Method for selectively extracting salt from saline water or salt water

Abstract

A method for the selective extraction of a salt to be extracted from salt water to be treated, the salt water to be treated containing said salt to be extracted, characterized in that: the salt water to be treated (S1) is fed into an initial extraction stirrer mixer (1a), in which the initial extraction stirrer mixer further introduces a liquid hydrophobic organic phase, mixing the salt water and the liquid hydrophobic organic phase in the initial extraction stirrer mixer (1a) and feeding the mixture into an initial decanter separator (1d, 1dc) to obtain: salt water from which the salt to be extracted has been removed; and an organic phase containing the salt to be extracted; recovering the saline water from which the salt to be extracted is removed; the organic phase containing the salt to be extracted is sent to an initial washing stirrer mixer decanting centrifugal separator (5a, 5dc or 7a, 7dc) to obtain a purified organic phase containing the salt to be extracted, which is sent to a so-called organic heat exchanger (1e), heated therein, and then sent to a so-called initial regeneration column (1c; 2c) in which a countercurrent exchange is carried out with hot water at a temperature higher than that of the initial saline water to obtain: water containing the salt to be extracted; and a regenerated organic phase after the salt to be extracted is lost; the regenerated organic phase is recycled to the initial extraction stirrer mixer (1a).

Description

Technical Field

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

[0002] The upgrading of alkaline, alkaline earth, transition metal or rare earth metal salts dissolved in natural or precipitated saline waters remains a relatively complex, costly and environmentally intensive operation, with considerable room for improvement.

[0003] The separation of various salts has historically been based on crystallization, which depends on the temperature variation of their respective solubilities and on their differences in water solubility, which, for example, at 50 °C follows the following order of increasing saturated moles:

[0004] Of course, for each brine to be treated, the number of ions and salts present is usually much smaller, which may result in solubility differences large enough to improve their separation capabilities. In fact, when separating alkaline and alkaline-earth salts, for example, the implementation of solar evaporation tanks or thermal crystallization systems allows the continuous precipitation of different salts according to their initial concentration and solubility in water. However, due to the variability of the mineral composition of the sources, their variability over time and the concentration of the brine to be treated, this creates a high variability in the crystallization process. This approach requires a good knowledge of the crystallization and co-crystallization conditions of the salts involved. Despite this, the mineral salts produced by such thermal crystallizations are produced in low to moderate yields of purity or are impure and require one or more additional downstream purification steps. Furthermore, given the high energy costs of water evaporation (682 kWh / ton at 20°C), even with the industrial implementation of modern evaporation processes combined with mechanical vapor compression and / or multiple-effect distillation, this method remains very expensive in terms of energy, i.e. between 20 and 60 kWh per ton of distilled water, or about 100 times higher relative to crystallized salts, considering the need to evaporate, for example, 75 to 150 tons of water per ton of crystallized sodium chloride. Therefore, in addition to applications involving salts of high economic value, solar evaporation ponds are used instead, which have the advantage of offering zero operating costs in terms of energy for water evaporation, but at the expense of a large amount of local water consumption, since more than 90% of the salt water to be treated is evaporated. This method is acceptable for seawater evaporation, but not for operations in desert, arid or inland areas.

[0005] Depending on the chemical composition of the water and the metals to be upgraded in the solution to be recovered, other separation strategies have been developed, in particular by hydrometallurgy, which is a liquid route metal treatment process that includes a stage in which the metals are dissolved for purification using acids (H2SO4) or oxidants (Cl2, H2O2…). This leaching or dissolution is particularly used for many transition metals such as zinc, copper, nickel, cobalt, manganese, rare earths, uranium, etc.

[0006] Once dissolved, some of them can precipitate as insoluble hydroxides, carbonates or sulfides. This is a precipitation-based method. These precipitates depend primarily on the pH of the medium, the solubility product and the redox potential of the medium. Most commonly, the pH of the solution is increased and precipitated hydroxides are formed. The metals are then recovered as solids by simple decantation within the limits of their solubility product, which typically results in concentrations at the process outlet that frequently exceed environmental discharge standards. In addition to the potential need for further processing downstream, this method may require large amounts of reagents (neutralizing bases and acids, or Na2CO3, etc.), or may create operational difficulties due to scaling problems, or may require upstream pretreatment to precipitate a product of acceptable purity, all of which can make this method unattractive from an economic point of view while generating large amounts of waste, which is typically stored in residue retention dams.

[0007] Another common method is ion exchange. This involves contacting the aqueous solution with an organic material or agent capable of exchanging the metal to be extracted, usually with protons H + or sodium + One disadvantage of this approach is that if Na + To perform ion exchange, the ion exchange resin, whether solid or liquid, is initially regenerated with acid to desorb the metal ions and then neutralized with NaOH soda before it can be used for extraction again. + / ionic metals, the pH of the contact water will be reduced and reagents will also need to be used to extract, wash and regenerate the material (neutralize the acid and / or base), which may be in large quantities. In both cases, this will incur huge operating costs and may use local chlorine and caustic soda plant facilities to produce these acids and bases. Another disadvantage may be that due to the regeneration or acid washing of the ion exchange material, if it is inorganic, it may also experience leaching and therefore degrade itself over time and may be discharged into the environment of transition metals and / or heavy metals Mn, Ti, Sn, Cu, Al, V or Sb, depending on the original composition of the selected ion exchange material. This is more problematic because the ecotoxicity thresholds of these metals are very low and they are difficult to capture completely.

[0008] Finally, we note that there is no direct extraction of neutral salts, nor is there any systematic addition of salts / wastes to the system and / or its environment, such as Na2SO4 produced by the bulk combination of 2NaOH and H2SO4.

[0009] Another approach is salt adsorption. This involves the physical or electrochemical fixation of metal salts to the surface or within the pores of an adsorbent material. Few applications have been developed so far, but we can cite the selective adsorption of lithium and capacitive deionization for desalination of low-salinity brackish waters. Naturally, adsorption is economically limited to the removal of small amounts of ions, i.e. 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, due to the need for very large porous adsorption surfaces. Summary of the invention

[0010] The aim of the present invention is to obtain a technical and economic solution capable of selectively extracting salts in a wide concentration range, with high extraction rates, energy consumption close to the theoretical minimum (related to the extraction of salts from water), the ability to produce such extracted salts with high purity, also with a water consumption close to the theoretical minimum and zero reagent consumption.

[0011] To this end, the present invention relates to a method for selectively extracting a salt to be extracted comprising at least one cation and at least one anion from brine or salt water to be treated containing the salt to be extracted and salts other than the salt to be extracted, characterized in that:

[0012] - The salt water to be treated is fed into a so-called initial extraction agitator mixer, into which a liquid hydrophobic organic phase is also introduced, said organic phase comprising:

[0013] o at least one first hydrophobic protic organic compound which solvates the anion of the salt to be extracted and has a pKa of at least 9 in water at 25°C; and

[0014] o at least one second hydrophobic organic compound which selectively extracts cations of the salt to be extracted and has a complex constant for the cations to be extracted, whose logarithmic K value in methanol at 25° C. is greater than 1, preferably greater than 2,

[0015] - mixing said salt water and said liquid hydrophobic organic phase in said extraction agitator mixer and sending the mixture to a so called primary extraction decanter separator to obtain:

[0016] ο salt water from which the salt to be extracted has been removed; and

[0017] o an organic phase containing the salt to be extracted;

[0018] - Recovering the brine from which the salts to be extracted have been removed;

[0019] - The organic phase leaving the initial decanter separator is sent to a new agitator mixer, called the initial wash agitator mixer, where it is mixed with wash water and the mixture is sent to a decanter centrifuge, called the initial wash decanter centrifuge, to obtain:

[0020] o the purified organic phase containing the salts to be extracted is sent to a so-called "organic" heat exchanger for heating; and

[0021] o Feeding the wash water loaded with impurities into the initial extraction agitator mixer;

[0022] - The organic phase containing the salts to be extracted, heated in the so-called "organic" heat exchanger, is sent to the so-called initial regeneration column, where it is exchanged in countercurrent with hot water of a higher temperature than the initial salty water, in order to obtain:

[0023] o water containing the salt to be extracted; and

[0024] o regenerated organic phase losing the salt to be extracted;

[0025] - said organic phase containing the salts to be extracted is heated in a heat exchanger by the organic phase leaving the regeneration column,

[0026] - The regenerated organic phase is recycled to the initial extraction agitator mixer.

[0027] In particular, the salt to be extracted can be extracted cold, ie at a temperature of -35° C. to +45° C. The regeneration in the regeneration column then takes place at a higher temperature.

[0028] In particular, the organic phase may be washed at a temperature of 5°C to 45°C.

[0029] The temperature difference between the extraction and regeneration steps is at least 30° C., preferably more than 40° C., even more preferably more than 50° C. The regeneration step may be carried out at a temperature ranging from 60° C. to 150° C.

[0030] In particular, the organic phase leaving the regeneration column can be heated before entering the "organic" heat exchanger.

[0031] In particular, the first organic compound that solvates the anions 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.

[0032] In a particular embodiment, the second organic compound is a lithium cation extractant compound selected from the group consisting of:

[0033]

[0034] in:

[0035] - R1 and R2 may be the same or different, regardless of their position on the nitrogen atom, and are independently selected from a linear or branched C1-C 12 Alkyl, aryl, C4-C8 cycloalkyl; or

[0036] - R1 and R2, together with the nitrogen atom carrying them, form a five-membered, six-membered, seven-membered or eight-membered ring;

[0037] -R3 is selected from hydrogen, linear or branched C1-C8 alkyl, C4-C8 cycloalkyl, C2-C6 alkoxyalkyl and alkoxyalkaryl;

[0038] -R4 is selected from hydrogen, linear or branched C1-C3 alkyl;

[0039] - R5 is selected from hydrogen, linear or branched C1-C3 alkyl;

[0040] - R6 is selected from hydrogen, linear or branched C1-C3 alkyl.

[0041] In a particular embodiment:

[0042] - R1 and R2 can be, regardless of their position on the nitrogen atom, independently selected from methyl, ethyl, n-propyl, isopropyl, n-butyl, isobutyl, sec-butyl, tert-butyl, n-pentyl, isopentyl, 2-methylbutyl, 2-ethylpropyl, n-hexyl, n-heptyl, n-octyl, n-nonyl, n-decyl, 2-ethylhexyl, phenyl, cyclobutyl, cyclopentyl, cyclohexyl, cycloheptyl, cyclooctyl; or,

[0043] - R1 and R2, together with the nitrogen atom carrying them, can form a pyrrolidine, piperidine, azepane or azoethane ring;

[0044] -R3 may be selected from hydrogen, methyl, ethyl, propyl, isopropyl, n-butyl, isobutyl, sec-butyl, tert-butyl, n-pentyl, isopentyl, 2-methylbutyl, 2-ethylpropyl, n-hexyl, cyclohexyl, methoxymethyl, methoxyethyl, methoxypropyl, methoxybutyl and -CH2-O-CH2-phenyl; and

[0045] - R4, R5 and R6 may be hydrogen or methyl,

[0046] If the saltwater to be treated has a calcium concentration greater than 10 g / L and / or preferably Li + / Ca 2+ Optionally, R1 and R2 are advantageously selected from butyl, pentyl, hexyl, heptyl, octyl, nonyl, decyl or phenyl; and

[0047] R1 and R2 are advantageously selected from isopropyl, isobutyl, sec-butyl, tert-butyl, isopentyl, 2-methylbutyl, 2-ethylpropyl, cyclobutyl, cyclopentyl, cyclohexyl, cycloheptyl, cyclooctyl, or when the salt water to be treated has a calcium concentration of less than 10 g / L and / or when Li + / Na + Preferably, R1 and R2, together with the nitrogen atom carrying them, form a pyrrolidine, piperidine, azetidine or azoalkane ring.

[0048] The lithium cation extractant compound may be selected from but is not limited to:

[0049]

[0050]

[0051]

[0052]

[0053] In particular, the second organic compound for extracting the cation of the salt to be extracted can be selected from 4-tert-butylcalix[4]arene acid tetraethyl ester, 4-tert-butyl-calix[6]arene acid hexaethyl ester, 2-[2,2-bis[[2-(dicyclohexylamino)-2-oxoethoxy]methyl]butoxy]-N,N-dicyclohexylacetamide and N,N,N',n',N",N"-hexacyclohexyl-4,4',4"-propyltri(3-oxobutyramide).

[0054] The organic phase may also contain a hydrophobic polar organic diluent. Preferably, this is an aromatic diluent, which may be selected from polar aromatic compounds such as 1-chloro-2-bromobenzene, 1,2-dibromobenzene, 2-bromotoluene or 3,4-dibromotoluene.

[0055] 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 the solvating molecules of the complementary anionic species to the extracting molecules of the cationic species is greater than 1.

[0056] The second organic compound which extracts the cation of the salt to be extracted may be present in the liquid hydrophobic organic phase in a concentration of at least 0.1 mol / L.

[0057] Before recovering the salty water from which the salts to be extracted have been removed, said salty water may be sent to another extractor agitator-mixer and then to an extractor decanter separator, which receives the regenerated organic phase in order to obtain, after decanting, a salty water with less salts to be extracted, which is recovered or sent again to an extractor agitator-mixer and then to an extractor decanter separator to obtain a salty water with less salts to be extracted than before. In this case, said operation of extracting the salts to be extracted from the salty water may be repeated at least once, the organic phase loaded with the salts to be extracted being returned from a given decanter separator to the lower agitator-mixer in order to recover the organic phase loaded with salts at the outlet of the initial decanter separator, wherein the organic phase is sent to a so-called "organic" heat exchanger.

[0058] In particular, multiple extractor stirrer-mixer-decanter-separators can be operated in countercurrent. The aqueous phase entering the extractor stirrer mixer contacts the organic phase coming from the upper extractor decanter separator.

[0059] At least one further regeneration column may be provided, which is connected in parallel with the initial regeneration column, which is supplied by part of the flow of a so-called “organic” heat exchanger.

[0060] In particular, the regeneration tower may be a stirred tower or a pulse tower.

[0061] The water used as regeneration water can be fed into a so-called "water" heat exchanger in which it is heated by water containing the salt to be extracted leaving the regeneration tower or towers (if several are provided) to obtain heated regeneration water, which can be further heated before being introduced into the regeneration tower or towers and from which cooling water containing the salt to be extracted is obtained.

[0062] Before being sent to the initial extraction agitator mixer, the wash water loaded with impurities can be sent to a purification unit to obtain:

[0063] ο the retentate or concentrate or raffinate brine fed to the initial extraction agitator mixer; and

[0064] o Return purified water to the initial wash agitator mixer.

[0065] In particular, before the washed organic phase leaving the so-called wash decanter separator is sent to a so-called "organic" heat exchanger, it is sent to another so-called wash agitator mixer and then a so-called wash decanter centrifugal separator receiving wash water, so as to obtain, after decanting, an organic phase subjected to a further washing step, which is sent to the so-called "organic" heat exchanger, said operation of washing the organic phase being repeated at least once, the wash water or retentate or concentrate or raffinate brine loaded with impurities being returned to the initial wash agitator mixer after purification.

[0066] In particular, a plurality of agitators-mixers-wash decanters-centrifuges can be operated in countercurrent. The aqueous phase entering the wash agitator mixer contacts the organic phase coming from the upper decanter centrifuge.

[0067] The water from the regeneration tower or the so-called "water" heat exchanger, loaded with the salt to be extracted, can be sent to a purification unit to obtain:

[0068] - water containing concentrated and purified salts to be extracted;

[0069] - Purified water, sent to the regeneration tower or so-called "water" heat exchanger.

[0070] The purification unit may be a reverse osmosis unit, or a membrane distillation unit, or an evaporation / condensation unit, or a unit for absorbing water by a temperature-regenerable organic phase, or a combination of at least two of these units.

[0071] At least one of the following:

[0072] - distilled water;

[0073] -freshwater;

[0074] - Desalinated water; and

[0075] - extracting at least some of the water containing the concentrated and purified salt from the purification unit (2f), possibly after cooling,

[0076] Can be used as washing water.

[0077] The initial decanter separator and / or the wash decanter-separator, if any, can be a decanting centrifugal separator allowing the removal of a phase consisting of solid waste or water wash from the organic phase, or the separation of a liquid two-phase medium with an aqueous phase / organic phase ratio of less than 1.0 for the initial decanter separator and less than 0.2 for the wash decanter separator.

[0078] The salt to be selectively extracted can be selected from sodium salts such as NaCl, NaCN, NaNO3, NaBr, NaI, potassium salts such as KCl, KCN, KNO3, KBr, KI, lithium salts such as LiCl, LiCN, LiNO3, LiBr, LiI, or other salts that can be selectively extracted in the presence of other salts.

[0079] The salts preferably extracted are so-called diionic salts, i.e. salts containing monoatomic cations carrying a single positive charge, such as Li + Cation, K + Cation or Na + Cations, and monoatomic or polyatomic anions carrying a single negative charge, such as the monoatomic anion Cl - or polyatomic anion NO3- of salt.

[0080] The volume ratio between the wash water introduced into the so-called wash agitator-mixer and the organic phase can be between 0.01 and 0.1, preferably between 0.03 and 0.06.

[0081] The volume ratio between the hot water introduced into the so-called regeneration column and the organic phase may be between 0.05 and 0.2, preferably between 0.08 and 0.15.

[0082] The following examples illustrate the invention without limiting its scope.

[0083] Example 1

[0084] Figure 1 The facility shown is intended for treating a saline or brackish water effluent S containing n dissolved salts consisting of cations and anions, from which salts can be extracted at least by a dedicated organic preparation effluent O circulating in a closed loop.

[0085] The ionic composition of the salt water or brackish water S, the ionic composition of the organic agent O and the ionic composition of the aqueous phase A vary throughout the process and are indicated by successive symbols S1 to S4, O1 to O6 and A1 and A2 respectively.

[0086] The desired product is in the form of a hot water phase P1 loaded with at least the salts to be extracted, which is recovered after cooling as production effluent P2.

[0087] Figure 1 Facilities shown include:

[0088] - three extractor stirrer mixers 1a, 2a, 3a; these extractor stirrer mixers each ensure a thorough mixing of the aqueous and organic phases supplied to them;

[0089] - three decanter phase separators 1d, 2d and 3d, respectively associated with the extractor stirrer mixers 1a, 2a, 3a; these decanter phase separators (the first of which may be of the centrifugal type, in this case denoted 1dc) allow the separation of the organic and aqueous phases of the effluent after the aqueous phase of the effluent has passed through the respective extractor stirrer mixers 1a, 2a, 3a;

[0090] - two heat exchangers 1e, 2e, each allowing heating and cooling respectively of the organic and aqueous phases supplied to them, the exchanger 1e being the main exchanger;

[0091] - two differential contactors 1 c , 2 c , which are here static or stirred or pulsed towers, each allowing the thermal desorption or deextraction of the organic phase supplied to it, at least of the salts to be extracted which were previously absorbed by the organic phase in the extractor stirrer mixers 1 a to 3 a ;

[0092] - Heating unit 1b, its function is as follows.

[0093] The salt water or brine S1 to be treated is mixed with the organic phase O3 in the stirred reactor 1a to produce an organic phase dispersed in a continuous aqueous phase (or vice versa). The two-phase mixture is then transferred to a centrifugal decanter 1dc to separate the two liquid phases and produce an effluent S2 sent to the stirred reactor 2a and an effluent O4 sent to the main heat exchanger 1e. It is also possible to remove the organic phase O4 before sending the effluent O4 to the heat exchanger 1e. Figure 1 ), or removing a third phase consisting of waste, particularly solid waste.

[0094] The salt water or brine S2 is then mixed with the organic phase O2 in the stirred reactor 2a to produce an organic phase dispersed in a continuous aqueous phase (or vice versa). The two-phase mixture is then transferred to a decanter 2d for gravity separation of the two liquid phases and produces an effluent S3 sent to the stirred reactor 3a and an effluent O3 sent to the stirred reactor 1a.

[0095] The brine or salt water S3 is then mixed with the organic phase O1 in the stirred reactor 3a to produce an organic phase dispersed in a continuous aqueous phase (or vice versa). The two-phase mixture is then transferred to a decanter 3d for gravity separation of the two liquid phases and produces a treated effluent S4 (raffinate) and an effluent O2 that is fed to the stirred reactor 2a.

[0096] The organic phase O4 loaded with at least the salt to be extracted is then pumped to the main heat exchanger 1e where it is reheated to obtain a hot effluent O5 loaded with the salt to be extracted, which is then sent to the top of the columns 1c and 2c operated in parallel (it is considered here that the density of O5 is greater than that of P1) so as to produce two organic effluents at the bottom of these columns, which are then combined to form O6, a hot organic phase free of salts, regenerated and pumped to the main heat exchanger 1e for cooling and then recovered in the form of organic effluent O1.

[0097] The aqueous phase A1, generally fresh, desalinated or deionized water, is pumped to the secondary heat exchanger 2e where it is heated to obtain a hot effluent A2 which is then fed to the bottom injection of the columns 1c and 2c operated in parallel for producing at the top of the two columns two aqueous effluents containing at least the salts to be extracted, which are then combined to form the product P1, a hot water phase containing at least the salts to be extracted and which is pumped to the secondary heat exchanger 2e for cooling and then recovered as production effluent P2.

[0098] The heating unit 1 b compensates for the heat losses caused by the temperature difference of the outflows at the cold ends of the heat exchangers 1 e and 2 e .

[0099] The towers 1c and 2c are preferably used with a continuously descending organic phase and an ascending dispersed aqueous phase, with a high organic / aqueous (O / A) flow rate ratio in the regeneration (hot water desorption), which is greater than 5, preferably greater than 10, possibly greater than 15 or even greater than 20. This makes it possible to minimize the consumption of fresh water for desorption of the salts to be extracted, while increasing the reconcentration of the salts 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 a higher O / A.

[0100] The aqueous effluent P1 does not see any transfer of organic phase droplets, just as the tops of towers 1c and 2c do not see any gaseous organic headspace, due to the system in which the aqueous phase is installed at the top of these towers, which allows any organic phase to coalesce, where it is associated with an appropriate residence time for the precipitation of organic matter. It should be noted that this design is relevant even if the O / A flow ratio during regeneration is high, due to the absence of an organic gas phase at the top of the tower.

[0101] It is also conceivable that organic traces can be removed from the aqueous effluents S4, P1 or P2 by installing an oil-water separation unit, which can be a decanter and / or a coalescer, combined with an adsorption unit using an adsorbent, which can be activated carbon, silica gel, diatomaceous earth and / or another similar method, allowing the organic components to be recycled to the process and / or transported to a disposal plant by incineration together with the adsorption medium.

[0102] This cycle is called 3-0-Z, 3 being the number of unit operations for extracting the salt to be extracted, 0 being the number of washing operations, and Z being the theoretical number of stages for regenerating the desorption with hot water.

[0103] Example 2

[0104] Figure 2 The facilities shown are Figure 1 The same is shown except that the decanter separators 2d and 3d have been replaced by centrifugal decanters 2dc and 3dc, and the reheating unit 1b used to compensate for heat losses at the cold end of exchangers 1e and 2e has been placed on the organic effluent O6 instead of the aqueous effluent A2.

[0105] The salt water or brine S used in this example is difficult to separate by gravity separation, so centrifugal separation is required. In addition, the economic conditions of the process may favor the use of a decanting centrifugal separator rather than a gravity decanter separator.

[0106] Example 3

[0107] Figure 3The facility shown is intended for treating a saline or brackish water effluent S containing n dissolved salts consisting of cations and anions, from which salts can be extracted at least by a dedicated organic preparation effluent O circulating in a closed loop.

[0108] The ionic composition of the salt water or brackish water S and the ionic composition of the organic agent O vary throughout the process and are indicated by successive symbols S1 to S6 and O1 to O8, respectively.

[0109] The desired product is in the form of a hot water phase P1 loaded with at least the salts to be extracted, which is recovered as production effluent P2 after cooling and as production effluent P3 after reconcentration.

[0110] Figure 3 Facilities shown include:

[0111] - five agitator mixers 1a, 2a, 3a, 4a, 5a; each of these agitator mixers ensures a thorough mixing of the aqueous and organic phases supplied to them, agitator mixers 1a, 2a, 3a and 4a being extraction agitator mixers, agitator mixer 5a being a washing agitator mixer;

[0112] - five decanter phase separators 1d, 2d, 3d, 4d and 5dc, respectively connected to the stirrer mixers 1a, 2a, 3a, 4a and 5a, the last of these decanter phase separators being of centrifugal type, indicated as 5dc, allowing the separation of the organic and aqueous phases of the effluent after the aqueous phase of the effluent has passed through the respective stirrer mixers 1a, 2a, 3a, 4a and 5a;

[0113] - two heat exchangers 1e, 2e allowing heating and cooling respectively of the organic and aqueous phases supplied to them, the exchanger 1e being the main exchanger;

[0114] - two differential contactors 1 c , 2 c , which are here static or stirred or pulsed towers, each allowing the organic phase supplied to it to be thermally desorbed or deextracted, at least the salts to be extracted which were previously absorbed by the organic phase in the stirrer mixers 1 a to 4 a ;

[0115] - Heating unit 1b, its function is as follows;

[0116] - Two units for reverse osmosis and / or membrane distillation and / or evaporation / condensation and / or absorption of water by a temperature-regenerable organic phase 1f, 2f, allowing separation of the retentate or concentrate or raffinate and the permeate or condensate or strip.

[0117] The brine or salt water S1 to be treated and the aqueous retentate / concentrate / raffinate S8 are mixed with the organic phase O4 in the stirred reactor 1a to produce a dispersed organic phase in a continuous aqueous phase (or vice versa). The two-phase mixture is then transferred to a decanter 1d to separate the two liquid phases and produce an effluent S3 sent to the stirred reactor 2a and an effluent O5 sent to the stirred reactor 5a.

[0118] The salt water or brine S3 is then mixed with the organic phase O3 in the stirred reactor 2a to produce an organic phase dispersed in a continuous aqueous phase (or vice versa). The two-phase mixture is then transferred to a decanter 2d to separate the two liquid phases and produce an effluent S4 that is sent to the stirred reactor 3a and an effluent O4 that is sent to the stirred reactor 1a.

[0119] The salt water or brine S4 is then mixed with the organic phase O2 in the stirred reactor 3a to produce an organic phase dispersed in a continuous aqueous phase (or vice versa). The two-phase mixture is then transferred to a decanter 3d to separate the two liquid phases and produce an effluent S5 that is sent to the stirred reactor 4a and an effluent O3 that is sent to the stirred reactor 2a.

[0120] The salt water or brine S5 is then mixed with the organic phase O1 in the stirred reactor 4a to produce an organic phase dispersed in a continuous aqueous phase (or vice versa). The two-phase mixture is then transferred to a decanter 4d to separate the two liquid phases and produce a treated effluent S6 (raffinate) and an effluent O2 that is sent to the stirred reactor 3a.

[0121] Then, in the stirred reactor 5a, the organic phase O5, containing at least the salt to be extracted, is mixed 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 a centrifugal decanter 5dc to separate the two liquid phases and produce an effluent S7 sent to a water recovery unit 1f, such as a reverse osmosis unit, a membrane distillation unit and / or an evaporation / condensation unit and / or a unit for absorbing water by a temperature-regenerable organic phase, and to produce an effluent O6 sent to the main heat exchanger 1e. The purpose of this step is to wash the organic phase O5 with a very small water flow A7 to remove salt impurities, thereby improving the purity of the salt to be extracted contained in the organic effluent O6. The centrifugal decanter 5dc can also be used to remove aqueous washes from the organic phase O5 before sending the effluent O6 to the heat exchanger 1e and before sending the effluent S7 to the water recovery unit 1f, while removing organic washes from the aqueous phase S7, or to remove a third phase consisting of solid waste or other materials.

[0122] The aqueous phase A7 is usually composed of the aqueous phase A5, which is distilled water, fresh water or desalted water and recycled water A6, and its flow rate is relatively small compared to the organic flow, and the flow rate ratio A / O of water washing (washing with cold water of neutral pH, the pH value has no effect on the washing performance) is less than 0.25, preferably less than 0.1, and may be less than 0.05, or as low as 0.01. This makes it possible to increase the purity of the salt extracted from the production P1 while minimizing the consumption of fresh water for washing, while minimizing the loss of the salt to be extracted during the washing process and / or the loss of the total extraction rate of the salt to be extracted.

[0123] Then, the aqueous phase S7 is pumped from the temperature-regenerable organic phase 1f to a reverse osmosis unit and / or a membrane distillation unit and / or an evaporation / condensation unit and / or a water absorption unit, and is separated into brine S8 (permeate or concentrate or residual liquid) which is sent to the stirred reactor 1a and an aqueous effluent A6 (permeate or condensate or de-extracted product) which is recycled to the inlet of the stirred reactor 5a.

[0124] The organic phase O6 containing the salt to be extracted is then pumped to the main heat exchanger 1e, where it is heated to obtain a hot organic effluent O7 containing the salt to be extracted, which is then sent to the top of the columns 1c and 2c operated in parallel for producing two organic effluents at the bottom of the columns, which are then combined to form effluent O8, i.e., the hot organic phase free of salt, which is regenerated and pumped to the main heat exchanger 1e for cooling and then recovered in the form of a cold organic effluent O1 (it is assumed here that the density of O7 is greater than the density of P1).

[0125] The aqueous phase A1, usually fresh water, desalinated water or deionized water, can be combined with the permeate condensate extract A2 from the water recovery unit 2f, the reverse osmosis unit, the membrane distillation unit and / or the evaporation / condensation unit and / or the unit for absorbing water by the temperature-regenerable organic phase, and pumped as aqueous effluent A3 to the secondary heat exchanger 2e, where it is heated to obtain a hot water effluent A4, which is then sent to be injected at the bottom of the columns 1c and 2c, operating in parallel, producing two aqueous effluents at the top of the two columns. The effluents, which contain at least the salt to be extracted, are then combined to form product P1, i.e. a hot water phase containing at least the salt to be extracted, and are pumped to a secondary heat exchanger 2e for cooling and then recovered as production effluent P2, which can be reconcentrated by a water recovery unit 2f, a reverse osmosis unit, a membrane distillation unit and / or an evaporation / condensation unit and / or a unit for absorbing water by a temperature-regenerable organic phase to produce a permeate condensate de-extracted A2 and a concentrated production effluent P3.

[0126] The heating unit 1 b compensates for the heat losses caused by the temperature difference of the outflows at the cold ends of the heat exchangers 1 e and 2 e .

[0127] The towers 1c and 2c are preferably used with a continuously descending organic phase and an ascending dispersed aqueous phase, with a high O / A flow ratio in the regeneration (hot water desorption), which is greater than 5, preferably greater than 10, possibly greater than 15 or even greater than 20. This makes it possible to minimize the consumption of fresh water for desorption of the salts to be extracted, while increasing the reconcentration of the salts to be extracted between the water to be treated and the regeneration or desorption water to be produced. The higher the regeneration temperature of the organic phase in 1c and 2c, the higher the O / A ratio.

[0128] The aqueous effluent P1 does not see any transfer of organic phase droplets, just as the tops of towers 1c and 2c do not see any gaseous organic sky, due to the installation of a system at the top of these towers that allows the liquid phase of any organic phase to coalesce, associated with an appropriate residence time for the precipitation of organic substances. It should be noted that this design is relevant even if the flow ratio O / A during regeneration is high, due to the absence of an organic gas phase at the top of the tower.

[0129] The devices 5a and 5dc can be combined to form a single device commonly referred to as a centrifugal extractor. The device can be doubled, tripled or quadrupled by connecting in series (preferably countercurrently) to improve the washing of O5, thereby producing a higher purity salt to be extracted as effluent P1. The device can be multiplied and used in parallel to be able to process an increased flow of O5 effluent.

[0130] It is also considered that organic traces can be removed from the aqueous effluents S6, S7, P1, P2 or P3 by installing an oil-water separation unit, which can be a decanter and / or a coalescer combined with an adsorption unit, which uses an adsorbent, which can be activated carbon, silica gel, diatomaceous earth and / or another similar method, allowing the organic components to be recycled into the process and / or transported to a disposal facility by incineration together with the adsorption medium.

[0131] When other outflows ( Figure 3 Upper right corner) operates at a higher temperature, but below the boiling point of water and salt water effluents A4 and P1 at the operating pressures of 1c, 2c and 2e, all effluents to the left of heat exchanger 1e and below heat exchanger 2e are preferably operated at conditions close to ambient temperature (5 to 35°C).

[0132] This configuration alternative can be implemented when the brine or saltwater aqueous phase to be treated does not allow sufficient gravity settling to be economically feasible due to its density, viscosity and / or other inherent characteristics. In this case, it is preferred to use mixer centrifugal decanters 1dc, 2dc, 3dc and 4dc instead of extractor mixer gravity decanters 1d, 2d, 3d and 4d, or possibly a centrifugal extractor with increased internal mixing time compared to the prior art.

[0133] This cycle is called 4-1-Z, 4 being the number of unit operations for extracting the salt to be extracted, 1 being the number of washing operations, and Z being the theoretical number of stages for regenerating the desorbed organic phase using hot water.

[0134] Example 4

[0135] When processing brine or salt water, where the relative concentration of the salt to be extracted is low compared to other salts present, it may be necessary to use several cold wash / scrub stages in series to improve washing and removal of impurities from the organic phase.

[0136] therefore, Figure 4 The facilities shown are Figure 3 The same as shown, with the addition of two further agitator mixers 6a and 7a, and decanting centrifugal separators 6dc and 7dc.

[0137] Then, in the stirred reactor 7a, the organic phase O5 containing the salt to be extracted is mixed with the salt water S10 to produce an aqueous phase dispersed in the continuous organic phase (or vice versa). The two-phase mixture is then transferred to a centrifugal decanter 7dc to separate the two liquid phases and the resulting effluent S7 is sent to a water recovery unit 1f, such as a reverse osmosis unit, a membrane distillation unit and / or an evaporation / condensation unit and / or a unit for absorbing water by a temperature-regenerable organic phase, and the resulting effluent O6 is sent to the agitator mixer 6a.

[0138] The aqueous phase S7 is then pumped to a reverse osmosis unit, a membrane distillation unit and / or an evaporation / condensation unit and / or a water absorption unit via the temperature-regenerable organic phase 1f, and is separated into brine S8 (permeate or concentrate or residual liquid) which is sent to the stirred reactor 1a and an aqueous effluent A6 (permeate or condensate or de-extracted product) which is recycled to the inlet of the stirred reactor 5a.

[0139] Then, the organic phase O6 containing the salt to be extracted is mixed in the stirred reactor 6a to produce an aqueous phase dispersed in the continuous organic phase (or vice versa). The two-phase mixture is then transferred to the centrifugal decanter 6dc to separate the two liquid phases and produce an effluent S10 sent to the stirred reactor 7a and an effluent O7 sent to the agitator mixer 5a.

[0140] Then, the organic phase O7 containing the salt to be extracted is mixed in the stirred reactor 5a to produce an aqueous phase dispersed in the continuous organic phase (or vice versa). The two-phase mixture is then transferred to a centrifugal decanter 5dc to separate the two liquid phases and produce an effluent S9 sent to the stirred reactor 6a and an effluent O8 sent to the heat exchanger 1e.

[0141] Adding one, two or three agitator mixers for cold washing of the organic phase can increase the purity and selectivity of the salts to be extracted.

[0142] This cycle is called 4-3-Z, 4 being the number of unit operations for extracting the salt to be extracted, 3 being the number of washing operations, and Z being the theoretical number of stages for regenerating the desorbed organic phase using hot water.

[0143] Figure 5 Facilities and Figure 4 Same as in, where the water recovery unit 2f is moved from the cold zone to the hot zone of the secondary "water" heat exchanger 2e, the water recovery unit 1f is removed.

[0144] The water recovery unit 2f can then be a membrane distillation unit and / or an evaporation / condensation unit and / or a unit using a temperature-regenerable organic phase to absorb water.

[0145] Example 5

[0146] In this example, the salt to be selectively extracted is sodium chloride NaCl.

[0147] The organic phase O comprises:

[0148] - As Na + Ion extractant, sodium ion carrier X, also known as 4-tert-butyl-calix[4]arene acid tetraethyl ester (C 60 H 80 O 12 , CAS n°: 97600-39-0), concentration 0.275M (mol / L), Na in methanol at 25°C + The ion pair number K is 5;

[0149] - As an anionic solvating agent, N-(3,5-dichlorophenyl)-octanamide (C 14 H 19 Cl2NO, CAS No.: 20398-46-3), with a concentration of 0.825 M and a pKa of 13.83 + / - 0.70 in water; and

[0150] -1-Chloro-2-bromo-benzene as diluent.

[0151] At 20°C, the density of the organic phase was 1.48 g / L.

[0152] The salt water to be treated includes sodium chloride and potassium chloride to be extracted. The salinity of the salt water is 209.81 g / L, of which Na + 2.26 g / L, K + It is 107g / L.

[0153] The aqueous solution of salt water (O / A) with a volume ratio of salt water to organic phase of 0.05, 0.2, 0.5, 1 and 4 was prepared at room temperature (20°C). 提取contact, a high relative extraction of sodium in the form of NaCl relative to potassium or KCl was obtained, such as Figure 6 shown.

[0154] like Figure 7 As shown, this sodium extraction is selective and does not alter potassium concentration. 提取 The potassium concentration remained stable at about 107 g / L in the tests with a ratio of 0.05, 0.2, 0.5 and 1. 提取 The residual sodium concentration in the brine treated with a ratio of 1 was 168 mg / L.

[0155] Figure 8 The ion concentration of each compound in the organic phase after extraction is shown. 提取 When the ratio is less than 2, NaCl occupies the majority in the organic phase. 提取 When it is lower, NaCl is even more.

[0156] Then, in order to prove that KCl is poorly retained in the organic phase, the (O / A) 提取 Preparations loaded with Na, K and Cl were considered after extraction followed by water wash at room temperature (20°C) in which very small amounts of water accounted for 1% to 20% of the volume of the organic phase, also known as the A / O wash ratio.

[0157] The results are as follows Fig. 9 As shown. It is clear that once the organic phase is in contact with 1% volume of water, 80.8% of the KCl absorbed in the organic phase is returned to the water, while this proportion of NaCl is only 1.9%. Similarly, in a washing step, when the water volume is 4%, 94.8% of KCl is desorbed / transferred to water, while the desorption / transfer rate of NaCl is 6.4%.

[0158] It can also be seen that NaCl is difficult to desorb at room temperature, and when the volume of water is 20% of the organic phase, only 21.1% of NaCl is transferred to water.

[0159] In order to allow NaCl to enter the aqueous phase and regenerate the organic phase, the preparation obtained by washing with 4% volume of water was thermally regenerated at 80°C by contacting with 2%, 5%, 10%, 20% and 50% hot water by volume of the organic phase. This ratio is recorded as (A / O) 再生 .

[0160] The results are as follows Fig.10 It is evident that at 80°C, when the water content is 20%, 67.3% of NaCl is desorbed from the organic phase (compared to 21.1% previously at 20°C).

[0161] Fig.11The ion concentration of the brine obtained after regeneration at 80°C is shown. After a washing stage with an A / O washing ratio of 4%, the ion concentration of the brine obtained after regeneration at 80°C is shown. 再生 The brine obtained in the thermal regeneration of the catalyst consists of 93 to 99% by mass of NaCl, which enables the purification of NaCl with minimal NaCl losses.

[0162] Parameter studies were performed on cycles 4-1-3 (the cycle of Example 3 with three theoretical stages for regenerating the desorption with hot water) and 4-2-3 (with two cold water wash operations and three theoretical stages for regenerating the desorption with cold water).

[0163] The result of loop 4-1-3 is as follows Fig.12 shown.

[0164] By extending the curve to the left, i.e., in the case where the A / O wash ratio tends to 0, we will have a cycle performance of 4-0-3 without water washing. It is obvious here that this washing by increasing the A / O wash ratio can increase the purity of the NaCl produced at the outlet of the regeneration tower to a high level of about 94 mass NaCl (excluding water). In addition, for sufficiently high (O / A) 提取 The NaCl extraction rate at 20°C can be maintained above 90%, while the purity of the extracted NaCl is increased by washing with water at room temperature (here 20°C). For this cycle 4-1-3, for a NaCl purity of 91.8 mass, the selected filling point gives a NaCl extraction rate of 93.2%, (O / A) 提取 The ratio is 1.1, (A / O) 洗涤 The ratio is 5% (circled points).

[0165] Fig.13 The effect of adding a second countercurrent washing stage at room temperature is shown. The performance of the process can be further improved by switching from cycle 4-1-3 to cycle 4-2-3, adding a second washing stage at room temperature, in order to obtain a high sodium extraction yield and a higher purity of NaCl produced (high level of 98% mass NaCl) at the outlet of the regeneration stage (removal).

[0166] For this 4-2-3 cycle, the packing point selected gave a NaCl extraction yield of 92.8% for a NaCl purity of 97.4 mass, an O / A extraction ratio of 1.1 and an A / O wash ratio of 5% (circled point at the intersection of the two curves).

[0167] Table 1 shows the material balance of cycle 4-1-3, where (O / A) 提取 The ratio is 1.1, (A / O) 洗涤 The ratio is 5%, (A / 0) 再生 The ratio is 10%.

[0168] [Table 1]

[0169]

[0170] The table shows that the raffinate S6 now contains only 0.13% by mass of NaCl to obtain a KCl and K2SO4 purity of 99.87% by mass (excluding water), where K + The amount of ions is almost unchanged compared to the amount present in the salt water S1 to be treated.

[0171] Table 2 shows the material balance of cycle 4-2-3, where (O / A) 提取 The ratio is 1.1, (A / O) 洗涤 The ratio is 5%, (A / 0) 再生 The ratio is 10%.

[0172] [Table 2]

[0173]

[0174] Compared with Table 1, it can be seen that the P3 concentrate obtained with two washing stages contains more NaCl (97.34% by mass) than the P3 concentrate obtained with one washing stage (91.81% by mass).

[0175] Example 6

[0176] In this example, the salt to be selectively extracted is potassium chloride, KCl.

[0177] The organic phase O comprises:

[0178] -As K + The extractant of ions, 4-tert-butyl-calix[6]arene acid hexaethyl ester (C 90 H 120 O 18 , CAS n°: 92003-62-8), the concentration is 0.175M (mol / L), and its K + The ion pair number K is 4.8;

[0179] - As an anionic solvating agent, N-(3,4-dichlorophenyl)-octanamide (C 14 H 19 Cl2NO, CAS No: 730-25-6), with a concentration of 0.525 M and a pKa of 13.63 + / - 0.70 in water; and

[0180] -1-Chloro-2-bromo-benzene as diluent.

[0181] At 20°C, the density of the organic phase was 1.43 g / L.

[0182] The salt water to be treated includes potassium chloride and sodium chloride to be extracted. The salinity of the salt water is 168g / L, of which Na + 62.7 g / L, K + It is 4.53g / L.

[0183] By mixing the salt water with the salt water aqueous solution (O / A) at room temperature (20°C) 提取 The volume ratios of the organic phase used were 0.1; 0.4; 0.8; 1.5 and 4, respectively, and good relative extraction of potassium was obtained in the form of KCl compared to sodium or NaCl, such as Fig.14 shown.

[0184] like Fig.15 As shown, potassium extraction is selective without changing the sodium concentration. 提取 The ratio is 0.1, 0.4, 0.8 and 1.5, using (O / A) 提取 In the salt water with a potassium concentration of 2.8 g / L treated with a ratio of 1.5, the sodium concentration was stabilized at around 62 g / L.

[0185] Fig.16 The ion concentrations of various compounds in the organic phase after extraction at 20°C are shown. 提取 When the ratio is less than 2, KCl occupies the majority in the organic phase.

[0186] Then, in order to prove that the content of NaCl in the organic phase is very low, the (O / A) 提取 After extraction and washing with a very small amount of water ranging from 1% to 20% of the volume of the organic phase at room temperature (20°C), a preparation loaded with Na, K and Cl is considered, which is also called (A / O) 洗涤 Compare.

[0187] The results are as follows Fig.17 It is obvious that once the organic phase comes into contact with 1% volume of water, 48.3% of the NaCl absorbed in the organic phase will return to the water, while only 2.1% of the KCl will be absorbed.

[0188] Likewise, in one wash stage, when the water volume was 4%, 76.4% of the NaCl was desorbed / transferred to water, while KCl was 4.9% (10% water desorbed 90.3% of the NaCl and KCl desorbed 8.9%).

[0189] It can also be seen that KCl is difficult to desorb at room temperature, and when the volume of water is 20% of the organic phase, only 14.8% of KCl is transferred to water.

[0190] In order to allow KCl to enter the aqueous phase and regenerate the organic phase, the preparation obtained by washing with 4% volume of water was thermally regenerated at 80°C by contacting with 2%, 5%, 10%, 20% and 50% hot water by volume of the organic phase. This ratio is recorded as (A / O) 再生 .

[0191] Fig.18 The results obtained are shown, showing that, in 20% by volume of water, 83.7% of the KCl was desorbed from the organic phase (previously 14.8% at 20° C.).

[0192] Fig.19 Shows the saline ion concentration obtained after thermal regeneration at 80°C. 洗涤 After a washing stage with a ratio of 4%, the brine obtained from the thermal regeneration consists of 78 to 90% KCl by mass, which enables the KCl to be purified and KCl for upgrading to be produced.

[0193] Parameter studies were performed on these cycles 4-1-3 (cycle of Example 3 with three theoretical stages of hot water regeneration desorption) and 4-2-3 (cycle with two cold water wash operations and three theoretical stages of hot water regeneration desorption).

[0194] The result of loop 4-1-3 is as follows Fig. 20 shown.

[0195] By extending the curve to the left, that is, at (A / O) 洗涤 As the ratio approaches 0, we will have a cycle performance of 4-0-3 without water washing. It is obvious here that by increasing (A / O) 洗涤 This washing can increase the purity of KCl produced at the outlet of the regeneration tower to a high level of about 91% mass KCl (excluding water). In addition, for sufficiently high (O / A) 提取 The KCl extraction rate can be maintained above 90% while the purity of the extracted KCl is increased by washing with water at room temperature (here 20°C). For this cycle 4-1-3, for a KCl purity of 78.5% by mass, the selected filling point gives a KCl extraction rate of 85.0%, (O / A) 提取 The ratio is 2.2, (A / O) 洗涤 The ratio is 5% (circled points).

[0196] Fig.21 The effect of adding a second countercurrent washing stage at room temperature is shown. The performance of the process can be further improved by switching from cycle 4-1-3 to cycle 4-2-3, adding a second washing stage at room temperature, resulting in high potassium extraction rates and even higher purity KCl produced (high levels of 98% mass KCl) at the outlet of the regeneration stage (removal).

[0197] For this cycle 4-2-3, when the KCl purity is 91.5% by mass, (O / A) 提取 The ratio is 2.2, (A / O) 洗涤 At a ratio of 5%, the selected filler point gave a KCl extraction of 85.6% (circled point).

[0198] Table 3 shows the material balance of cycle 4-1-3, where (O / A) 提取 The ratio is 2.2, (A / O) 洗涤 The ratio is 5%, (A / 0) 再生 The ratio is 10%.

[0199] [Table 3]

[0200]

[0201] The table shows that the raffinate S6 now contains only 0.76% by mass of KCl, of which Na + The amount of ions is almost unchanged compared with that in the salt water S1 to be treated.

[0202] Table 4 shows the material balance of cycle 4-2-3, where (O / A) 提取 The ratio is 2.2, (A / O) 洗涤 The ratio is 5%, (A / 0) 再生 The ratio is 10%.

[0203] [Table 4]

[0204]

[0205] Comparison with Table 3 shows that the concentrate P3 obtained with two washing stages contains more KCl (91.46% by mass) than the concentrate P3 obtained with a single washing stage (78.46% by mass).

[0206] Examples 5 and 6 demonstrate that the organic selective extraction preparations and processes used in accordance with the present invention are capable of obtaining relatively pure effluents in NaCl and / or KCl.

[0207] Example 7

[0208] In this example, the salt to be selectively extracted is lithium chloride, LiCl.

[0209] The organic phase O comprises:

[0210] - As Li ion + Extractant, lithium ion carrier VIII, also known as N, N, N', N', N", N'-hexacyclohexyl-4, 4', 4"-propynyl tris (3-oxobutyric acid amide) (C48 H 83 N3O6, CAS No: 1133338-85-9), the concentration is 0.25M (mol / L), and the logarithm K in methanol is 2.2;

[0211] - As an anionic solvating agent, N-(3,5-dichlorophenyl)-octanamide (C 14 H 19 Cl2NO, CAS No.: 20198-46-3), with a concentration of 0.75M and a pKa of 13.83+ / -0.70 in water; and

[0212] -1-Chloro-2-bromo-benzene as diluent.

[0213] At 20°C, the density of the organic phase was 1.43 g / L.

[0214] The brine to be treated contains lithium chloride. + 、Na + , K + Mg 2+ , Ca 2+ , boron, SO4 2- and Cl - The brine has a salinity of 209.81 g / L, of which Li + 2.26g / L, Na + 2.26mg / L, K + 107g / L.

[0215] Since most lithium-rich brines have a Ca / Li molar ratio of less than 0.5 (exceptions are: Maricunga, Chile: Ca / Li = 2; Tres-Quebradas, Argentina: Ca / Li = 7.4), and the Na / Li molar ratio can vary between 1 and 150, the key to direct lithium extraction is obviously a very good selectivity between lithium and sodium.

[0216] As an example, the average brine composition of the Maricunga Blanco Project was examined, with representative Na / Li and Ca / Li ratios of 22.8 and 2.1, respectively. Table 5 shows the minimum, average, and maximum brine concentrations for the Maricunga Blanco Project.

[0217] [Table 5]

[0218]

[0219] Laboratory tests combined the reconstruction of the ion balance for cycles 4-0-3 (in the case of Example 1), 4-1-3 (in the case of Example 3) and 4-2-3, all of which were carried out at 20°C for lithium extraction and washing and at 80°C for desorption, as Fig. 22 and 23 As shown, the lithium yield can be as high as 97%, and the purity of the produced LiCl depends on three complementary parameters, namely (O / A) 提取 Ratio, Cycle Selection and (A / O) 洗涤 Compare.

[0220] The purity of the LiCl product (in % mass, excluding water) varies from a maximum of 65% for cycle 4-0-3 to 90% for cycle 4-1-3 and 97.5% for cycle 4-2-3.

[0221] From these two figures, we can see that increasing (O / A) 提取 The lithium yield can be increased by the co-extraction of NaCl and CaCl2, while the purity of the LiCl product can be reduced.

[0222] Surprisingly, when pure water is used for the wash stage (for one or two wash stages in series), the (A / O) of cycle 4-1-3 洗涤 When the A / O wash ratio of cycle 4-2-3 is kept below 10% (0.100) and below 5% (0.050), the purity of the LiCl product is greatly improved and the loss of lithium yield is very limited.

[0223] Fig.23 Table 6 shows the great advantage of washing with water to improve the purity of LiCl product. A salt water component is given at the outlet of the desorption column (P1).

[0224] [Table 6]

[0225]

[0226] Sodium and calcium salts are removed from the LiCl product while the lithium remains unchanged. The application of this process provides a high level of protection for local water resources, both for the water contained in the lithium-rich brine (if the lithium content of the lithium-rich brine to be treated is greater than 100 mg / kg water, preferably greater than 150 mg / kg water, the water in this brine will not be evaporated for re-concentration purposes), and for the fresh water required to operate the washing and regeneration desorption stages, because the high purity of the LiCl product and the low extraction rate of NaCl and CaCl2 allow efficient recovery of water by using a reverse osmosis membrane unit, a membrane distillation unit, and / or a multiple-effect evaporation condensation unit, and / or a unit for regenerating the organic phase by temperature for absorbing water, and / or a combination of at least two of these units on the output stream of the washing and regeneration desorption stages.

[0227] Tables 7 and 8 give the detailed mass balances of cycles 4-1-3 and 4-2-3 when the extraction and washing operations are performed at 20°C and the organic phase is regenerated in the desorber at 80°C. In these two cycles, the fresh water recovery is 84% ​​and 87%, respectively, resulting in very low fresh water requirements of 6.03 tons of water / ton of LiCl and 5.61 tons of water / ton of LiCl, respectively (37.7 tons and 43.2 tons of water / ton of LiCl if no fresh water recovery is performed). It should be noted that K + Mg 2+ and SO4 2- It does not appear in the product of effluent P3, nor does boron B, since 100% of it remains in the raffinate B6. Here, when water is removed, in effluent P3, the cycle 4-1-3 (O / A) 再生 =10, the purity of the produced LiCl is 85.9% by mass, cycle 4-2-3 (O / R) 再生 =10, the purity is 95.5% by mass.

[0228] It should be noted that by 再生 Reducing the regeneration water flow rate from 10 to 15 can produce LiCl with approximately the same purity as the produced LiCl, but will reduce the lithium extraction yield by about 10 percentage points. On the other hand, the produced LiCl is more concentrated, with a gain of up to 50% (61 g LiCl / kg water compared to 40 g LiCl / kg water).

[0229] [Table 7]

[0230]

[0231] [Table 8]

[0232]

[0233] Apart from Fig.25 In addition to the parametric study showing the possible performance of cycle 4-3-3, Tables 9 and Fig.24 The optimal dimensions adopted in the case of selective lithium extraction from Maricunga brine are also summarized for cycles with four extraction stages and three theoretical regeneration stages, respectively, of 0, 1, 2, and 3 washing stages.

[0234] It is clear that the selectivity of the organic agent chosen allows the LiCl content in the brine produced at the outlet of the regeneration process (removal) to increase from 2.26% to 51.6% by mass.

[0235] It is clear that using more and more countercurrent washing stages it is possible to complete the purification of the produced LiCl, rising from 51.6% mass without washing to 99.1% with 3 washing stages (washing), keeping the total lithium extraction yield above 90%.

[0236] [Table 9]

[0237]

[0238] The total salinity (TDS) obviously decreases with increasing LiCl purity (up to 99.09% dry weight), which also makes it possible to reduce the osmotic pressure of this effluent to 47 bar, making it easier to separate the LiCl from the water for recovery.

[0239] Compound formula:

[0240]

[0241] The target compound was synthesized in three consecutive stages.

[0242] Stage 1: Synthesis of secondary amines

[0243]

[0244] Place ketone (10 mmol, 1 eq), solvent (17 Vol), amine (45 mmol, 4.5 eq) and reagents in a clean, dry flask. Heat if necessary, depending on the reagent. After conversion, perform TLC and the initial ketone disappears. Evaporate the solvent (sometimes residual amine) under reduced pressure. Filter on Fontainebleau sand if necessary, then add methanol (12 volumes). If necessary, add NaBH4 (30 mmol, 3 eq) in batches and stir at room temperature for 1-15 hours.

[0245] Purify on silica gel if necessary. (DCM to DCM / ethyl acetate). Yield: 20-77%.

[0246] Stage 2: Synthesis of Chloroacetamide

[0247]

[0248] The previously formed amine (10mmol, 1eq), dichloromethane (3Vol, 15eq) and triethylamine (30mmol, 3eq) were placed in a flask. Chloroacetyl chloride (20-25mmol, 2-2.5eq) was added cold under an argon atmosphere. Shake at room temperature for 5-24h. Two volumes of water were added and the aqueous phase was stripped twice with two volumes of DCM. The organic phase was concentrated on a rotary evaporator.

[0249] The product was purified on a silica gel column (100% DCM eluent). Yield: 30-70%.

[0250] Phase III: Synthesis of target compounds

[0251]

[0252] In the triol, NaH (3.5-4eq) is added to 10 volumes of anhydrous THF. The medium is heated to reflux under argon. Heat is introduced into the triol dissolved in 10 volumes of THF, and then the synthetic chloroacetamide dissolved in 15-20 volumes of THF is added. The medium is stirred and refluxed for 1-24h. The medium is neutralized by adding 10 volumes of water. The aqueous phase is back-extracted twice with 5 volumes of DCM. The organic phase is then washed once or twice with 5V of water. The organic phase is concentrated in a rotary evaporator.

[0253] Purify the crude product obtained by silica gel chromatography (eluent: heptane / ethyl acetate). The yield is generally between 50-70%.

[0254] Table 10 shows the formulas of the synthesized compounds, and Li VIII was renamed CE00.

[0255] [Table 10]

[0256]

[0257]

[0258]

[0259] Tables 11 and 12 show the characteristics of the synthesized compounds.

[0260] [Table 11]

[0261]

[0262] [Table 12] (Me = methyl)

[0263]

Claims

1. A method for selectively extracting a salt to be extracted from a saltwater or brackish water to be treated, wherein the salt to be extracted comprises at least one cation and at least one anion, and the saltwater or brackish water to be treated contains the salt to be extracted and salts other than the salt to be extracted, characterized in that: - The salt water to be treated (S1) is fed into a so-called initial extraction agitator mixer, wherein the initial extraction agitator mixer also introduces a liquid hydrophobic organic phase, said organic phase comprising: ° at least one first hydrophobic protic organic compound which solvates the anion of the salt to be extracted and has a pKa in water at 25°C of at least 9; and ° at least one second hydrophobic organic compound which selectively extracts the cations of the salt to be extracted and has a complex constant for the cations to be extracted whose logarithmic K value in methanol at 25° C. is greater than 1, preferably greater than 2, - mixing said salt water and said liquid hydrophobic organic phase in said extraction agitator mixer (1a) and sending the mixture to a so-called primary extraction decanter separator (1d; 1dc) to obtain: ° Saline water from which the salt to be extracted has been removed; and ° an organic phase containing the salt to be extracted; - recovery of brine from which the salts to be extracted have been removed; - The organic phase leaving the initial decanter separator (1d; 1dc) is sent to a new agitator mixer, called initial wash agitator mixer (5a; 7a), where it is mixed with wash water and the mixture is sent to a decanting centrifuge (5dc; 7dc), called initial wash decanting centrifuge, to obtain: o the purified organic phase containing the salts to be extracted is sent to a so-called "organic" heat exchanger for heating; and o Feeding the wash water loaded with impurities into the initial extraction agitator mixer (1a); - The organic phase containing the salts to be extracted, heated in the so-called "organic" heat exchanger (1e), is sent to the so-called initial regeneration tower (1c; 2c), where it is exchanged in countercurrent with hot water of a higher temperature than the initial salty water, in order to obtain: o water containing the salts to be extracted; and o Regenerated organic phase losing the salts to be extracted; - said organic phase containing the salts to be extracted is heated in a heat exchanger by the organic phase leaving the regeneration column (1c; 2c), - The regenerated organic phase is recycled to the initial extraction agitator mixer (1a).

2. The method according to claim 1, characterized in that Before recovering the salty water from which the salts to be extracted have been removed, the salty water can be sent to another extraction agitator mixer (2a) and then to an extraction decanter separator (2d), which receives the regenerated organic phase so as to obtain, after decanting, salty water with less salts to be extracted, which is recovered or sent again to an extraction agitator mixer (3a) and then to an extraction decanter separator (3d) to obtain salty water with less salts than the previous salts to be extracted. The operation of extracting the salts to be extracted from the salty water can be repeated at least once, the organic phase containing the salts to be extracted being returned from a given decanter separator to the lower-stage agitator mixer so as to recover the organic phase containing salts at the outlet of the initial decanter separator, wherein the organic phase is sent to a so-called "organic" heat exchanger.

3. The method according to claims 1 and 2, characterized in that At least one further regeneration column (2c) is provided, which is connected in parallel with the initial regeneration column (1c), which is supplied by part of the flow of a so-called "organic" heat exchanger (1e).

4. The method according to any one of claims 1 to 3, characterized in that The regeneration tower is a stirring tower or a pulse tower.

5. The method according to any one of claims 1 to 4, characterized in that The water used as regeneration water can be fed into a so-called "water" heat exchanger (2e) in which it is heated by water containing the salt to be extracted leaving the regeneration tower (1c; 2c) or leaving multiple regeneration towers if provided, to obtain heated regeneration water, which can be further heated before being introduced into the regeneration tower or multiple regeneration towers (1c; 2c) and cooling water containing the salt to be extracted is obtained from the heat exchanger (2e).

6. The method according to any one of claims 1 to 5, characterized in that Before being sent to the initial extraction agitator mixer (1a), the wash water loaded with impurities can be sent to a purification unit (1f) to obtain: o retentate or concentrate or raffinate brine which is sent to the initial extraction agitator mixer (1a); and o Send purified water back to the initial wash agitator mixer (5a; 7a).

7. The method according to claims 1 to 6, characterized in that Before the washed organic phase leaving the so-called wash decanter separator (7dc) is sent to a further so-called wash agitator mixer (6a), then a so-called wash decanter centrifugal separator (6dc) receiving wash water, in order to obtain, after decanting, an organic phase which has undergone a further washing step and which is sent to the so-called "organic" heat exchanger (1e), said operation of washing the organic phase can be repeated at least once, the wash water or retentate or concentrate or raffinate brine loaded with impurities being returned to the initial wash agitator mixer (7a) after purification.

8. The method according to any one of claims 1 to 7, characterized in that The water containing the salts to be extracted, coming from the regeneration tower (1c; 2c) or the so-called "water" heat exchanger (2e), can be sent to a purification unit (2f) to obtain: - water containing concentrated and purified salts to be extracted; - Purified water, sent to the regeneration tower (1c; 2c) or the so-called "water" heat exchanger (2e).

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

10. The method according to any one of claims 1 to 9, characterized in that At least one of the following: - distilled water; -freshwater; - desalinated water; and - At least some of the water containing the concentrated and purified salts is extracted from the purification unit (2f), possibly after cooling, and used as washing water.

11. The method according to any one of claims 1 to 10, characterized in that The initial decanter separator and / or the wash decanter-separator, when used, are decanting centrifugal separators (1dc, 5dc, 6dc, 7dc) which allow the removal of a phase consisting of solid waste or water wash from the organic phase, or the separation of a liquid two-phase medium with an aqueous phase / organic phase ratio of less than 1.0 for the initial decanter separator and less than 0.2 for the wash decanter separator.

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

13. The method according to any one of claims 1 to 12, characterized in that The salt to be selectively extracted can be selected from sodium salts such as NaCl, NaCN, NaNO3, NaBr, NaI, potassium salts such as KCl, KCN, KNO3, KBr, KI, lithium salts such as LiCl, LiCN, LiNO3, LiBr, LiI, or other salts that can be selectively extracted in the presence of other salts.

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

06.

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

Images