the recycling of the organic components to the process and / or their transport, with the adsorption media, to a disposal facility by incineration.

[0107] All effluents located to the left of heat exchanger 1c and below heat exchanger 2e are preferably operated near ambient temperature (5 to 35°C) when the other effluents (top right of Figure 3) operate at higher temperature but below the boiling point of the water and brine effluents A4 and PI for operating pressures 1c, 2c and 2e.

[0108] A variant of this configuration may be implemented when the saline aqueous phase or brine to be treated, due to its intrinsic properties of density, viscosity and / or others, does not allow sufficient gravity decantation to be economically viable. In this case, instead of the gravity extraction mixer-settlers Id, 2d, 3d and 4d, centrifugal mixer-settlers Idc, 2dc, 3dc and 4d are preferably used or, potentially, centrifugal extractors with increased internal mixing time compared to the current state of the art.

[0109] This cycle is denoted 4-1-Z, 4 being the number of unit operations for the extraction of the salt to be extracted, 1 being the number of washing operations and Z being the number of theoretical stages for the regeneration-desorption of the organic phase with hot water.

[0110] Example 4

[0111] In the case of the treatment of salt water or brine with low relative concentration of the salt to be extracted compared to the other salts present, it may prove necessary to implement several stages of cold washing / scrubbing in series in order to improve the washing and the elimination of impurities from the organic phase.

[0112] The installation of Figure 4 is therefore identical to that of Figure 3 with the presence of two additional agitator-mixers 6a and 7a to which centrifugal decanter-separators 6dc and 7dc are associated.

[0113] The organic phase 05, loaded with the salt to be extracted, is then mixed in the stirred reactor 7a with the brine S10 for the production of an aqueous phase dispersed in a continuous organic phase (or vice versa). The two-phase mixture is then transferred to the centrifugal decanter 7dc for the separation of the two liquid phases, and the generation of the effluents S7 sent to a water recovery unit If such as a reverse osmosis unit, a membrane distillation unit and / or an evaporation / condensation unit and / or a water absorption unit by a temperature-regenerable organic phase, and 06 sent to the agitator-mixer 6a.

[0114] The aqueous phase S7 pumped to a reverse osmosis unit, a membrane distillation and / or an evaporation / condensation unit and / or a water absorption unit by an organic phase regenerable at temperature If is then separated into a brine S8 (retentate or concentrate or raffinate) sent to the stirred reactor 1a and into an aqueous effluent A6 (permeate or condensate or back-extract) recycled to the inlet of the stirred reactor 5a.

[0115] The organic phase 06, loaded with the salt to be extracted, is then mixed in the stirred reactor 6a for the production of an aqueous phase dispersed in a continuous organic phase (or vice versa). The two-phase mixture is then transferred to the centrifugal decanter 6dc for the separation of the two liquid phases, and the generation of the effluents S10 sent to the stirred reactor 7a and 07 sent to the agitator-mixer 5a.

[0116] The organic phase 07, loaded with the salt to be extracted, is then mixed in the stirred reactor 5a for the production of an aqueous phase dispersed in a continuous organic phase (or vice versa). The two-phase mixture is then transferred to the centrifugal decanter 5dc for the separation of the two liquid phases, and the generation of the effluents S9 sent to the stirred reactor 6a and 08 sent to the heat exchanger le.

[0117] The addition of one, two or three mixers for cold washing of the organic phase improves the purity and selectivity of the extraction of the salt to be extracted.

[0118] This cycle is denoted 4-3-Z, 4 being the number of unit operations for the extraction of the salt to be extracted, 3 being the number of washing operations and Z being the number of theoretical stages for the regeneration-desorption of the organic phase with hot water.

[0119] The installation of Figure 5 is identical to that of Figure 4 with the relocation of the water recovery unit 2f from the cold zone to the hot zone of the secondary “water” heat exchanger 2e and the removal of the water recovery unit If.

[0120] The water recovery unit 2f can then be a membrane distillation unit and / or an evaporation / condensation unit and / or a water absorption unit by an organic phase that can be regenerated at temperature.

[0121] Example 5

[0122] In this example the salt to be selectively extracted is sodium chloride NaCl.

[0123] The organic phase 0 includes: - as extractant of the Na ion + , sodium ionophore X, also called 4-tert-butyl-Calix

[0004] arene acid tetraethyl ester (CeoHsoO^, CAS No.: 97600-39-0) at a concentration of 0.275 M (mol / L) and whose log K for the Na ion + in methanol at 25°C is 5;

[0124] - as an anionic solvate N-(3,5-dichlorophenyl)-octanamide (C14H19CI2NO, CAS No.: 20398-46-3) at a concentration of 0.825 M and whose pKa in water is 13.83 + / - 0.70; and

[0125] - l-chloro-2-bromo-benzene as diluent.

[0126] The organic phase has a density of 1.48 g / L at 20°C.

[0127] The brine to be treated includes sodium chloride to be extracted as well as potassium chloride. The salinity of this brine is 209.81 g / L including 2.26 g / L of Na + and 107 g / L of K + .

[0128] By contacting this brine at room temperature (20°C) with volume ratios of the organic phase used relative to the aqueous brine (O / A) extraction of respectively 0.05; 0.2; 0.5; 1 and 4, a high relative extraction of sodium, in the form of NaCl, was obtained relative to potassium, or KC1 according to the graph in Figure 6.

[0129] This sodium extraction is selective without changing the potassium concentration as shown in Figure 7. The potassium concentration remains stable around 107 g / L for the tests at (O / A)extraction ratios of 0.05; 0.2; 0.5 and 1 and a residual sodium concentration of 168 mg / L in the brine treated with an (O / A)extraction ratio of 1.

[0130] Figure 8 shows the ionic concentrations of the different compounds in the organic phase after extraction. NaCl is predominant in the organic phase for an (O / A) extraction ratio less than 2 and even more predominant when the (O / A) extraction ratio is low.

[0131] Then, in order to demonstrate the very low resistance of KC1 within the organic phase, the formulation loaded with Na, K, Cl was considered after extraction at the (O / A) extraction ratio of 0.5 and washing with water was carried out, at room temperature (20°C), with very small quantities of water ranging from 1% to 20% of the volume of the organic phase, also noted ratio (A / O)i avage .

[0132] The results obtained are shown in Figure 9. It appears that upon contact with 1% volume of water, 80.8% of the KC1 absorbed in the organic phase returns to the water compared to only 1.9% of the NaCl. It also appears that with 4% volume of water, in one washing step, 94.8% of the KC1 is desorbed / transferred to the water compared to 6.4% of the NaCl.

[0133] It can also be observed that NaCl is hardly desorbed at room temperature with only 21.1% of NaCl transferred to water when the volume of water is 20% of the organic phase.

[0134] In order to allow this passage of NaCl into the aqueous phase and to regenerate the organic phase, a hot regeneration is used, at 80°C, of ​​the formulation resulting from washing with 4% of the volume of water, by bringing into contact a volume of hot water of 2%, 5%, 10%, 20% and 50% of the volume of the organic phase, this ratio being noted (A / O) regeneration •

[0135] The results obtained are shown in Figure 10. It appears that with 20% by volume of water, 67.3% of the NaCl was desorbed from the organic phase at 80°C (compared to 21.1% previously at 20°C).

[0136] Figure 11 shows the ionic concentrations of the brine obtained after regeneration at 80°C. After a washing step at a (A / O) wash ratio of 4%, the brine resulting from the hot regeneration for these various (A / O) regenerations consists of 93% to 99% mass of NaCl, which allows purification of the NaCl associated with minimal loss of NaCl.

[0137] A parametric study was carried out on cycles 4-1-3 (cycle of Example 3 with 3 theoretical stages for regeneration-desorption in hot water) and 4-2-3 (cycle with two cold water washing operations and 3 theoretical stages for regeneration-desorption in hot water).

[0138] The results of cycle 4-1-3 are shown in Figure 12.

[0139] By the extension of the curves to the left, that is to say at a ratio (A / 0)i aVage tending towards 0, we would have the performance of a cycle without water washing 4-0-3. It appears clearly here that this washing, carried out by increasing the (A / O) washing ratio, makes it possible to increase the purity of the NaCl produced at the outlet of the regeneration columns until reaching a high level of the order of 94% mass of NaCl (water excluded). In addition, for sufficiently high (0 / A) extraction ratios, the NaCl extraction rate at 20°C can be maintained at more than 90% while increasing the purity of the NaCl back-extracted by washing with water at room temperature (20°C here). For this 4-1-3 cycle, the chosen sizing point allows a NaCl extraction yield of 93.2% for a NaCl purity of 91.8% by mass, with an extraction (0 / A) ratio of 1.1 and a washing (A / 0) ratio of 5% (circled points).

[0140] Figure 13 shows the effect of adding a second countercurrent washing stage at room temperature. Adding a second scrubbing stage at room temperature, by switching from the 4-1-3 cycle to the 4-2-3 cycle, allows further improvement in process performance to be achieved, achieving both a high sodium extraction yield and an even higher purity of the NaCl produced (high level at 98% by mass of NaCl) at the outlet of the regeneration stage (strip).

[0141] For this 4-2-3 cycle, the chosen sizing point allows a NaCl extraction yield of 92.8% for a NaCl purity of 97.4% by mass, with an extraction (0 / A) ratio of 1.1 and a washing (A / 0) ratio of 5% (circled points at the intersection of the two curves).

[0142] Table 1 shows the material balance for a 4-1-3 cycle with an extraction (0 / A) ratio of 1.1, a washing (A / 0) ratio of 5% and a regeneration (A / 0) ratio of 10%.

[0143] [Table 1]

[0144] This table shows that the S6 raffinate only contains 0.13% by mass of NaCl to obtain a purity of 99.87% by mass (water excluded) of KC1 and K2SO4, the quantity of K ions + in this one being practically not modified compared to the quantity present in the brine to be treated SI.

[0145] Table 2 shows the material balance for a 4-2-3 cycle with an extraction (O / A) ratio of 1.1, a wash (A / O) ratio of 5% and a regeneration (A / O) ratio of 10%.

[0146]

[0147] Compared to 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).

[0148] Example 6

[0149] In this example the salt to be selectively extracted is potassium chloride KC1.

[0150] Organic phase 0 includes:

[0151] - as an extractant of the K ion + , 4-tert-butyl- Calix

[0006] arene acid hexaethyl ester (C90H120O18, CAS No.: 92003-62-8) at a concentration of 0.175 M (mol / L) and whose log K for the K ion + in methanol is 4.8;

[0152] - as anionic solvate N-(3,4-dichlorophenyl)-octanamide (C14H19CI2NO, CAS No.: 730-25-6) at a concentration of 0.525 M and whose pKa in water is 13.63 + / - 0.70; and

[0153] - l-chloro-2-bromo-benzene as diluent.

[0154] The organic phase has a density of 1.43 g / L at 20°C.

[0155] The brine to be treated includes potassium chloride to be extracted as well as sodium chloride. The salinity of this brine is 168 g / L including 62.7 g / L of Na + and 4.53 g / L of K+ . By contacting at room temperature (20°C) this brine with volume ratios of the organic phase used compared to the aqueous brine (O / A) extraction of 0.1; 0.4; 0.8; 1.5 and 4, a good relative extraction of potassium, in the form of KC1, compared to sodium, or NaCl was obtained according to the graph in Figure 14.

[0156] This potassium extraction is selective without changing the sodium concentration as shown in Figure 15. The sodium concentration remains stable around 62 g / L for the tests with (O / A) extraction ratios of 0.1; 0.4; 0.8 and 1.5 and a potassium concentration of 2.8 g / L in the brine treated with an (O / A) extraction ratio of 1.5.

[0157] Figure 16 shows the ionic concentrations of the different compounds in the organic phase after extraction at 20°C. KC1 is predominant in the organic phase for an extraction ratio (O / A) lower than 2.

[0158] Then, in order to demonstrate the very low resistance of NaCl within the organic phase, the formulation loaded with Na, K, Cl was considered after extraction at the (O / A) extraction ratio of 0.8 and implementation of a water wash (scrub), at room temperature (20°C), with very small quantities of water ranging from 1% to 20% of the volume of the organic phase, also noted ratio (A / O)i avage .

[0159] The results are shown in Figure 17. It appears that upon contact with 1% volume of water, 48.3% of the NaCl absorbed in the organic phase returns to the water against only 2.1% of the absorbed KC1. It also appears that with 4% volume of water, in one washing step, 76.4% of the NaCl is desorbed / transferred to the water against 4.9% of the KC1 (and 90.3% of NaCl desorbed at 10% water for 8.9% of desorbed KC1).

[0160] It can also be observed that KC1 is hardly desorbed at room temperature with only 14.8% of KC1 transferred to water when the volume of water is 20% of the organic phase.

[0161] In order to allow this passage of KC1 into the aqueous phase and to regenerate the organic phase, a hot regeneration is used, at 80°C, of ​​the formulation resulting from washing with 4% of the volume of water, by bringing into contact a volume of hot water of 2%, 5%, 10%, 20% and 50% of the volume of the organic phase, this ratio being noted (A / O) regeneration •

[0162] The results obtained are shown in Figure 18. This time it appears that with 20% water volume, 83.7% of the KC1 was desorbed from the organic phase (compared to 14.8% previously at 20°C).

[0163] Figure 19 shows the ionic concentrations of the brine obtained after hot regeneration, at 80°C. After a washing step at a (A / O) wash ratio of 4%, the brine resulting from hot regeneration consists of 78% to 90% mass of KC1, which allows purification of KC1 associated with production of KC1 for its recovery.

[0164] A parametric study was carried out on these cycles 4-1-3 (cycle of Example 3 with three theoretical stages for regeneration-desorption in hot water) and 4-2-3 (cycle with two washing operations in cold water and 3 theoretical stages for regeneration-desorption in hot water).

[0165] The results of cycle 4-1-3 are shown in Figure 20.

[0166] By extending the curves on the left, i.e. at a washing ratio (A / O) tending towards 0, we would have the performance of a cycle without washing with water 4-0-3. It appears clearly here that this washing, carried out by increasing the washing ratio (A / O), makes it possible to increase the purity of the KC1 produced at the outlet of the regeneration columns until reaching a high level of around 91% by mass of KC1 (water excluded). In addition, for sufficiently high extraction ratios (O / A), the extraction rate in KC1 can be maintained at more than 90% while increasing the purity of the KC1 de-extracted by washing with water at room temperature (20°C here). For this 4-1-3 cycle, the chosen sizing point allows a KC1 extraction yield of 85.0% for a KC1 purity of 78.5% by mass, with an extraction (O / A) ratio of 2.2 and an (A / O) ratio i aV age of 5% (circled points).

[0167] Figure 21 shows the effect of adding a second countercurrent washing stage at room temperature. The addition of a second scrubbing stage at room temperature, by switching from the 4-1-3 cycle to the 4-2-3 cycle, allows further improvement of the process performance to obtain both a high potassium extraction yield and an even higher purity of the KOI produced (high level at 98% by mass of KOI) at the outlet of the regeneration stage (strip). For this 4-2-3 cycle, the design point adopted allows a KC1 extraction yield of 85.6% for a KC1 purity of 91.5% by mass with an extraction ratio (0 / A) of 2.2 and an (A / O)i ratio avage of 5% (circled points)

[0168] Table 3 shows the material balance for a 4-1-3 cycle with an extraction (0 / A) ratio of 2.2, a wash (A / O) ratio of 5% and a regeneration (A / 0) ratio of 10%.

[0169]

[0170] This table shows that the S6 raffinate now only contains 0.76% mass of KC1, the quantity of Na ions + in this one being practically not modified compared to the quantity present in the brine to be treated SI.

[0171] Table 4 shows the material balance for a 4-2-3 cycle with an extraction (O / A) ratio of 2.2, a wash (A / O) ratio of 5% and a regeneration (A / O) ratio of 10%.

[0172]

[0173] Compared to Table 3, it can be seen that the P3 concentrate obtained with two washing steps contains more KC1 (91.46% by mass) than the P3 concentrate obtained with a single washing step (78.46% by mass).

[0174] These examples 5 and 6 show a good capacity of the organic selective extraction formulations and of the process implemented according to the invention, to allow the obtaining of effluents relatively pure in NaCl and / or in KC1.

[0175] Example 7

[0176] In this example the salt to be selectively extracted is lithium chloride LiCl.

[0177] Organic phase 0 includes:

[0178] - as an extractant of the Li ion + , lithium ionophore

[0179] VIII, also called N,N,N',N',N",N"- Hexacyclohexyl-4,4',4"-propylidynetris(3-oxabutyramide) (C48H83N3O6, CAS No. 133338-85-9) at a concentration of 0.25 M (mol / L) and whose log K in methanol is 2.2;

[0180] - as an anionic solvate N-(3,5-dichlorophenyl)-octanamide (C14H19CI2NO, CAS No.: 20198-46-3) at a concentration of 0.75 M and whose pKa in water is 13.83 + / - 0.70; and

[0181] - l-chloro-2-bromo-benzene as diluent.

[0182] The organic phase has a density of 1.43 g / L at 20°C. The brine to be treated contains lithium chloride. This brine is a Li brine + , N / A + , K + , Mg 2+ , That 2+ , Boron, SO4 2 - and Cl~ of 209.81 g / L of salinity, including 2.26 g / L of Li + , 2.26 g / L of Na + and 107 g / L of K + .

[0183] 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 seems that the key to direct lithium extraction is to have very good selectivity between lithium and sodium.

[0184] For illustration, the average brine composition of the Maricunga Blanco project was examined, which has representative Na / Li and Ca / Li ratios (22.8 and 2.1, respectively). Table 5 shows the minimum, average and maximum brine concentrations of the Maricunga Blanco project.

[0185] [Table 5]

[0186] Thanks to laboratory tests associated with a work of reconstruction of the ionic balances for the cycles 4-0-3 (case of Example 1), 4-1-3 (case of Example 3) and 4-2-3, all operating at 20°C for the extraction of Lithium and the washing when it is carried out, and at 80°C for the desorption, it appears from figures 22 and 23 that the lithium yield can reach 97% and that the purity of the LiCl produced depends on three complementary parameters, the ratio (O / A) extraction, the choice of the cycle and the ratio (A / 0)i avag e.

[0187] The purity of the LiCl product (in % by mass, excluding water) varies from a maximum of 65% for the 4-0-3 cycle, to 90% for the 4-1-3 cycle and to 97.5% for the 4-2-3 cycle.

[0188] As can be seen in these two graphs, an increase in the (O / A)extraction ratio allows an increase in lithium yield while reducing the purity of the LiCl product by co-extraction of NaCl and CaCl2.

[0189] The surprising point is that when pure water is used for the washing steps (for one or two washing steps in series), the purity of the LiCl product is very considerably improved with a very limited loss on the lithium yield when the (A / O) washing ratio remains below 10% (0.100) for the 4-1-3 cycle and below 5% (0.050) for the 4-2-3 cycle.

[0190] Figure 23 and Table 6 show the great advantage of water washing to gain purity in LiCl production. The brine compositions are given at the outlet of the desorption column (PI).

[0191] [Table. 6]

[0192]

[0193] Sodium and calcium salts are eliminated from LiCl production when lithium is maintained.The application of this process allows a high protection of local water resources, whether it is the water contained in the lithium-rich brine (because the water in this brine is not evaporated for reconcentration purposes if the lithium content in the lithium-rich brine to be treated is higher than 100 mg / kg of water, preferably higher than 150 mg / kg of water) or the fresh water required for the operation of the process for the washing and regeneration-desorption steps as a high purity of the LiCl product and a low extraction rate of NaCl and CaCl2 allow an efficient recycling of the water via the use of a reverse osmosis membrane unit, a membrane distillation unit and / or a multi-effect evaporation-condensation unit and / or a water absorption unit by a temperature-regenerable organic phase and / or a combination of at least two of these units on the outlet stream(s) of the washing steps and regeneration-desorption.

[0194] Tables 7 and 8 give a detailed mass balance for cycles 4-1-3 and 4-2-3 when operating at temperatures of 20°C for the extraction and washing operations and 80°C in the desorption columns for the regeneration of the organic phase. For these two cycles, the freshwater recycling is 84% ​​and 87%, respectively, generating very low freshwater requirements of 6.03 tonnes of water / tonne LiCl and 5.61 tonnes of water / tonne LiCl, respectively (and 37.7 and 43.2 tonnes of water / tonne LiCl without freshwater recycling). It should be noted that K + , Mg 2+ and SO4 2 - do not appear in the production of the P3 effluent and that boron B is not presented as 100% of it remains in the B6 raffinate. Here, the LiCl produced is 85.9% pure by mass for the 4-1-3 (O / A) cycle r agenerous r has tion= 10 and at 95.5% by mass for the cycle 4- 2-3 (O / A) regeneration=l 0 in the effluent P3 when the water is excluded. Note that a reduction in the regeneration water flow rate by a change from the 0 / A ratio ré g énération from 10 to 15 generates approximately the same purity of LiCl produced but degrades the lithium extraction yield by about ten percentage points. On the other hand, the LiCl produced is more concentrated with a gain of up to 50% (61 g LiCl / kg water compared to 40 g LiCl / kg).

[0195]

[0196]

[0197] In addition to Figure 25 presenting the parametric study of the possible performances for the 4-3-3 cycle, Table 9 and Figure 24 provide a summary of the optimum dimensions retained for the case of selective lithium extraction for the Maricunga brine, for the cycles with four extraction stages and three theoretical regeneration stages for respectively 0, 1, 2 and 3 washing stages.

[0198] It appears that the selectivity of the chosen organic formulation makes it possible to go from 2.26% by mass of LiCl in the brine to be treated to 51.6% by mass of LiCl in the brine produced at the regeneration outlet (strip).

[0199] It also appears that the implementation of an increasing number of countercurrent washing stages makes it possible to complete the purification of the LiCl produced, going from 51.6% by mass without washing to 99.1% with three washing stages (scrub) to maintain the overall lithium extraction yield at more than 90%.

[0200] [Table 9]

[0201] It appears that the total salinity (TDS) decreases with the increase in the purity of LiCl (up to 99.09% in dry mass) which also makes it possible to lower the osmotic pressure of this effluent to 47 bars for greater ease of separation of LiCl from water, for its recycling.

[0202] The synthesis of the compounds of interest is carried out in three successive stages.

[0203] Step 1: Synthesis of the secondary amine

[0204] In a clean, dry flask, introduce the ketone (10 mmol, 1 eq), the solvent (17 Vol), the amine (45 mmol, 4.5 eq) then the reagent(s). Depending on the reagent, heat if necessary. The conversion is monitored by TLC with the disappearance of the starting ketone. Evaporation of the solvent (and sometimes the residual amine) under reduced pressure. Filtration if necessary on Fontainebleau sand then addition of methanol (12 Vol). Addition in portions of NaBH4 (30 mmol, 3 eq) if necessary, stirring for 1-15 h at RT.

[0205] Purification on silica gel if necessary. (DCM at

[0206] DCM / ethyl acetate). Yield: 20-77% Step 2: Synthesis of chloroacetamide

[0207] In a flask, introduce the previously formed amine (10 mmol, 1 eq), dichloromethane (3 Vol, 15 eq), triethylamine (30 mmol, 3 eq). Add cold and under an argon atmosphere chloroacetyl chloride (20-25 mmol, 2-2.5 eq). Stir at room temperature for 5-24 h. Add 2 volumes of water and carry out two counter-extractions of the aqueous phase with 2 volumes of DCM. Concentrate the organic phase on a rotary evaporator.

[0208] Purification of the product on a silica gel column (eluent DCM 100%). Yield: 30-70%

[0209] Step 3: Synthesis of the compound of interest

[0210] In a three-necked flask, introduce NaH (3.5-4 eq) into 10 volumes of anhydrous THF. Heat the medium to reflux under argon. Hot introduction of the triol solubilized in 10 volumes of THF then the synthesized chloroacetamide, solubilized in 15-20 volumes of THF. Stir the medium at reflux for 1-24 hours. Neutralize the medium by adding 10 volumes of water. Perform two counter-extractions of the aqueous phase with 5 volumes of DCM. Then perform one or two washes of the organic phase with 5 volumes of water. Concentrate the organic phase on a rotary evaporator.

[0211] 5 Purify the crude product obtained by chromatography on silica gel (eluent: heptane / ethyl acetate). Yields are generally between 50-70%.

[0212] Table 10 presents the formulas of the synthesized 0 compounds, with Li VIII being renamed CE00.

[0213] [Table. 10]

[0214]

[0215] Tables 11 and 12 present the characteristics of the synthesized compounds. [Table. 11]

[0216] [Table 12] (Me = Methyl)

[0217]

Claims

REDEMPTION CAT IONS

1. - Process for the selective extraction of a salt to be extracted comprising at least one cation and at least one anion from salt water or brine to be treated containing said salt to be extracted and salts other than the salt to be extracted, characterized in that: - the brine to be treated (SI) is sent to an extraction mixer-agitator (la) called initial into which a liquid hydrophobic organic phase is also introduced, said organic phase comprising: o at least one first organic compound solvating the anion of the salt to be extracted, protic and hydrophobic, whose pKa in water at 25°C is at least 9; and o at least one second hydrophobic organic compound selectively extracting the cation of the salt to be extracted, having a complexation constant of the cation to be extracted whose log K value, in methanol at 25°C, is greater than 1, preferably greater than 2, - said brine and said liquid hydrophobic organic phase are mixed in said extraction agitator-mixer (la), and the mixture is sent to an extraction decanter-separator (Id; Idc) called initial, to obtain: o a brine from which the salt to be extracted has been removed; and o an organic phase loaded with salt to be extracted; - the brine is recovered from which the salt to be extracted has been removed; - we address the organic phase coming out of the decanter- initial separator (Id; Idc) to a new agitator-mixer called initial washing (5a; 7a) in which it is mixed with washing water, and the mixture is sent to a centrifugal decanter-separator (5dc; 7dc) called initial washing to obtain: o an organic phase loaded with purified salt to be extracted, which is sent to a heat exchanger called "organic" (le) in which it is heated; and o washing water loaded with impurities which is sent to the initial extraction agitator-mixer (la); - the organic phase loaded with salt to be extracted, heated in the so-called "organic" heat exchanger (le), is sent to a column (le; 2c) called the initial regeneration column in which a countercurrent exchange is carried out with hot water, at a temperature higher than that of the starting brine, to obtain: o water containing the salt to be extracted; and o a regenerated organic phase, having lost the salt to be extracted; said organic phase loaded with salt to be extracted being heated in the heat exchanger (le) by the organic phase leaving the regeneration column (le; 2c), - the regenerated organic phase is recycled into the initial extraction agitator-mixer (la).

2. - Method according to claim 1, characterized in that before recovering the brine from which the salt to be extracted has been removed, said brine to another extraction agitator-mixer (2a) followed by an extraction decanter-separator (2d), which receives the regenerated organic phase, to obtain, after decantation, a brine having even less salt to be extracted which is recovered or which is sent to another extraction agitator-mixer (3a) followed by an extraction decanter-separator (3d) to obtain a brine having even less salt to be extracted than previously, said operation of extracting salt to be extracted from the brine being able to be repeated at least once, and the organic phase loaded with salt to be extracted is returned from a given decanter-separator to the lower level agitator-mixer in order to recover at the outlet of the initial decanter-separator, the organic phase loaded with salt which is sent to the heat exchanger called “organic”.

3. - Method according to one of claims 1 and 2, characterized in that at least one other regeneration column (2c) is provided, mounted in parallel with said initial regeneration column (le), which is supplied by part of the flow coming from the so-called "organic" heat exchanger (le).

4. - Method according to one of claims 1 to 3, characterized in that the regeneration column(s) (le; 2c) are stirred columns or pulsed columns.

5. - Method according to one of claims 1 to 4, characterized in that the water which will serve as regeneration water is sent to a heat exchanger called "water" (2e) in which it is heated by the water loaded with salt to be extracted leaving the regeneration column(s) (le; 2c) if several are provided, to obtain heated regeneration water which can be reheated again before being introduced into the regeneration column(s) (le; 2c), and from said heat exchanger (2e), cooled water loaded with salt to be extracted is obtained.

6. Method according to one of claims 1 to 5, characterized in that before being sent to the initial extraction agitator-mixer (la), the washing water loaded with impurities is sent to a purification unit (If) to obtain: o a retentate brine or concentrate or raffinate which is sent to the initial extraction agitator-mixer (la); and o purified water which is readdressed to the new initial washing agitator-mixer (5a; 7a).

7. - Method according to claims 1 to 6, characterized in that before sending the washed organic phase leaving the decanter-separator called washing (7dc) to the heat exchanger called "organic" (le), said washed organic phase is sent to another agitator-mixer (6a) called washing followed by a decanter-separator-centrifuge called washing (6dc) which receives washing water, to obtain, after decantation, an organic phase which has undergone a new washing step which is sent to the heat exchanger called organic (le), said operation of washing the organic phase being able to be repeated at least once, and the washing water loaded with impurities, or the brine retentate or concentrate or raffinate, is returned, after purification, to the initial wash agitator-mixer (7a).

8. - Method according to one of claims 1 to 7, characterized in that the water loaded with salt to be extracted, which comes either from the regeneration column(s) (le; 2c) or from the so-called "water" heat exchanger (2e), is sent to a purification unit (2f) to obtain: - water loaded with salt to be extracted, concentrated and purified; - purified water which is sent either to the regeneration columns (2c), or to the heat exchanger called “water” (2e).

9. - Method according to one of claims 1 to 8, characterized in that the purification unit (If; 2f) is a reverse osmosis unit or membrane distillation unit or an evaporation / condensation unit or a unit for absorption of water by an organic phase regenerable at temperature or a combination of at least two of these units.

10. - Method according to one of claims 1 to 9, characterized in that at least one of the following is used as washing water: - distilled water; - fresh water; - desalinated water; and - at least part of the water loaded with salt to be extracted concentrated and purified coming from the purification unit (2f), possibly after cooling.

11. - Method according to one of claims 1 to 10, characterized in that the initial decanter-separator and / or the washing decanter(s)-separator(s) when they are present are centrifugal decanter-separators (Idc, 5dc, 6dc, 7dc), allowing the elimination of a phase composed of solid waste or aqueous carryover of the organic phase or the separation of a liquid two-phase medium having 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(s).

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

13. - Method according to one of claims 1 to 12, characterized in that the salt to be selectively extracted is chosen from sodium salts, such as NaCl, NaCN, NaNOs, NaBr, NaI, potassium salts, such as KCl, KCN, KNO3, KBr, KI, lithium salts, such as LiCl, LiCN, LiNOs, LiBr, LiI, or other salts selectively extractable in the presence of other salts.

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

06.

15. - Method according to one of the claims 1 to 14, characterized in that the volume ratio between the hot water and the organic phase introduced into the column(s) (le; 2c) known as the regeneration column is between 0.05 and 0.2, preferably between 0.08 and 0.15.

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