Selective extraction process for a salt to be extracted from a brine

Abstract

Process for selectively extracting a salt from a brine The present invention relates to a process for selectively extracting a salt from a brine containing said salt, characterized in that: the brine (S1) is sent to an extraction mixer-agitator (1a) into which a liquid hydrophobic organic phase is also introduced; said brine and said liquid hydrophobic organic phase are mixed in said extraction mixer-agitator (1a), and the mixture is sent to a decanter-separator (1d), to obtain: a brine from which salt has been removed; and an organic phase loaded with salt; the brine from which the salt has been removed is recovered; the organic phase loaded with salt is sent to an "organic" heat exchanger (1e) in which it is heated before being sent to a column (1c;2c) said regeneration in which a counter-current exchange is carried out with hot water, which is at a temperature higher than that of the starting brine, to obtain: water containing the salt to be extracted; and a regenerated organic phase, having lost the salt to be extracted; the regenerated organic phase is recycled in the extraction agitator-mixer (1a).

Description

Title of the invention: Method for the selective extraction of a salt from a brine

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

[0002] The valorization of alkali, alkaline-earth salts, transition metals or rare earth metals solubilized in a natural brine or leaching remains today a relatively complex, costly operation with a high environmental impact and therefore largely improvable.

[0003] Historically, it is based on separating the different salts by crystallization according to the temperature evolution of their respective solubilities coupled with their differences in solubility in water, which, for example at 50°C, follows the increasing order of the following saturation molalities: NaF < NaHCO3 ~ BaCl2 < Na2SO4 < MgSO4 < ZnSO4 < KCl < MgCl2 < NaCl < CdCl2 < MnCl2 < KN03 < NH4Cl < CaCl2 < CsCl < LiCl < ZnCl2 < NH4NO3...

[0004] 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 differences in solubility to improve their separation capacity. Indeed, for example, when it comes to separating alkali and alkaline earth salts, the use of solar evaporation ponds or thermal crystallization systems allows for the successive precipitation of the different salts according to their initial concentration and their solubility in water. However, this generates significant variability in the crystallization process, due to the variability in the mineral composition of the sources, their variability over time, and the degree of concentration of the brine to be treated. This approach requires a very good understanding of the crystallization and co-crystallization conditions of the salts present.Despite this, the mineral salts produced by this thermal crystallization are either pure with low to medium yields or impure and require one or more additional downstream purification steps. Furthermore, due to the high energy cost of water evaporation (682 kWh th / tonne at 20°C), even with the industrial implementation of modern evaporation processes incorporating mechanical vapor compression and / or multi-effect distillation, this approach remains very energy-intensive, requiring on the order of 20 to 60 kWh / tonne of distilled water, or, when applied to crystallized salt, approximately 100 times more energy due to the need, for example, to evaporate 75 to 150 tonnes of water per tonne of crystallized NaCl. Consequently, except for applications involving salts of high economic value, solar ponds are generally used. Evaporation methods have the advantage of ensuring zero operating costs in terms of water evaporation energy, but at the cost of considerable local water consumption, since more than 90% of the water in the brine being treated is evaporated. This approach, acceptable for water evaporation at the seaside, is much less so for operations in desert, arid, or landlocked locations.

[0005] Depending on the chemical composition of the water and the metals in solution to be recovered, other separation strategies have been developed, particularly hydrometallurgy, a liquid metal treatment process incorporating a step where the metal is solubilized for purification using an acid (H2SO4) or an oxidant (Cl2). This leaching or dissolution is particularly applied to many transition metals such as zinc, copper, nickel, cobalt, manganese, rare earth elements, uranium, and others.

[0006] Once solubilized, some of these compounds can precipitate as insoluble hydroxide, carbonate, or sulfide compounds. This is a precipitation approach. These precipitates depend primarily 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 then precipitate. The metal is then recovered as a solid by simple decantation, up to the limit of its solubility product, which often results in outlet concentrations exceeding environmental discharge standards. Beyond this potential need for further downstream treatment, 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 pretreatment to allow the precipitation of a product of acceptable purity, all of which can make this approach unattractive from an economic point of view.

[0007] 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 ion (Na+). A drawback of this approach is that if the ion exchange is with Na+, the extracted metal will be impure during regeneration, and if the ion exchange is with H+, it is again necessary to use reagents for extraction, washing, and regeneration of the material (neutralizing acids and bases), which can be required in significant quantities. This generates considerable operating costs and potentially necessitates the local installation of a chlorine and caustic soda plant for the production of these acids and bases.Another drawback may be due to the regeneration or acid washing of the ion exchange material which, if inorganic, may 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. Finally, note the absence of direct extraction of a neutral salt and the systematic addition of a salt / waste product to the system, such as Na2SO4 resulting from the combination of two NaOH molecules and H2SO4.

[0008] Another approach is salt adsorption. This involves physically or electrochemically attaching the metal salt to the surface or within the pores of an adsorbent material. Very few applications have emerged, but examples include selective lithium adsorption and capacitive deionization for the desalination of low-salinity brackish water. Adsorption is naturally economically limited to the removal of small quantities of ions due to the very large porous adsorption surfaces required: less than 5 g NaCl / L of solution in capacitive deionization (CDI) and less than 6.5 g LiCl / L of solution in selective extraction by adsorption.

[0009] The present invention aims to obtain a technical and economic solution enabling the selective extraction of a salt to be extracted for a wide range of concentration, with a high extraction yield, an energy consumption approaching the theoretical minimum (associated with the extraction of a salt from water), an ability to produce this extracted salt in high purity, a water consumption also close to the theoretical minimum and zero reagent consumption.

[0010] To this end, the present invention relates to a process for the selective extraction of a salt to be extracted comprising at least one cation and at least one anion from a 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 is sent to an extraction agitator-mixer into which a liquid hydrophobic organic phase is also introduced, said organic phase comprising: • at least one first organic compound that solvates the anion of the salt to be extracted, is protic and hydrophobic, and has a pKa in water at 25°C of at least 9; and • at least one second hydrophobic organic compound that selectively extracts the cation from 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 2, • The brine and the liquid hydrophobic organic phase are mixed in the extraction agitator-mixer, and the mixture is sent to an extraction decanter-separator 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 is recovered from which the salt to be extracted has been removed; • The organic phase, loaded with salt to be extracted, is sent to an "organic" heat exchanger where it is heated before being sent to a regeneration column where a counter-current exchange takes place with hot water, which is at a temperature higher than that of the starting brine, to obtain: • water containing the salt to be extracted; and • a regenerated organic phase, having lost the salt to be extracted;

[0011] said organic phase loaded with salt to be extracted being heated in the heat exchanger by the organic phase exiting the regeneration column, • the regenerated organic phase is recycled in the extraction agitator-mixer.

[0012] In particular, the extraction of the salt to be extracted can be carried out cold, i.e. at a temperature between -35°C and 35°C. Regeneration in the regeneration column then takes place at a higher temperature.

[0013] The temperature difference between the extraction and regeneration stages is at least 30°C, preferably more than 40°C and even more preferably more than 50°C.

[0014] In particular, the organic phase exiting the regeneration column can be heated before entering the so-called "organic" heat exchanger.

[0015] In particular, the first organic compound solvating the anion of the salt to be extracted can be chosen from N-(3,4-dichlorophenyl)octanamide and N-(3,5-dichlorophenyl)octanamide.

[0016] In particular, the second organic compound extracting the cation from the salt to be extracted can be chosen from 4-tert-butyl-Calix[4]arene acid tetraethyl ester and 4-tert-butyl-Calix[6]arene acid hexaethyl ester.

[0017] The organic phase may also include a fluidizer selected from polar aromatic compounds such as l-chloro-2-bromo-benzene.

[0018] Before recovering the brine from which salt to be extracted has been removed, said brine is 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 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 before, said operation of extracting salt from the brine being able to be repeated at least once more, and the organic phase loaded with salt to be extracted, coming from each of the extraction decanter-separators, is returned to the initial extraction agitator-mixer, the regenerated organic phase arriving in each case in the last extraction agitator-mixer.

[0019] The plurality of agitator-mixer-decanter-extraction separator can be in part The particular process operates in a counter-current manner. The aqueous phase entering an extraction mixer agitator is brought into contact with the organic phase coming from an upper extraction decanter-separator.

[0020] At least one other regeneration column mounted in parallel with the said initial regeneration column can be provided, which is supplied by a part of the flow from the so-called "organic" heat exchanger.

[0021] The water that will serve as regeneration water can be directed into a heat exchanger called "water" in which it is heated by the water loaded with salt to be extracted coming out of 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.

[0022] The organic phase exiting the initial separator-decanter can be directed to a new agitator-mixer, called a washing agitator, in which it is mixed with wash water, and the mixture is directed to a decanter-separator, called a washing clarifier, to obtain: • a purified organic phase loaded with salt to be extracted, which is returned to the so-called "organic" heat exchanger; and • an aqueous phase that is sent to a purification unit to obtain: • a brine that is re-tempered, concentrated, or refined, which is then sent to the initial extraction mixer-agitator; and • water which is then sent back to the new washing agitator-mixer.

[0023] In particular, before sending the washed organic phase from the washing decanter-separator to the "organic" heat exchanger, said washed organic phase is sent to another washing agitator-mixer followed by a washing decanter-separator which receives wash water, to obtain, after decantation, an organic phase which has undergone a further washing step which is sent to the organic heat exchanger, said washing operation of the organic phase being able to be repeated at least once more, and the final wash water, or the brine, re-tempered or concentrated or refined, after purification, is sent back to the initial washing agitator-mixer.

[0024] The plurality of centrifugal washing agitator-mixer-settlers can, in particular, operate in counter-current flow. The aqueous phase entering a washing agitator-mixer is brought into contact with the organic phase coming from an upper centrifugal washing settler.

[0025] 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 directed to a purification unit. Fication to obtain: • purified salt-laden water to be extracted; • purified water which is sent to the heat exchanger called "water".

[0026] A reverse osmosis unit or membrane distillation unit or an evaporation / condensation unit or a water absorption unit by a temperature-regenerable organic phase or a combination of at least two of these units can be used as a purification unit (the ; 2e).

[0027] At least one of the following can be used as washing water: • distilled water; • fresh water; • desalinated water; and • at least a portion of the purified salt-laden water to be extracted from the purification unit.

[0028] The initial decanter-separator and / or the washing decanter-separator when it is present may be centrifugal decanter-separators, allowing the removal of a phase composed of solid waste or of aqueous removal of the organic phase or the separation of a liquid biphasic medium having an aqueous phase / organic phase ratio of less than 0.2.

[0029] The salt to be extracted can be chosen from sodium salts, such as NaCl, NaCN, NaNO3, NaBr, Nal, potassium salts, such as KC1, KCN, KNO3, KBr, Kl, lithium salts or other salts that are selectively extractable in the presence of other salts.

[0030] The following examples illustrate the present invention without however limiting its scope. Example 1

[0031] The installation of [Fig.1] is intended to treat a 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.

[0032] The ionic composition of the brine S, that of the organic formulation O and that of the aqueous phase A vary throughout the process, receiving the successive notations respectively SI to S4, O1 to O6 and Al and A2.

[0033] The desired product is in the form of a hot aqueous phase PI charged at least with the salt to be extracted, which, after cooling, is recovered as production effluent P2.

[0034] The installation of [Fig. 1] comprises: - three extraction agitator-mixers 1a, 2a, 3a; each of these extraction agitator-mixers ensures an intimate mixing of the aqueous and organic phases addressed to them; - three phase separators Id, 2d and 3d, associated respectively with extraction mixers la, 2a, 3a; these phase separators - the first of which can be of the centrifugal type, then being noted Idc - allow the separation of the organic and aqueous phases of a stream after the latter has passed through the respective extraction mixer la, 2a, 3a; - two heat exchangers, the 2nd, each allowing the heating and cooling respectively of the organic phase and the aqueous phase addressed to them, the exchanger being the main exchanger; - two differential contactors le, 2c, which are here static or agitated or pulsed columns, each allowing the hot desorption or deextraction of the organic phase addressed to it of at least the salt to be extracted previously absorbed by the organic phase in the extraction agitator-mixers la to 3a; - a heating unit 1b, whose role is indicated below.

[0035] The brine to be treated SI is mixed in the stirred reactor Ia 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 I. It can also allow the removal of aqueous carryover in the organic phase O4 (not shown in [Fig. 1]) or the removal of a third phase composed of waste, particularly solid waste, before the effluent O4 is sent to the heat exchanger I.

[0036] The brine S2 is then mixed in the stirred reactor 2a with the organic phase 02 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 gravity separation of the two liquid phases, and the generation of the effluent S3 sent to the stirred reactor 3a and the effluent 03 sent to the stirred reactor 1a.

[0037] The brine S3 is then mixed in the stirred reactor 3a with the organic phase 01 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 gravity separation of the two liquid phases, and the generation of the treated effluent S4 (the raffinate) and the effluent 02 sent to the stirred reactor 2a.

[0038] The organic phase 04, loaded with at least the salt to be extracted, is then pumped to the main heat exchanger where it is heated to obtain the hot effluent 05 loaded with the salt to be extracted and then sent to the top (it is assumed here that the density of 05 is greater than the density of PI) of columns 1 and 2c, operating in parallel, to the production at the bottom of these columns of two organic effluents which are then combined to form O6, a hot salt-free organic phase, regenerated and pumped to the main heat exchanger to be cooled and then recovered as organic effluent O1.

[0039] The aqueous phase Al, generally fresh, desalinated or deionized water, is pumped to the secondary heat exchanger 2e where it is heated to obtain the hot effluent A2 which is then sent by bottom injection of columns le and 2c, operating in parallel, for the production at the top of the two columns of two aqueous effluents, loaded at least in salt to be extracted, which are then combined to form the production PI, a hot aqueous phase loaded in 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.

[0040] The heating unit 1b allows compensation of the heat losses generated by the temperature difference of the effluents at the cold end of the heat exchangers 1 and 2e.

[0041] Columns 1 and 2c are preferably used with a continuous descending organic phase and an ascending dispersed aqueous phase, with a high organic / aqueous phase (O / A) flow rate ratio during regeneration (hot water desorption), greater than 5, preferably greater than 10, potentially greater than 15, or even greater than 20. This minimizes the consumption of fresh water for desorption of the salt to be extracted while also increasing the reconcentration of the salt to be extracted between the water to be treated and the regeneration or desorption water to be produced. A higher regeneration temperature promotes higher O / A ratios.

[0042] The aqueous effluent PI does not experience any transfer of organic phase droplets, just as the tops of columns 1 and 2c do not experience any organic gaseous sky due to the installation at the top of these columns, in the aqueous phase, of systems enabling the coalescence of any organic phase associated with an appropriate residence time for the organic matter to settle. It should be noted that this design is also relevant due to the absence of organic gaseous phase at the column top, even if the O / A flow rate ratio during regeneration is high.

[0043] It can also be considered that aqueous effluents S4, PI or P2 can be subjected to the removal of organic traces by the installation of an oil-water separation unit which can be a coalescer combined with an adsorption unit using adsorbents which can be activated carbon, silica gel, diatomite earth and / or another similar approach allowing the recycling of organic components back to the process and / or their transport, with the adsorption medium, to an incineration disposal facility.

[0044] This cycle is denoted 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 hot water regeneration-desorption. Example 2

[0045] The installation of [Fig.2] is identical to that of [Fig.1] except that the decanters-separators 2d and 3d are replaced by centrifugal decanters 2dc and 3dc and that the heating unit 1b, necessary to compensate for the heat loss at the cold ends of the exchangers 1 and 2e, is placed on the organic effluent 06 in place of the aqueous effluent A2. Example 3

[0046] The installation of [Fig.3] is intended to treat a 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.

[0047] The ionic composition of brine S and that of the organic formulation O vary throughout the process, receiving the successive notations SI to S6, and 01 to 08 respectively.

[0048] The desired product is in the form of a hot aqueous phase PI charged at least with salt 1, which, after cooling, is recovered as production effluent P2, and after reconcentration, is recovered as production effluent P3.

[0049] The installation of [Fig.3] comprises: - five mixer-agitators la, 2a, 3a, 4a, 5a; these mixer-agitators each ensure an intimate mixing of the aqueous and organic phases addressed to them, mixer-agitators la, 2a, 3a and 4a being extraction mixer-agitators and mixer-agitator 5a being a washing mixer-agitator; - five phase separators Id, 2d, 3d, 4d and 5dc, associated respectively with the mixers la, 2a, 3a, 4a and 5a, these 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 a stream after the latter has passed through the respective mixer la, 2a, 3a, 4a and 5a; - two heat exchangers, the 2nd, each allowing the heating and cooling respectively of the organic phase and the aqueous phase addressed to them, the exchanger being the main exchanger; - two differential contactors, 2c, which are here static, agitated, or pulsed columns, each allowing desorption or deextraction at hot of the organic phase which is addressed to it of at least the salt to be extracted previously absorbed by the organic phase in the agitator-mixers la to 4a; - a heating unit 1b, whose role is indicated below; - two units of reverse osmosis and / or membrane distillation and / or evaporation / condensation and / or water absorption by a regenerable organic phase at temperature If, ​​2f allowing the separation of a retentate or concentrate or refiner and a permeate or condensate or extract.

[0050] The brine to be treated SI and the aqueous retentate / concentrate / raffinate S8 are mixed in the stirred reactor Ia with the organic phase O4 to produce a dispersed organic phase in a continuous aqueous phase (or vice versa). The two-phase mixture is then transferred to the decanter Id for the separation of the two liquid phases, and the generation of effluents S3 sent to the stirred reactor 2a and O5 sent to the stirred reactor 5a.

[0051] The brine S3 is then mixed in the stirred reactor 2a with the organic phase 03 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 effluent S4 is sent to the stirred reactor 3a and 04 to the stirred reactor 1a.

[0052] The brine S4 is then mixed in the stirred reactor 3a with the organic phase 02 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 effluent S5 is sent to the stirred reactor 4a and 03 to the stirred reactor 2a.

[0053] The brine S5 is then mixed in the stirred reactor 4a with the organic phase 01 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 treated effluents S6 (the raffinate) and 02 are sent to the stirred reactor 3a.

[0054] The organic phase 05, containing at least some of the salt to be extracted, is then mixed in the stirred reactor 5a with the aqueous phase A7 to produce 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 generated effluent S7 is sent to a water recovery unit 5dc, such as a reverse osmosis unit, a membrane distillation unit, and / or an evaporation / condensation unit, and / or a water absorption unit using a temperature-regenerable organic phase, and S6 is sent to the main heat exchanger. This step aims to wash the organic phase 05 with a very low flow rate. water A7 to remove salt impurities 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 aqueous carryover in the organic phase 05 or allow the removal of a third phase composed of solid waste or other, before the effluent 06 is sent to the heat exchanger le and before the effluent S7 is sent to the water recovery unit If.

[0055] The aqueous phase A7, generally composed of the aqueous phase A5 which is distilled water, fresh water or desalinated water and recycled water A6, is used as a minor flow rate compared to the organic flow rate, with an A / O flow rate ratio for water washing (water washing with neutral pH water, pH not having an impact on washing performance) of less than 0.25, preferably less than 0.1, potentially less than 0.05, or as low as 0.01. This makes it possible 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 overall loss of extraction yield of the salt to be extracted.

[0056] 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 a temperature-regenerable organic phase If is then separated into a brine S 8 (retentate or concentrate or refineate) sent to the stirred reactor 1a and into an aqueous effluent A6 (permeate or condensate or destract) recycled to the inlet of the stirred reactor 5a.

[0057] The organic phase 06, loaded with salt to be extracted, is then pumped to the main heat exchanger where it is heated to obtain the hot organic effluent 07 loaded with salt to be extracted and then sent to the top of columns 1 and 2c, operating in parallel, for the production at the bottom of columns of two organic effluents which are then combined to form the effluent 08, a hot organic phase without salt, regenerated and pumped to the main heat exchanger to be cooled and then recovered as cold organic effluent 01 (it is assumed here that the density of 07 is greater than the density of PI).

[0058] The aqueous phase A1, generally fresh water, desalinated or deionized water, which can be combined with the permeate-condensate A2 from 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 heated to obtain the hot aqueous effluent A4 which is then sent by bottom injection to columns 1e and 2c, operating in parallel, for the production at the top of the two columns of two aqueous effluents, loaded with at least 1% of the salt to be extracted which are then combined for the formation of production PI, a hot aqueous phase charged with at least salt to be extracted and pumped to the secondary heat exchanger 2e to be cooled 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-destract flow A2 and a concentrated production effluent P3.

[0059] The heating unit 1b allows compensation of the heat losses generated by the temperature difference of the effluents at the cold end of the heat exchangers 1 and 2e.

[0060] Columns Ie and 2c are preferably used with a continuous descending organic phase and an ascending dispersed aqueous phase, with a high O / A flow rate ratio during regeneration (hot water desorption), greater than 5, preferably greater than 10, potentially greater than 15, or even greater than 20. This minimizes the consumption of fresh water for desorption of the salt to be extracted while simultaneously increasing the reconcentration of the salt to be extracted between the water to be treated and the regeneration or desorption water to be produced. This O / A ratio can be even higher as the regeneration temperature of the organic phase in Ie and 2c increases.

[0061] The aqueous effluent PI does not experience any transfer of organic phase droplets, just as the tops of columns 1 and 2c do not experience any organic gaseous sky due to the installation at the top of these columns, in the aqueous phase, of systems enabling the coalescence of any organic phase combined with an appropriate residence time for the organic matter to settle. It should be noted that this design is also relevant due to the absence of organic gaseous phase at the column top, even if the O / A flow rate ratio during regeneration is high.

[0062] Equipment 5a and 5dc can be combined into a single unit commonly called a centrifugal extractor. This unit can be doubled, tripled, or quadrupled by connecting it in series to improve the washing of 1'05 in order to generate, as effluent PI, a product of greater salt purity for extraction. This unit can be multiplied and used in parallel to process effluents 05 with increasing flow rates.

[0063] It can also be considered that aqueous effluents S6, S7, PI, P2 or P3 can be subjected to the removal of organic traces by the installation of an oil-water separation unit which may be 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 organic components back to the process and / or their transport, with the adsorption medium, to a incineration disposal facility.

[0064] All effluents located to the left of heat exchanger le and below heat exchanger 2e are operated near ambient temperature (5 to 35°C) when the other effluents (top right of [Fig.3]) operate below the boiling point of water and brine effluents A4 and PI for operating pressures of le and 2c.

[0065] A variant of this configuration may be implemented when the aqueous saline or brine phase to be treated, due to its intrinsic properties of density, viscosity and / or other characteristics, does not allow for sufficient gravity settling to be economically viable. In this case, instead of the gravity-fed extraction mixer-settlers Id, 2d, 3d and 4d, centrifugal mixer-settlers Idc, 2dc, 3dc and 4dc are preferably used, or potentially centrifugal extractors with an increased internal mixing time compared to the current state of the art.

[0066] This cycle is denoted 4-1-Z, where 4 is the number of unit operations for extracting the salt, 1 is the number of washing operations, and Z is the number of theoretical stages for regeneration-desorption of the organic phase in hot water. Example 4

[0067] In the case of brine treatment with low relative concentration of the salt to be extracted compared to the other salts present, it may prove necessary to implement several cold washing / scrubbing stages in series in order to improve the washing and removal of impurities from the organic phase.

[0068] The installation of [Fig.4] is therefore identical to that of [Fig.3] with the presence of two additional agitator-mixers 6a and 7a to which are associated centrifugal decanter-separators 6dc and 7dc.

[0069] The organic phase 05, containing the salt to be extracted, is then mixed in the stirred reactor 7a with the brine S10 to produce 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 generated effluent S7 is 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 is sent to the agitator-mixer 6a.

[0070] The aqueous phase S7, pumped to 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 If, ​​is then separated into a brine S8 (retentate, concentrate, or refineate) sent to the stirred reactor and into an aqueous effluent A6 (permeate, condensate, or destractate) recycled to the reactor inlet agitated 5 a.

[0071] The organic phase 06, containing the salt to be extracted, is then mixed in the 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 the centrifugal decanter 6dc for the separation of the two liquid phases, and the generation of effluents S10 sent to the stirred reactor 7a and S17 sent to the agitator-mixer 5a.

[0072] The organic phase 07, containing the salt to be extracted, is then mixed in the 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 the centrifugal decanter 5dc for the separation of the two liquid phases, and the generation of effluents S9 sent to the stirred reactor 6a and 08 sent to the heat exchanger 1.

[0073] The addition of one, two or three agitator-mixers for the cold washing of the organic phase improves the purity and selectivity of the extraction of the salt to be extracted.

[0074] This cycle is denoted 4-3-Z, where 4 is the number of unit operations for extracting the salt, 3 is the number of washing operations, and Z is the number of theoretical stages for regeneration-desorption of the organic phase in hot water. Example 5

[0075] In this example the salt to be selectively extracted is sodium chloride NaCl.

[0076] The organic phase O comprises: • as an extractant of the Na+ ion, sodium ionophore X, also called 4-tert-butyl-Calix[4]arene acid tetraethyl ester (C6oH8oOi2, CAS No.: 97600-39-0) at a concentration of 0.275 M (mol / L) and whose log K in methanol at 25 °C is 5; • as an anionic solvent, N-(3,5-dichlorophenyl)octanamide (C14H19Cl2 NO, CAS No.: 20398-46-3) at a concentration of 0.825 M and with a pKa in water of 13.83 + / - 0.70; and • l-chloro-2-bromo-benzene as a diluent.

[0077] The organic phase has a density of 1.48 g / L at 20°C.

[0078] The brine to be treated comprises sodium chloride to be extracted and potassium chloride. The salinity of this brine is 209.81 g / L, of which 2.26 g / L is Na+ and 107 g / L is K+.

[0079] 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 strong relative extraction of sodium, in the form of NaCl, relative to potassium, or KCl, was obtained according to the graph in [Fig.5].

[0080] This extraction of sodium is selective without altering the concentration in

[0081] potassium as shown in [Fig.6]. 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 sodium concentration which is 168 mg / L in the brine treated with an (O / A)extraction ratio of 1. Figure 7 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 lower.

[0082] Next, in order to demonstrate the very low retention of KC1 within the organic phase, the formulation loaded with Na, K, Cl was considered after extraction at the ratio (O / A) extraction of 0.5 and a water wash 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) wash.

[0083] The results obtained are shown in [Fig. 8]. It appears that upon contact with 1% water by volume, 80.8% of the KCl absorbed in the organic phase is transferred back into the water, compared to only 1.9% of the NaCl. It also appears that with 4% water by volume, in a single washing step, 94.8% of the KCl is desorbed / transferred to the water, compared to 6.4% of the NaCl.

[0084] It can also be observed that NaCl is difficult to desorb at room temperature with only 21.1% of NaCl transferred to water when the water volume is 20% of the organic phase.

[0085] To enable the NaCl to pass into the aqueous phase and to regenerate the organic phase, a hot regeneration process is used at 80°C. This process involves washing the formulation with 4% water by volume, by contacting it with hot water volumes of 2%, 5%, 10%, 20%, and 50% of the volume of the organic phase, this ratio being denoted (A / O)regeneration*.

[0086] The results obtained are shown in [Fig.9]. 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).

[0087] Figure 10 shows the ionic concentrations of the brine obtained after regeneration at 80°C. After a washing step with a (A / O) washing ratio of 4%, the brine from the hot regeneration for these various (A / O) regenerations consists of 93% to 99% by mass of NaCl, which allows for NaCl purification with minimal NaCl loss.

[0088] A parametric study was carried out on cycles 4-1-3 (cycle of Example 3 with 3 theoretical stages for regeneration-desorption with hot water) and 4-2-3 (cycle with two cold water washing operations and 3 theoretical stages for regeneration-desorption with hot water).

[0089] The results are shown in [Fig. 11].

[0090] By extending the curves to the left, i.e., at an (A / O)iavage ratio tending towards 0, we would obtain the performance of a 4-O-3 cycle without water washing. It is clear here that this washing, achieved by increasing the (A / O)iavage ratio, makes it possible to increase the purity of the NaCl produced at the outlet of the regeneration columns up to a high plateau of approximately 94% NaCl mass (excluding water). Moreover, for sufficiently high (O / A)extraction ratios, the NaCl extraction rate at 20°C can be maintained above 90% while increasing the purity of the NaCl extracted by washing with water at room temperature (20°C here). For this 4-1-3 cycle, the selected sizing point allows a NaCl extraction yield of 93.2% for a NaCl purity of 91.8% by mass, with an (O / A)extraction ratio of 1.1 and an (A / O)iavage ratio of 5% (circled points).

[0091] Figure 12 shows the effect of adding a second counter-current washing stage at room temperature. Adding a second room-temperature washing (scrub) stage, by switching from cycle 4-1-3 to cycle 4-2-3, allows for further improvement of process performance, achieving both a high sodium extraction yield and even higher purity of the NaCl produced (high plateau at 98% by mass of NaCl) at the outlet of the regeneration (stripping) stage.

[0092] For this 4-2-3 cycle, the selected sizing point allows a NaCl extraction yield of 92.8% for a NaCl purity of 97.4% by mass, with an (O / A)extraction ratio of 1.1 and an (A / O)extraction ratio of 5% (circled points at the intersection of the two curves).

[0093] Table 1 shows the material balance for a 4-1-3 cycle with an (O / A) ex traction ratio of 1.1, an (A / O)iavage ratio of 5% and an (A / O)regeneration ratio of 10%. Opt 51 52 56 A7 57 A3 P2 58 A6 P3 A2 Al Æ 4-1-:-: Brine b treat Seumue input 155A ftrffimt Input 5c ru b Output Serti Input Strip Output Strip Concentrate rate Concentrate Strip Concentrate Strip Input Price n Os motiq ue 234 tara 34tnra b ra Work 29 tnra 0 tara 52tsra 278 tara Otsra 299 tara Otsra Otsra 0 tara O,m (ton H20 / h) 1000.0 1006.8 1006.8 55.50 50.1 100 1000 SB. 49.2 19.3 90.8 19.3 55 Û.V (nO / h) 1124.4 113Q9 " 1127.4 55.1 55.3 11Q1 111.5 65 49.3 21.7 90.8 195 55 Temperature fC) 0 61981 3 29 7 44 0 318 182 0 0 0 Liquid mite (kg / L) 0,999 Compta ition liq lide, wate r NB+ 2 678 3024 181 0 6 610 0 22 696 62 952 °P 129 689 qo 0,0 0 K+ 179 97 2 7 2701 112158 op 15 434 qo 0.0 0 SO4-- 4 076 4053 4052 0 7 0 0 68 0.0 0 0.0 0.0 0 cl- 443 198730 0.0 213990 0.0 0.0 0 NaCI fc / kg su ) 6.3 7.7 0.5 Op 16.8 OJO 57.7 16Q0 OP 3 29.7 qo 0.0 OP KCI 3. 36.31 u] 214 OP 5.1 213.8 OP 29.4 QO 0.0 OP KSO^fedeau] 7.4 7.4 7.4 O O OP 0.0 OP 0.1 OP 0.0 qo 0.0 0.0 Compcs ition liqiide,%01V. Na+ 0.ZV 0.64V 0.00V 214V 4582* QOCV 9.542V OjOOV 0.00V oxcv K+ 13.3 Z* 13.29V 13.31V 0.00V OjO1 1.13 V O.JOÜV 0.3V 1.3V Q8 O,OOV OjüOV 504,.0.30V 030V 030V 0.00V 0.00V OJOOV 0.0DW 0.006V QOOV 0 / »Z* O.OOV 0.0ÜV OJÜÜV cl- 1216 V 1218 V 11.87 V 0.00V 2 CIV 0 / KV 3.5ZV 14.465V QOCV 15.745 V OjOCV 0.00V H 20 74J0Z* 7 4JO1V 74-51V 100JÜCV 96.2Ztf 100J00V 94.09V 72785V 100.00V 73577V 100.00V 10Q0CV 100JÜÜV Composition liqLide,Wm (water excluded) NaCl 1.94V 219V 0.13V 0.00V 4280V OjOOV 91.81V 42S0V QOOV 91.81V 0.00V OjOÜV O.JÜÜV KOI 95.95V 95.7 Z* 97.7 Z* o,oov 57.17V OjOOV 8.19 V 57.17V QOCV 8.19V OJOCV 0.00V oxcv K2904 2HV 209V 215V 0.00V 0.00V 0.JO0V o,oov OjOSV QOCV OjOOV 0„0ÜV OjOÜV OJOOV Composition liqlide, rr^ / L solution Na+ 2382 2690 161 0 6579 0 22383 55 516 0 114907 0 0 0 K+ 160066 159 712 159371 0 11 721 0 2664 98 911 0 13 674 0 0 0 SO4-- 3 63 3604 3615 0 7 0 0 60 0 0 0 0 0 cl- 146140 146 308 142092 0 20 766 0 36 933 175 37 0 189600 0 0 0 H 20 8S9394 889 350 892119 999000 995 37 999000 986 30 881887 999000 886023 999 000 999 000 999 000 Débite hnHueE.kg / h Na+ 2 67S 3042 18 2 0 364 0 2497 364 0 2497 0 0 0 K+ 179972 180620 179675 0 6 48 0 297 648 0 297 0 0 0 SO4-- 4076 4076 4076 0 0 0 0 0 0 0 0 0 0 cl- 164314 165 46 2 160195 0 1148 0 4119 1148 0 v 4119 0 0 0 H 20 1 000 000 1005775 1006 775 55000 55000 110000 110000 5775 49 23 19 30 90750 19 30 5775 . 0094] [Tables 1]

[0095] This table shows that the raffinate S6 contains only 0.13% mass of NaCl to obtain a purity of 99.87% mass (excluding water) of KC1 and K2SO4, the quantity of K+ ions in it being practically unchanged compared to the quantity present in the brine to be treated SI.

[0096] Table 2 shows the material balance for a 4-2-3 cycle with an (O / A) ex traction ratio of 1.1, an (A / O)kvage ratio of 5% and an (A / O)regeneration ratio of 10%. Cycle SI 52 S6 A7 S7 A3 P2 50 A6 F5 A2 Al Æ f 4-2-3 SHLmLreB SsLmLTe input traiter T5SA ftrffirst Input Scrub Output Scrub Input Strip p Output Strip Concerted Corde rrest Stretch Concentrate Osmotic 234tar 234tar 227 tar Order 33 tar Order 50tar 279 tar Order 33 tar Order 0 tar 0 tar Q,m (good H20 / h) 100Q0 1007.4 1007.4 ' 55.0 55.4.0 110 13.7 a 1.3 13.7 7.4 Qy (m3 / h) 1124-4 - 11328 ' 1129.2 55.1 55.5 110.1 111.4 3.4 476 21.0 91.4 3.4 re (AC 2 p-7.7) 20 20 20 23 23 Total salinity (mg / Lj Lks / Ll 1.202 1.232 1.197 0.999 1X41 0.999 1J046 1.212 0.999 1.198 Q999 0.999 0.999 Compcsiticfi liquide, rrg / lç uater Na+ 2 673 313 192 0 3 307 0 22539 65 2 3 3 qo 132374 0.0 0 / 3 0 K+ 179 972 18 5Æ 179 776 0 322 109 455 qo 4336 0,0 op 0 SOi- 4076 4046 4046 0 7 0 0 53 qo 0 0,0 op 0 cl- 164 314 164 576 3904 21 199 309 qo 209 33 0.0 qo 0 NaCI Ë / kgsu) 6.3 8.0 0.5 □HO 22* HOURS 57.4 165.3 qo 3378 0.0 HO 0.0 KCI ^Ib=8 3 3 3 23.2 0..0 1.6 2Œ.6 qo 9.2 0.0 EYE EYE K2904fe / iBŒu) 7.4 7.3 7.3 EYE EYE EYE 01 QO EYE 0.0 0.0 EYE Nae+QZ0 liquid, “Ccsiticfi 0J01V Q.OOV 084V QOOV 213V 4.75V oxov 9.86V Ü.JOCV ÜJÜÜV OJOOV K+ 13.3 13.23V 13.3 IV QOOV 1.41V QOCV 0, QOff6 O.JOCV OjOOV OjOÜV S04- Q30V O^OV 0,30V Q.OOV oxov QOOV QOCV OJOOV o,ocv 0,00V OJOCV Ü.OCV OJOOV cl- 12.16 V 12.18 V ii^ev QOOV 257V QOCV 3.36V 14.54V OjOOV 155 4V 0,LÜV ojæv OJOOV H 20 7 4J01V, 74 95.19V 103.03V 9 4.43 V 72.75V lOOjOOV 7 4.24V 10Q0CV lOOOCV lOOjOOV Compas iticn liq nië, %m (es ue cell la ) NaCI 1.9« .001V Q4 9734V 44.3V oxov 9734V 0.OÜV OjOÜV OJOOV KCI 95^5V 95.6 4V 97.71V QOOV 55.70V QOOV 266V 55.70V OjOOV 2.66V OJOOV OJOOV OJOOV K 290 4 211V 2.09V 2.15V QOOV oxov QOÜV QOCV OJO0V 0X0V 0.00V OJODV occv OXÜV Liquid composition, mj / L solution Na+ 23S2 2792 171 0 3727 0 22311 57515 0 113 216 0 0 0 K+ 160 066 159 593 159 300 0 14643 0 312 96505 0 4302 0 0 0 SOi- 363 3 59S 3609 0 7 0 0 47 0 0 0 0 0 cl- 146 140 146 363 142 OŒ 0 26 730 0 35 142 176 16E 0 186 204 0 0 0 H 20 SS 394 339 336 392161 999 000 990952 999ÜOO 987696 331 683 999 cm 339683 999 000 999000 999030 Débite ioriqiES, kg / h Ma+ 2678 3 163 194 0 434 0 2435 484 0 2485 0 0 0 K+ 179 972 130 735 179 3S1 0 313 0 90 313 0 90 0 0 0 SOi- 4076 4076 4076 0 0 0 0 0 0 0 0 0 cl- 164 314 165 790 160 400 0 1 434 0 3914 1484 V 0 * :: 3914 0 0 H 20 1000 000 1007 43 1007 43 55 000 55 000 110000 110000 7 43 47 5 75 13 703 91300 18 703 743 . 0097] [Tableaux2]

[0098] By comparison 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). Example 6

[0099] In this example the salt to be extracted selectively is potassium chloride KC1.

[0100] The organic phase O comprises: • as an extractant of the K+ ion, 4-tert-butyl-Calix[6]arene acid hexaethyl ester (C9oHi2oOi8, CAS No.: 92003-62-8) at a concentration of 0.175 M (mol / L) and whose log K in methanol is 4.8; • as an anionic solvent, N-(3,4-dichlorophenyl)octanamide (Ci4Hi9C12NO, CAS No.: 730-25-6) at a concentration of 0.525 M and with a pKa in water of 13.63 + / - 0.70; and • l-chloro-2-bromo-benzene as a diluent.

[0101] The organic phase has a density of 1.43 g / L at 20°C.

[0102] The brine to be treated comprises potassium chloride to be extracted and sodium chloride. The salinity of this brine is 168 g / L, of which 62.7 g / L is Na+ and 4.53 g / L is K+.

[0103] 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 0.1; 0.4; 0.8; 1.5 and 4, a good relative extraction of potassium, in the form of KCl, relative to sodium, or NaCl, was obtained according to the graph in [Fig. 13].

[0104] This potassium extraction is selective without altering the sodium concentration, as shown in [Fig. 14]. The sodium concentration remains stable at around 62 g / L for the tests at (O / A)Extraction ratios of 0.1, 0.4, 0.8, and 1.5, and the potassium concentration is 2.8 g / L in the brine treated with an (O / A)Extraction ratio of 1.5.

[0105] Figure 15 shows the ionic concentrations of the different compounds in the organic phase after extraction at 20°C. KCl is predominant in the organic phase for an (O / A)extraction ratio of less than 2.

[0106] Next, in order to demonstrate the very low retention of NaCl within the organic phase, the formulation loaded with Na, K, Cl was considered after extraction at the ratio (O / A) extraction 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)iavage.

[0107] The results are shown in [Fig. 16]. It appears that upon contact with 1% of the volume of water, 48.3% of the NaCl absorbed in the organic phase returns to the water against only 2.1% of the KCl absorbed.

[0108] It also appears that with 4% water by volume, in one washing step, 76.4% of the NaCl is desorbed / transferred to water against 4.9% of the KC1 (and 90.3% of NaCl desorbed at 10% water for 8.9% of KC1 desorbed).

[0109] It can also be observed that KC1 is difficult to desorb at room temperature with only 14.8% of KC1 transferred to water when the water volume is 20% of the organic phase.

[0110] To enable the KC1 to pass into the aqueous phase and to regenerate the organic phase, a hot regeneration process is used at 80°C. The formulation obtained from washing with 4% water by volume is carried out by contacting hot water volumes of 2%, 5%, 10%, 20%, and 50% of the volume of the organic phase, this ratio being denoted (A / O Regeneration*).

[0111] The results obtained are shown in [Fig.17]. It appears this time that with 20% by volume of water, 83.7% of the KC1 was desorbed from the organic phase (compared to 14.8% previously at 20°C).

[0112] Figure 18 shows the ionic concentrations of the brine obtained after hot regeneration at 80°C. With the washing step at a (A / O)wash ratio of 4%, the brine from the hot regeneration consists of 78% to 90% by mass of KCl, which allows for KCl purification combined with KCl production for valorization.

[0113] A parametric study was carried out on these cycles 4-1-3 (cycle of Example 3 with 3 theoretical stages for regeneration-desorption with hot water) and 4-2-3 (cycle with two cold water washing operations and 3 theoretical stages for regeneration-desorption with hot water).

[0114] The results are shown in [Fig. 19].

[0115] By extending the curves to the left, i.e., at an (A / O)iavage ratio tending towards 0, we would obtain the performance of a 4-O-3 cycle without water washing. It is clear here that this washing, achieved by increasing the (A / O)iavage ratio, makes it possible to increase the purity of the KC1 produced at the outlet of the regeneration columns up to a high plateau of approximately 91% KC1 mass (excluding water). Moreover, for sufficiently high (O / A)extraction ratios, the KC1 extraction rate can be maintained above 90% while increasing the purity of the KC1 extracted by washing with water at room temperature (20°C here). For this 4-1-3 cycle, the selected sizing point allows a KC1 extraction yield of 85.0% for a KC1 purity of 78.5% by mass, with an (O / A)extraction ratio of 2.2 and an (A / O)watering ratio of 5% (circled points).

[0116] Figure 20 shows the effect of adding a second counter-current wash stage at room temperature. Adding a second wash stage at room temperature (scrub), by switching from cycle 4-1-3 to cycle 4-2-3, allows for further improvement. of the performance of the process to obtain both a high potassium extraction yield and an even higher purity of the KC1 produced (high plateau at 98% by mass of KC1) at the outlet of the regeneration stage (stripping).

[0117] For this 4-2-3 cycle, the selected sizing point allows a KC1 extraction yield of 85.6% for a KC1 purity of 91.5% by mass with an (O / A)extraction ratio of 2.2 and an (A / O)watering ratio of 5% (circled points)

[0118] Table 3 shows the material balance for a 4-1-3 cycle with an (O / A) ex traction ratio of 2.2, an (A / O)iavage ratio of 5% and an (A / O)regeneration ratio of 10%. Cycle 51 52 56 A7 57 A3 P2 S8 A6 P3 A2 Al A5 , .. Brine h treat Pr«ŒÎon Osmotic 137 tara Qm (ton H2O / h) lOOOjO Brine e inlet TSSÛ 139 tara 1011.6 feffirBt 130 tara Scrub Inlet Ota ra Scrub Outlet 31 tara Strip Inlet 0 tara Strip Outlet 27 tara Concentrate 35 tara Cord rente 0 tara Concentrate 36 tara Cord mte Strip Inlet 0 ta ra 0 ta ra Scrub Inlet 0 tara 1011.6 f 110.0 110J0 " 220.0 220.0 11J5 98.5 3.3 194.7 3.3 11.6 (m³ / h) Temperature (l / L) Total Salinity (l / L) 155 247 10659 20 157 207 1.106 1061.1 20 146 061 110.1 20 0 110.4 20 37 211 220.2 20 0 221.0 22 38374 130 22 315 641 98.5 20 0 283 22 39580 1949 33 20 20 0 0 11.6 20 0 Liquid Volume (l / L) 1.106 1099 0.999 1.033 0.999 1.034 1.203 0.999 1.193 0.999 0.999 0.999 Com poe rtion 1 iq ukie, wate r Na+ 61 222 K+ 4105 61956 4 278 59 812 609 0 0 13184 2018 0 0 3 266 2 15861 19 ^ 529 137 9 18 on oo op op 0 0 cl- asiM " 99 423 92790 0 22161 0 19 419 211 067 0.0 858 oo 00 0 NaCI fe / lç au) 155.5.08 1.302.1 319.2 hours 722 hours 0.0 hours KCI fe / kg ea u) 7.8 8.2 1.2 0.0 3.3 hours 30.2 36.7 hours 263.0 hours 0.0 hours &a2|e4.0 hours 0 hours QO 0.0 0.0 EYES 0.0 oo oo EYES Liquid composition, %m Na+ 5.26V 5.32V 5.19 V Ü.CCV 1.27V o,ocv 03IV 9.31V o.ocv 2.127 V O.31V Q+LCV^OCV 0.05V OjÜCV 0.19V O.ÜOV 1.53V 1.418 V o.ocv 10330V O.LCV O.LCV o.ocv Cl- 8.43V H2O 85.95V 8.53V 85.71V OjVO 8.ŜLCO 2.14V 96.40V o.ocv 1CO..0CV 137V 96.3V 15.567V 73.755V o.ocv ilci.o:v 12647V 7 4,896V OCCV OJOOV lOOjCCV 1Ü0O0V 0,00V 100, oov Compas rtion liquide, %m (eau exclue) NaCI 95.21V 95J0BV 99.24V O.LC0 89,7VV 134V oc. 89.70V ov 21.54V cv ov cv KCI 4.79V Oompa rtion liquid, rr^ / L sol il ion 492V 0.76V OJOQV 1Q30V o.ocv 78.46V 1030V ov 78.46 V ov cv Na 57021 0 13 130 0 3 32 111 375 0 3 385 0 0 0 K+ 3 893 4060 581 0 2010 0 15790 17051 0 13271 0 0 0 0 0 0 Cl- 0 293 H2Q19 94351 9 48 985 88 460 953 332 0 999OÜO 22071 995942 0 999OÜO 19332 995 549 187 215 887 033 0 99900 150 580 2499763 0 999000 999000 0 999OOÜ Ionic flow rates, kg / h Na+ 61 222 K 62672 60503 0 1450 0 719 1 450 222 2 438 " 11550 0 ..........0.....'....98 450 719 3 489 4272 3 300 0 0 0 K+ 4106 4327 616 0 222 0 3 489 0 0 0 cl- 98 134 " 100572 93862 0 2 438 0 4 272 0 0 0 H 20 1ÜÜOÜOO " 1011550 1011550 110000 110000 220000 220 000 194700 3 300 11550 . 0119] [Tables3]

[0120] This table shows that the raffinate S6 contains only 0.76% mass of KC1, the quantity of NaCl ions in it being practically unchanged compared to the quantity present in the brine to be treated SI.

[0121] Table 4 shows the material balance for a 4-2-3 cycle with an (O / A) ex traction ratio of 2.2, an (A / O)kvage ratio of 5% and an (A / O)regeneration ratio of 10%. Cycle SI 52 S6 A7 S7 .43 P2 58 A6 F5 A2 Al A5 Brine b treat Brine and inlet TSSÜ h-ffirst Entry Scrub Output Scrub Inlet Strip Output Strip Broken cord Conde remaining Broken cord Inlet Strip Scrub Output Scrub 137 tons 140 tons 130 tons 0 tons 41 tons □ tons 23 tons 34 tons Otara 227 tons Otara Otara Otara Qm (tons H20 / h) iLOj.O 1015.4 1015.4 * 11Q.0 11Q0 f 220.0 220.0 220.0 120.0 41J 154 1966 196.0 210 15.4 Qy (m3 / h) 105X9 1070 2 1065.3 11Q1 111.0 220.2 220.5 17.3 94.7 24.6 196.2 220 15.4 Temperature fC) Total salinity (n^ / L) 947 0 33 224 313 170 0 29793 0 0 0 De ns ite liquid (Ice / L) 0.999 0.999 Compcsiticn liqùde, rr^ / kg lester Na+ 61 222 62 219 60061 0 17 779 0 1119 126 993 0.0 11194 op 0 2. 0 53 4 23 ' 15 974 15 651 0.0 159 744 qo op 0 cl- 93 134 99331 93133 0 29 404 0 16 211 210 030 0.0 162113 HO op 15856.0 NaCI) 15X7 0 / 3 45.2 0.0 13 3228 0.0 Z.5 00 OP 0.0 MCI Is^»] 7 .B 8.2 1.1 Op 4.2 0.0 30.5 3.3 0.0 30.O 30.5 3.3 0.0 3O4J6 0.0 & uuj2 0.0 3O4J6 .00 & CI|0 0 / 3 0 / 3 0.0 0.0 0.(3 0.0 op QO OP 0.0 Liquid composition, W Tl Na+ 5.26V 5.33V 5.20V OjOOV 1^9V 0.00V 0.000CV 0..4V Oj. 0.00V K+ 035V 0.37V 0 / EV OCCV 0.21V O.OÜV 1.55V 1.157V O.OÜV 11.963V 0.00V 0,00V 0,0ÜV cl- 8,43V 3,56V 8,07V OjOOV 2,80V OJOOV 1,57V 15,527V OOCV 1X16IV OJOOV 0,00V 0,00« H 20 85,95V 85,74V 86,67V 100JO0V 95,29 V 10Q00V 96,78V 73,9 SV 10Q00V 75,016V lOOjÜCV 100,00V 100,00V Compcsiticn liqùde,9tim (eauexclue) NaCI 95,21V 95J09V 99, SV OjOÜV 91,54V 0,00V 8,54V 91,54V OV 8,54V OV OV OV KOI 4,79V 4,9 IV 0,72V OjOOV 8,46V OjOÜW 91,46 V 8,46 V OV 91,46 V OV cv OV Composition liquide, rrg / L sol utbn Na+ 53 145 59031 57 236 0 17 6S 0 1 117 112768 0 10014 0 0 0 K+ 3899 4061 555 0 217 2 0 15935 13 898 0 142896 0 0 0 cl- 93 203 94715 38 769 0 29 150 0 16 172 186 504 0 145015 0 0 0 H 20 949 751 948 750 953136 999000 9913 49 999 000 997 565 88 7 986 999 000 894532 999 000 999 000 999 000 Débits icniques, kg / h Na+ 61 222 63 177 60975 0 1956 0 246 1956 0 246 0 0 0 K+ 4105 4346 591 0 241 0 3514 241 0 3514 0 0 0 cl- 93 134 101 363 94567 0 3 234 0 3 566 3 234 0 3556 0 0 0 H 20 1000 000 1015 400 1015 400 110000 110000 220 000 220000 15 400 9 4600 22000 198 000 22000 15 400 . 0122] [Tables4]

[0123] By comparison with 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).

Claims

Demands

1. - A selective extraction process for a salt to be extracted comprising at a cation and at least one anion from a brine to be treated containing said salt to be extracted and salts other than the salt to be extracted, characterized by the fact that: • The brine to be treated (SI) is sent to an extraction agitator-mixer (la) into which a liquid hydrophobic organic phase is also introduced, said organic phase comprising: • at least one first organic compound that solvates the anion of the salt to be extracted, is protic and hydrophobic, and has a pKa in water at 25°C of at least 9; and • at least one second hydrophobic organic compound that selectively extracts the cation from 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 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), 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 is recovered from which the salt to be extracted has been removed; • the organic phase exiting the separator-decanter is addressed (Id; Idc) to a new agitator-mixer called a washing agitator (5a; 7a) in which it is mixed with wash water, and the mixture is sent to a decanter-separator (5dc; 7dc) called a washing decanter to obtain: • a purified organic phase loaded with salt to be extracted, which is returned to an "organic" heat exchanger where it is heated; and • an aqueous phase which is sent to a purification unit (If) to obtain: • a brine that has been re-tempered, concentrated, or refined, which is then sent to the extraction mixer-agitator initial (la); and • water which is re-addressed to the new washing agitator-mixer (5a; 7a); • the organic phase loaded with salt to be extracted, reheated in the so-called "organic" heat exchanger (le), is addressed to a column (le; 2c) called the regeneration column in which a counter-current exchange is carried out with hot water, which is at a temperature higher than that of the starting brine, to obtain: • water containing the salt to be extracted; and • a regenerated organic phase, having lost the salt to be extracted; said organic phase loaded with salt to be extracted being reheated in the heat exchanger (le) by the organic phase exiting the regeneration column (le; 2c), • the regenerated organic phase is recycled in the extraction agitator-mixer (la).

2. - A process according to claim 1, characterized in that before recovering the brine from which salt to be extracted has been removed, said brine is sent 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 sent again 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 before, said operation of extracting salt from the brine being able to be repeated at least once more, and the organic phase loaded with salt to be extracted, from each of the extraction decanter-separators (2d; 3d;4d), towards the initial extraction agitator-mixer (la), the regenerated organic phase arriving in each case in the last extraction agitator-mixer.;

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

4. - A method according to any one of claims 1 to 3, characterized in that the water which will serve as regeneration water is directed into a heat exchanger called "water" (2e) in which it is heated by the water loaded with salt to be extracted coming out of the regeneration column(s) (the ; 2c) if several are provided, to obtain heated regeneration water which is introduced, if necessary after a further heating, into the regeneration column(s) (the ; 2c), and cooled water loaded with salt to be extracted is obtained from said heat exchanger (2e).

5. - A process according to any one of claims 1 to 4, characterized in that before sending the washed organic phase from the washing decanter-separator (7dc) to the "organic" heat exchanger (le), said washed organic phase is sent to another washing agitator-mixer (6a) followed by a washing decanter-separator (6dc) which receives wash water, to obtain, after decantation, an organic phase which has undergone a further washing step which is sent to the organic heat exchanger (le), said washing operation of the organic phase being able to be repeated at least once more, and the final wash water, or the re-stained or concentrated or refined brine, after purification, is returned to the initial washing agitator-mixer (7a).

6. - A method according to any one of claims 1 to 5, characterized in that the salt-laden water to be extracted, which comes either from the regeneration column(s) (the ; 2c) or from the heat exchanger called "water" (2e), is sent to a purification unit (2f) to obtain: • purified salt-laden water to be extracted; • purified water which is sent to the heat exchanger called "water" (2e).

7. - A process according to any one of claims 1 to 6, characterized in that the purification unit (If; 2f) is a reverse osmosis unit or a membrane distillation unit or an evaporation / condensation unit or a water absorption unit by a temperature-regenerable organic phase or a combination of at least two of these units.

8. - A process according to any one of claims 1 to 7, characterized in that at least one of the following is used as wash water: • distilled water; • fresh water; • desalinated water; and • at least a portion of the purified salt-laden water to be extracted from the purification unit (If).

9. - A method according to any one of claims 1 to 8, characterized in that the initial decanter-separator and / or the washing decanter-separator when present are centrifugal decanter-separators, allowing the removal of a phase composed of solid waste or of aqueous removal of the organic phase or the separation of a liquid two-phase medium having an aqueous phase / organic phase ratio of less than 0.

2.

10. - A process according to any one of claims 1 to 9, 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.

11. - A process according to any one of claims 1 to 10, characterized in that the salt to be extracted is chosen from sodium salts, such as NaCl, NaCN, NaNO3, NaBr, Nal, potassium salts, such as KCl, KCN, KNO3, KBr, Kl, lithium salts or other salts selectively extractable in the presence of other salts.

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