Electrochemical oxidation methods and processes

Abstract

The present invention relates to a method and system for producing an oxidant solution, comprising an electrochemical cell having an anode and a cathode. The method includes the steps of: (i) supplying a feed electrolyte to a reaction zone between the anode and the cathode, wherein the feed electrolyte contains sulfate ions (SO4 2- ) and divalent iron ions (Fe 2+ ) and (ii) in the operation cycle, the raw electrolyte is electrolyzed to produce peroxodisulfate (SO) 2- ) and trivalent iron ions (Fe 3+ and (iii) providing the oxidizing acid solution. The oxidizer solution is preferably used to leach metals from waste materials, including e-waste, and metal-bearing ores.

Description

Detailed Description of the Invention

[0001] [Field of the Invention] The present invention relates to the field of electrochemistry.

[0002] In one aspect, the present invention relates to the generation of an oxidant solution using an electrochemical cell.

[0003] In one particular embodiment, the present invention is suitable for use in metal extraction, such as from solid or solid particulate waste such as mined ore, or e-waste sources.

[0004] In another particular aspect, the present invention is suitable for use in a metal leaching process, or alternatively, an electrowinning process.

[0005] While it is convenient to describe the invention below in the context of e-waste, particularly copper-containing e-waste, it should be understood that the invention is not limited to that use alone and may be applied to other feedstocks or waste streams that contain one or more other metals or metal ions, such as, for example, those produced by mining.

[0006] [Background technology] Any discussion of documents, devices, acts, or knowledge in this specification should be understood to be included to explain the context of the invention. Moreover, the discussion throughout this specification arises out of the inventor's awareness and / or the inventor's identification of certain related art problems. Furthermore, any discussion of material, such as documents, devices, acts, or knowledge, in this specification is included to explain the context of the invention in terms of the inventor's knowledge and experience, and therefore, any such discussion should not be construed as an admission that any of the material formed part of the prior art standard or general knowledge in the relevant technical field, in Australia or elsewhere, prior to the priority date of the disclosure and claims herein.

[0007] Many raw materials and waste streams contain at least one metal compound. These include ores, landfill residues, sludges, tailings, slags, ash, filter dust from incinerators, blanks or e-waste, and wafers from electronic circuits. E-waste is defined as waste generated from a wide range of electronic devices, such as computers, mobile phones, televisions, and home appliances.

[0008] E-waste is a growing concern due to the increasing volumes it generates, resulting from rapid advances in technology and growing consumer demand for electronic products. Many countries have introduced legislation and policies for e-waste management. For example, in 2011, Australia implemented the National Television and Computer Recycling Scheme (NTCRS). The NTCRS aims to achieve 80% e-waste recycling by 2030 and to make industry-funded e-waste recycling available to households and small businesses to dispose of used electronics. The implementation of this type of recycling scheme provides an opportunity to extract and recycle valuable metals.

[0009] Metals can be recovered from waste streams by electrochemical means using oxidizing acids. In such processes, the oxidizing acids can be generated by oxidizing a starting acid at the anode of an electrochemical cell, the starting material is then washed with or immersed in the oxidizing acid to dissolve the metal or metal compound, and finally the dissolved metal is deposited at the cathode of the electrochemical cell.

[0010] Electrochemistry utilizes the flow of electrons to drive redox reactions. Electrochemical cells typically include two half-cells, one associated with an anode (positive electrode) and the other with a cathode (negative electrode), with an electrolyte between them to facilitate the reaction and movement of ions.

[0011] In electrochemical cells, oxidation of a metal occurs at the anode and reduction of a metal occurs at the anode in so-called "redox" reactions. Redox reactions occur in the presence of an oxidizing agent, usually in an electrolyte, which oxidizes another substance by accepting electrons and being reduced. Oxidizing acids are particularly useful in electrolytes used in electrochemical reactions because they can often oxidize metals that are less reactive with other acids.

[0012] For example, German Patent No. 10 2015 110 179 (DE 10 2015 110 179) describes the use of an oxidizing acid and a doped diamond cathode in an electrochemical cell for extracting metals from solid sources. The method has been successfully tested on solutions of filter dust containing Zn, Cu, Fe, Ni, and Sn, Chilean copper slag (CuFe), chalcopyrite (CuFeS2), and copper orifice.

[0013] According to DE 10 2015 110 179, the use of an oxidizing acid, such as peroxodisulfate, results in a significant increase in the amount of metal brought into solution compared to other acids, often at least doubling the amount of metal in solution. DE 10 2015 110 179 also asserts that neither the process parameters nor the concentration of the oxidizing acid are critical, but that the concentration of the oxidizing acid should be at least 0.1 mol / L, preferably at least 0.5 mol / L, and more preferably at least 1 mol / L. Concentrations greater than this are not required.

[0014] The use of oxidizing acids, such as peroxodisulfates, is also discussed in electrochemical systems for water disinfection and organic removal in U.S. Patent Nos. 10,259,727 and 10,046,898 to Advanced Diamond Technologies. Operation at high current densities with diamond anodes (and optionally diamond cathodes) is claimed to result in high current efficiency, longer operating life, and improved cost effectiveness.

[0015] [Summary of the Invention] It is an object of the present invention to enable a more efficient method for the electrochemical generation of oxidized analyte solutions.

[0016] Another object of the present invention is to improve the economics of processes that use oxidized analyte solutions to extract metals, including improving metal extraction and reducing energy consumption of the process.

[0017] Another object of the present invention is to provide a more effective oxidation analyte solution for extracting metals or metal ions from raw materials or waste streams.

[0018] It is yet another object of the present invention to provide an electrochemical process that can regenerate and reuse key electrolytic ions.

[0019] It is a further object of the present invention to alleviate at least one disadvantage associated with the related art.

[0020] It is an aim of the embodiments described herein to overcome or mitigate at least one of the disadvantages of the related art systems mentioned above, or at least to provide a useful alternative to the related art systems.

[0021] In a broad sense, the present invention relates to the synthesis of sulfate ions (SO4 2- ) and divalent iron ions (Fe 2+ ) from electrolyte raw materials containing peroxodisulfate (SO8 2- ) and trivalent iron ions (Fe 3+ The present invention relates to a method for electrochemically generating an analyte solution containing

[0022] In a first aspect of the embodiments described herein, there is provided a method for generating an oxidant solution using an electrochemical cell having an anode and a cathode, the method comprising: (i) supplying a raw electrolyte solution to a reaction zone between the anode and the cathode, wherein the raw electrolyte solution contains sulfate ions (SO 2- ) and divalent iron ions (Fe 2+ ) and (ii) During the operation cycle, the raw electrolyte is electrolyzed to produce peroxodisulfate (SO). 2- ) and trivalent iron ions (Fe 3+ ), and (iii) providing the oxidizing acid solution; Includes.

[0023] In a second aspect of the embodiments described herein, there is provided a method for producing an oxidant solution using an electrochemical cell having an anode half-cell and a cathode half-cell, the method comprising: (i) A step of supplying a raw electrolyte solution flowing into and out of a fluid path formed between the anode and the cathode to an anode half-cell, wherein the raw electrolyte solution is converted into sulfate ions (SO4 2- ) and divalent iron ions (Fe 2+ ) and (ii) During the operation cycle, the raw electrolyte is electrolyzed to produce peroxodisulfate (SO). 2- ) and trivalent iron ions (Fe 3+ ), and (iii) delivering the analyte solution from the anode half-cell; Includes.

[0024] Typically, the anode half-cell and the cathode half-cell are separated by a separator, such as a porous membrane or an ion exchange membrane. In a preferred embodiment, the ion exchange membrane is a cation exchange membrane or an anion exchange membrane, more preferably a monovalent anion selective membrane. The use of a separator advantageously allows the oxidizing power of the acid to be improved. As used herein, the term "separate cell" structure refers to a structure in which a separator is inserted to separate the anode and the cathode.

[0025] In another embodiment, the anode half-cell and cathode half-cell may be combined, for example, by removing the separator, to form a single electrochemical cell, with the analyte solution serving as the electrolyte for the electrochemical cell. As used herein, the term "non-separated cell" configuration refers to a configuration without any separator or other barrier between the anode and cathode.

[0026] The operating cycle of the present method may include a single pass or may be operated in a recycle mode. The operating cycle may include electrolyzing the electrolyte in a single batch or continuous flow system. Recycling and electrolyzing batch volumes of electrolyte may be used to increase the concentration of oxidant converted from the feedstock.

[0027] In one embodiment, the analyte solution from the chemical cell may be fed to a downstream process, such as an electrowinning cell. In another embodiment, the analyte solution may be generated on-site and fed directly to a process for hydrometallurgical leaching of metals from ore deposits, mines, mining waste dumps, and landfills. For example, the analyte may be generated at the mine, injected into a wellbore, propelled along the length of the wellbore, and then returned to the surface, where separation of the metal from the analyte may occur. In another embodiment, the analyte solution may be generated on-site and fed to replace the leach solution in common hydrometallurgical processes, such as heap leaching and various variations of tank and vat leaching.

[0028] The use of aqueous chemistry for the recovery of metals from ores, concentrates, and recycled or residual materials is known as hydrometallurgy. Hydrometallurgy offers a more economical and environmentally friendly alternative to metallurgical processes. It has shown advantages in processing low-quality ores, and mainly includes heap leaching, in-situ leaching, and tank leaching. However, hydrometallurgical processes have not been widely used in the processing of some sulfide ores, such as chalcopyrite, mainly due to the production of polysulfides (Sn 2- ) and elemental sulfur (SO) protective layers are responsible for the slow dissolution rate in acid. Therefore, researchers 3+ , O2, H2O2, Cr2O7 2- , ClO4 - , O3, MnO4 - and S2O8 - Leaching efficiency is enhanced by adding various oxidizing agents, including HCl, ...

[0029] In another embodiment, the analyte solution from the chemical cell may be used in a process for the leaching and selective recovery of metals from e-waste, including base metals such as Cu, Ni, Fe, precious metals such as Au, Ag, platinum group metals such as Pt, Pd, Rh, Ir, and Ru, rare metals such as Te, Ga, Se, Ta, and Ge, and hazardous metals such as Pb, Cd, In, Sb, etc. In a particularly preferred embodiment, the metals extracted from e-waste are Cu, Ni, Zn, and Al.

[0030] Typically, sulfate ions (SO4 2- ) concentration is between 0.1 molar and 5 molar.

[0031] Typically, sulfate ions (SO4 2- ) is sulfuric acid (H2SO42- ) from a raw material. The present invention uses a lower concentration of acid compared to equivalent prior art processes that use 15 to 20 moles of acid. The lower concentration of acid reduces raw material costs and improves process safety.

[0032] Typically, divalent ferrous ions (Fe 2+ The concentration of is between 0.1 and 0.5 molar.

[0033] Typically, the aqueous starting electrolyte has a current of about 50 to 200 mA / cm -2 It is electrolyzed at a current density of

[0034] Preferably, SO4 2- :Fe 2+ The ratio is between 1:0.05 and 1:0.5, more preferably between 1:0.05 and 1:0.1.

[0035] Typically the anode is a doped diamond electrode, for example a boron doped diamond electrode, however other electrode materials such as carbon composite, stainless steel, copper or titanium are also suitable.

[0036] Typically, the cathode is a doped diamond electrode, however other electrode materials such as carbon composite, stainless steel, copper, or titanium are also suitable.

[0037] The method described above may also include the regeneration and reuse of key electrolytic ions. For example, sulfate ions (SO4 2- ) and divalent iron ions (Fe 2+ ) is a raw electrolyte consisting of peroxodisulfate (SO8 2- ) and trivalent iron ions (Fe 3+ ) may be used to generate an analyte solution containing ferric ions (Fe 3+ ) is used to leach metals from waste, resulting in the production of ferrous ions (Fe 2+ ) is converted into ferrous ions (Fe2+ ) is peroxodisulfate (SO8 2- ), which reacts with sulfate ions (SO4 2- ) and trivalent iron ions (Fe 3+ ), thus completing the regeneration of ions from the starting electrolyte.

[0038] In a third aspect of the embodiments described herein, the method for regenerating the oxidant solution described above comprises: (iii) Ferrous ion (Fe 3+ ) is a divalent iron ion (Fe 2+ supplying the oxidizing acid solution to the metal-containing waste so that the metal-containing waste is reduced to (iv) Peroxodisulfate (SO) 2- ) to divalent iron ions (Fe 2+ ) to produce a solution containing ferric ions (Fe 3+ ) and sulfate ions (SO4 2- ), and Further includes:

[0039] Again, regeneration can be carried out with or without a separator between the anode and cathode half-cells.

[0040] In a fourth aspect of the embodiments described herein, (a) an anode and a cathode defining a reaction area of ​​an electrochemical cell; (b) an inlet port and a flow controller for passing the aqueous raw electrolyte between the electrodes, and the raw electrolyte is sulfate ions (SO4 2- ) and divalent iron ions (Fe 2+ ), and sulfate ions and divalent iron ions are peroxodisulfate (SO8 2- ) and trivalent iron ions (Fe 3+ an inlet port and flow controller selected for electrochemical generation of an oxidant including (c) current means for supplying a current for electrolyzing the aqueous raw electrolyte solution to produce an oxidant solution containing the oxidant in the reaction region; (d) an outlet port for supplying an oxidant solution from the electrochemical cell; An oxidizer solution generating system is provided, comprising:

[0041] The oxidant solution may be suitable for the extraction of metals from e-waste or minerals, including, for example, waste streams associated with mineral processing.

[0042] In a fifth aspect of the embodiments described herein, (a) an anode half-cell and a cathode half-cell; (b) an inlet port and flow controller for the passage of aqueous feed electrolyte through the anode half-cell, where the feed electrolyte is sulfate ions (SO4 2- ) and divalent iron ions (Fe 2+ ), and sulfate ions and divalent iron ions are peroxodisulfate (SO8 2- ) and trivalent iron ions (Fe 3+ an inlet port and flow controller selected for electrochemical generation of an oxidant including (c) current means for supplying a current for electrolyzing the aqueous feedstock electrolyte to produce an oxidant solution containing the oxidant in the anode half-cell; (d) an outlet port for supplying oxidant solution from the electrochemical cell from the anode half-cell;

[0010] A system for on-site generation of oxidant solution for metal extraction is provided, comprising:

[0043] The oxidant solution generated on-site may be suitable for the extraction of metals from e-waste or minerals, including, for example, waste streams associated with mineral processing.

[0044] In another aspect of the embodiments described herein, there is provided a method for leaching metals from metal-bearing waste, the method comprising providing an oxidant solution of the present invention and contacting the oxidant solution with metal-bearing waste.

[0045] Metals leached from metal-bearing waste typically include copper.

[0046] Other aspects and preferred configurations are disclosed herein and / or are defined in the appended claims, which form part of the description of the invention.

[0047] Essentially, embodiments of the present invention involve the use of ferrous ions (Fe 2+ ) can substantially improve the generation of oxidant solutions using electrochemical cells. 2- ) and trivalent iron ions (Fe 3+ ) for the electrochemical generation of oxidants containing sulfate ions (SO4 2- ) combined with ferrous ions (Fe 2+ It has also been recognized that the combination of oxidant and acid (peroxodisulfate (SO2)) provides superior oxidants for metal extraction. 2- ) and trivalent iron ions (Fe 3+ It has also been recognized that it is also advantageously possible to regenerate

[0048] The advantages provided by the present invention include: Improved efficiency of electrochemical cells in terms of power consumption Improved electrochemical cell efficiency from an economical perspective The oxidant / acid can be regenerated between multiple operating cycles of the electrochemical cell, thus lowering operating costs and reducing waste. Lower concentrations of acid can be used compared to conventional techniques, thus reducing raw material costs and improving process safety Complete extraction of certain metals can be achieved in a relatively short time

[0049] Further scope of applicability of embodiments of the present invention will become apparent from the detailed description given hereinafter. It should be understood, however, that the detailed description and specific examples, while indicating preferred embodiments of the invention, are given by way of illustration only, since various changes and modifications within the spirit and scope of the disclosure herein will become apparent to those skilled in the art from this detailed description. [Brief explanation of the drawings]

[0050] Further disclosure, objects, advantages and aspects of preferred and other embodiments of the present application may be better understood by those skilled in the relevant art by reference to the following description of embodiments, which is given by way of example only and therefore does not limit the disclosure herein, in conjunction with the accompanying drawings. [Figure 1] FIG. 1 shows a typical electrochemical cell. [Figure 2] FIG. 2 shows a reactor comprising an electrochemical cell with a membrane separating the anode half-cell from the cathode half-cell. [Figure 3] Figure 3A shows an electrochemical cell of the type used in the present invention. Figure 3 is a plot of the concentration of peroxodisulfate ion (SO 2 ) versus time obtained using the electrochemical cell shown in Figure 3A. [Figure 4] Figure 4A shows an electrochemical cell of the type used in the present invention, and Figure 4B is a plot of the conversion of ferrous iron (Fe2+) to ferrous iron (Fe3+) versus time in the cell of Figure 4A. [Figure 5] FIG. 5 shows an electrochemical cell for in situ copper extraction according to the present invention. [Figure 6] 6A and 6B are plots showing the extent of copper conversion in the electrochemical cell of FIG. [Figure 7] FIG. 7 is a plot of copper conversion versus reaction time for the electrochemical cell of FIG. [Figure 8] FIG. 8 is a plot of iron composition (Fe2+ / Fe3+) versus reaction time measured using the electrochemical cell of FIG. [Figure 9] FIG. 9 is a plot of power consumption measured using the electrochemical cell of FIG. 5 for leaching media of different compositions. [Figure 10] FIG. 10 is a plot of power consumption measured using the electrochemical cell of FIG. 5 versus reaction time. [Figure 11] FIG. 11 is a schematic diagram of an electrolytic cell for e-waste leaching. [Figure 12] FIG. 12 is a schematic diagram of a pilot-scale electrolyzer having a separated cell configuration according to the present invention. [Figure 13] FIG. 13 is a schematic diagram of an electrodeposition reactor for metal recovery according to the present invention. [Figure 14] FIG. 14 is a flow chart showing the general concept of process development for metal leaching. [Figure 15] FIG. 15 is a schematic diagram of oxidant generation (Fe3+ and / or peroxodisulfate) and in-situ leaching of e-waste in a non-separated cell. [Figure 16] FIG. 16 is a schematic diagram of oxidant generation (Fe3+ and peroxodisulfate) in a separate cell and single-step leaching. [Figure 17] Figure 17 shows the concentration profile of peroxodisulfate at different initial sulfate concentrations and under the same reaction conditions (i.e., current density = 150 mA cm-2 and flow rate = 50 mL min-1). [Figure 18] Figures 18A and 18B show the profiles of iron(III) and peroxodisulfate ions as a function of reaction time in a separation cell configuration with an anion exchange membrane at various initial sulfate concentrations: (A) 0.5 M Fe2+ and (B) 0.25 M Fe2+ under the same reaction conditions (i.e., current density = 150 mA cm-2 and flow rate = 50 mL min-1). [Figure 19] Figures 19A and 19B show profiles of (A) extracted copper (wt%) and (B) molar fraction of ferric ions as a function of reaction time at various SO4 2- / Fe2+ ratios under the reaction conditions of current density = 150 mA cm-2 and flow rate = 50 mL min-1. [Figure 20] Figure 20 is a plot of the concentration of extracted metals as a function of reaction time in a laboratory-scale electrolytic cell under reaction conditions of SO4 2- / Fe2+ = 1 / 0.1, current density = 150 mA cm-2, and flow rate = 50 mL min-1. [Figure 21]Figure 21 is a plot of metals extracted in a pilot-scale electrolytic cell as a function of reaction time. Reaction conditions: SO4 2- / Fe2+ = 1 / 0.1, current density = 150 mA cm-2, and flow rate = 1 L min-1. [Figure 22] Figure 22 is a plot of the concentration of extracted metals as a function of reaction time for three batches of e-waste input in a pilot-scale electrolyzer under the following conditions: initial sulfate concentration of 2.5 M, current density = 150 mA cm-2, and flow rate = 1 L min-1. [Figure 23] Figure 23 shows a proposed design for a commercial electrolyzer. [Figure 24] FIG. 24 is a plot of copper recovery as a function of reaction time. [Figure 25] FIG. 25 is a plot of the oxidation-reduction potential (ORP) versus time as a function of reaction time for laboratory and pilot-scale electrolyzers under conditions of SO 4 2− / Fe 2+ = 1 / 0.1 and current density = 150 mA cm −2 . [Figure 26] Figure 26 is a plot of ORP versus time as a function of reaction time for three batches of e-waste input under the conditions of an initial sulfate concentration of 2.5 M, current density = 150 mA cm-2, and flow rate = 1 L min-1 for the pilot-scale electrolyzer. [Figure 27] FIG. 27 is a plot of copper recovery from chalcopyrite as a function of time and temperature for a laboratory-scale reactor under conditions of SO 4 2− / Fe 2+ = 1 / 0.1, current density = 100 mA cm −2 , and 700 RPM. [Figure 28] FIG. 28 is a plot of copper recovery from chalcopyrite as a function of time depicting the effect of mixing at 700 rpm with or without ultrasound under conditions of SO / Fe = 1 / 0.1, current density = 100 mA cm, and 700 RPM. [Figure 29] FIG. 29 is a plot of the effect of stirring speed on copper recovery from chalcopyrite under conditions of SO4 2- / Fe2+=1 / 0.1 and current density=100 mA cm-2. [Figure 30] Figure 30 is a plot of the concentration of metals extracted for one batch of nickel ore 1 as a function of reaction time in a laboratory-scale electrolytic cell under the following conditions: initial sulfate concentration 2.5 M, SO4 2- / Fe2+ = 1 / 0.1, current density = 150 mA cm-2, and flow rate = 50 mL min-1. [Figure 31] Figure 31 is a plot of the concentration of metals extracted for one batch of nickel ore 2 as a function of reaction time in a laboratory-scale electrolytic cell under the following conditions: initial sulfate concentration 2.5 M, SO4 2- / Fe2+ = 1 / 0.1, current density = 150 mA cm-2, and flow rate = 50 mL min-1. [Figure 32] Figure 32 shows the improvement in metal recovery after 2 hours of treatment in an electrolytic cell with analyte solution relative to standard acid leaching using 2.5 M H2SO4. The electrolytic cell conditions were: initial sulfate concentration 2.5 M, SO4 2- / Fe2+ = 1 / 0.1, current density = 150 mA cm-2, and flow rate = 50 mL min-1. [Figure 33] FIG. 33 is a simplified flow diagram depicting the standard heap leaching process. [Figure 34] FIG. 34 is a simplified schematic of the Shanks System. [Figure 35] FIG. 35 is a schematic diagram of a front view (FIG. 35A) and a top view (FIG. 35B) of a Rotocel extraction device. DETAILED DESCRIPTION OF THE INVENTION

[0051] [Detailed explanation] For purposes of this description, terms such as “upper,” “lower,” “right,” “left,” “rear,” “front,” “vertical,” “horizontal,” “interior,” and “exterior,” as well as derivatives thereof, refer to the present invention as oriented in FIG. 5 . However, it should be understood that the present invention may assume various alternative orientations unless expressly specified otherwise. The specific devices and processes illustrated in the accompanying drawings and described in the following specification are also to be understood as merely exemplary embodiments of the inventive concepts defined in the appended claims. Therefore, specific dimensions and other physical characteristics related to the embodiments disclosed herein are not to be considered limiting unless expressly claimed in the claims. Furthermore, unless otherwise specified, discussion of a particular feature of a component as extending in or along a certain direction, etc., should not be understood to imply that the feature or component follows a straight line or axis in such direction, or extends solely in such direction or on such plane, without other directional components or biases, unless otherwise specified.

[0052] Figure 1 shows a typical electrochemical cell comprising two half-cells, one associated with an anode (positive electrode) (1) and the other associated with a cathode (negative electrode) (2), with an electrolyte between them. In the associated redox reaction, a metal is oxidized at the anode and the oxidized metal is reduced at the cathode. Redox reactions occur in the presence of an oxidizing agent, usually in an electrolyte, which oxidizes another substance by accepting electrons and being reduced.

[0053] Figure 2 shows an electrochemical cell with a membrane (3) that allows only negatively charged ions to migrate from the anode (1) to the cathode (2). Positively charged ions remain on the anode side of the membrane.

[0054] Oxidizing acids are often included in the electrolyte of electrochemical cells because they can oxidize some metals that are less reactive with other acids. The present invention provides a method for oxidizing sulfate ions (SO4 2- ) and divalent iron ions (Fe2+ ) from electrolyte raw materials containing peroxodisulfate (SO8 2- ) and trivalent iron ions (Fe 3+ It involves the electrochemical generation of an analyte solution containing peroxodisulfate ions (SO8 2- ) is a sulfate group (SO4 ·- Peroxodisulfate ions have strong oxidizing properties due to their ability to produce sulfuric acid (SFA). Chemical dissolution of metals with peroxodisulfate ions, also known as oxidative dissolution or metal leaching, involves the process of dissolving metals in solution by utilizing peroxodisulfate salts as oxidizing agents. As a result of this dissolution, the dissolved metal will participate in the activation of peroxodisulfate ions, generating sulfate groups. This provides an environmentally friendly alternative to traditional leaching agents such as cyanide and sulfuric acid, which pose significant environmental risks.

[0055] Preferably, the oxidant solution is generated using an electrochemical cell comprising an anode half-cell and a cathode half-cell. The feed electrolyte is passed between the anode and the cathode. If a porous membrane is used, the feed electrolyte passes through the anode half-cell. Sulfate ions (SO4 2- ) and divalent iron ions (Fe 2+ ), which is oxidized during the operation cycle of the electrochemical cell to form peroxodisulfate (SO8 2- ) and trivalent iron ions (Fe 3+ ) to produce an oxidizing acid solution containing

[0056] The oxidizing acid solution from the electrochemical cell can be generated in the reaction zone between the anode or cathode. Alternatively, if the anode and cathode half-cells are separated by a separator, such as a porous film or ion exchange membrane, the oxidizing acid solution can be generated in the anode half-cell. A reservoir in fluid communication with the anode half-cell can be used to circulate and store the analyte solution.

[0057] The operating cycle of the electrochemical cell may include recycling and electrolyzing a batch volume of electrolyte to increase the concentration of oxidant converted from the feedstock.

[0058] [Metal-containing waste] The operating cycle of the electrochemical cell may involve electrolyzing metal-containing waste in a single pass or continuous flow system. The dissolved metal may be separated by deposition at the cathode and the acid may be regenerated by reoxidation at the anode, thus allowing the process to be operated in a single pass. Alternatively, the process may be operated continuously where leached waste material may be removed from the reaction zone and new waste material may be supplied.

[0059] The process can be used to leach metals from liquids, solutions, or solid materials, such as circuit boards or wafers, and deposit them at a cathode. Metals deposited at a cathode can be easily removed from the cathode by a variety of well-known physical methods. Alternatively, electrochemical means may be employed to return the deposited metal to solution and redeposit the metal in pure form at another cathode.

[0060] For example, the oxidizing acid solution from the chemical cell may be fed to downstream processes to extract valuable metals from e-waste. E-waste typically contains 40% metals, 30% ceramics (silica, alumina, alkaline earth oxides), and 30% plastic materials (polyethylene, polypropylene, polyvinyl chloride, polystyrene, epoxy, nylon, etc.). These metals can be further classified into several categories, including base metals (Cu, Ni, Fe), precious metals (Au, Ag), platinum-group metals (Pt, Pd, Rh, Ir, and Ru), rare metals (Te, Ga, Se, Ta, and Ge), and toxic metals (Pb, Cd, In, Sb).

[0061] More specifically, the most abundant element in e-waste is copper (Cu) (approximately 32 wt%) in e-waste, while the concentrations of other elements, such as iron (Fe), aluminum (Al), tin (Sn), and zinc (Zn), are 13, 5.7, 1.9, and 1.7 wt%, respectively.

[0062] The general concept of metal recovery from e-waste is depicted in Figure 14. The present invention aims to integrate the generation of oxidant using leaching, followed by the regeneration of the oxidant from spent solution, and the recovery of selective metals from e-waste. The various steps of this process could be integrated into a single operation or separated into a series of unit operations.

[0063] In another process, the oxidizing acid solution from the chemical cell may be fed to a downstream process, such as an electrowinning cell. In another embodiment, the oxidizing acid solution may be generated and fed on-site and directly input into processes for leaching metals from ore deposits, mines, mining waste dumps, and landfills. For example, the oxidizing acid solution may be generated at the mine, injected into a mine borehole, propelled along the length of the borehole, and then returned to the surface where separation of the metals from the analytes can occur.

[0064] The use of electrochemical oxidation (EO) to generate oxidants in situ within a reactor system is a sustainable and promising approach due to the improvement of renewable energy technologies that reduce electricity costs. The indirect oxidation process is widely recognized as the primary mechanism for ore electrolysis. Therefore, generating oxidants within a reactor system is a key step for ore electrooxidation.

[0065] OER is a competing reaction in conventional slurry electrolysis processes, resulting in reduced current efficiency. It is clear that electrodes with low OER overpotentials (high OER catalysis), such as IrO2, RuO2, and platinum, generate strongly absorbed radical species, which subsequently lead to O2 formation and reduce the oxidation potential of ROS. In contrast, electrodes with high OER overpotentials (low OER catalysis), such as SnO2, PbO2, and BDD, are effective in suppressing O2 production. At the same time, they exhibit stronger oxidation potential for weakly adsorbed ROS, making them an ideal choice for electrolytic slurries. BDD anodes have attracted attention in electrochemical oxidation processes due to their unique properties, such as a very wide potential window, strong corrosion protection stability, high energy efficiency, an inert surface with low adsorption properties, and a wide operating pH range. [Example]

[0066] The invention will now be further described with reference to the following non-limiting examples.

[0067] Example 1: Figure 3A shows an electrochemical cell of the type used in the following examples. The electrochemical cell has a working volume of 80 mL. On the left side of the half-cell, sulfuric acid (H2SO4 2- ) from sulfate ions (SO4 2- ) is oxidized to peroxodisulfate ion (SO8 2- On the right side of the half-cell, catholyte (5) contains copper ions, which are reduced at the cathode and deposited as copper metal (Cu).

[0068] Example 2: Figure 3B shows the peroxodisulfate ion (SO) extracted using the electrochemical cell shown in Figure 3A. 2- ) concentration versus time. 2- The concentrations of sulfuric acid used were 0.615M, 1.3M, 2.3M, and 5.0M. The higher the sulfuric acid concentration, the greater the amount of peroxodisulfate ions (SO) formed in the analyte. 2-) increases, thereby improving the oxidizing ability of the electrolyte to extract copper.

[0069] Example 3: FIG. 4A shows an electrochemical cell with a working volume of 80 mL, where 4.5 M sulfuric acid (HSO) in analyte (4) 2- ) and 0.5M ferrous sulfate (Fe2SO4 2- ) from sulfate ions (SO4 2- ) is oxidized to form peroxodisulfate ions (SO8 2- ) and trivalent iron ions (Fe 3+ ) is produced. A separator in the form of an ion exchange membrane helps to improve the oxidizing power of the acid produced. In the right-hand half-cell, the catholyte contains copper ions, which are reduced at the cathode and deposited as copper metal (Cu). The current density used is 50 mA cm -2 is.

[0070] This arrangement is similar to the on-site supply of analyte solution to a metal leaching process from an electrochemical cell, which may involve, for example, injecting the analyte into a wellbore in a mine shaft, traveling down the length of the wellbore, and then returning it to the surface where separation of the metal from the analyte can take place.

[0071] Figure 4B shows the relationship between ferrous ions (Fe 2+ ) and trivalent iron ions (Fe 3+ ) versus time. The plot shows the mole fraction of peroxodisulfate ion (SO8 2- ) is converted into ferrous ions (Fe 2+ ) and trivalent iron ions (Fe 3+ ) and the porous membrane prevents the migration of ions to the right cell where they would otherwise be reduced at the cathode. The plot also shows that an oxidizing solution for metal extraction can be obtained after 3 hours of reaction time.

[0072] Example 4: Figure 5 shows an electrochemical cell for in situ copper extraction. The cell shown is the same as the cell shown in Figure 4A, with an ion exchange membrane (3) as a separator between the anode and cathode half-cells. In this embodiment of the invention, divalent copper ions (Cu 2+ ) was added to the analyte solution (4) and reduced to copper (Cu), which precipitated in the analyte container.

[0073] FIG. 5 also shows the reactions that facilitate the regeneration and reuse of key electrolyte ions. As shown, during the operating cycle, sulfate ions (SO 2- ) and divalent iron ions (Fe 2+ The raw electrolyte consists of peroxodisulfate (SO8 2- ) and trivalent iron ions (Fe 3+ ) may be used to generate an analyte solution containing ferric ions (Fe 3+ ) is a divalent copper ion (Cu 2+ ), resulting in the leaching of ferrous ions (Fe 2+ ) is converted into ferrous ions (Fe 2+ ) is peroxodisulfate (SO8 2- ), which reacts with sulfate ions (SO4 2- ) and trivalent iron ions (Fe 3+ ), thus completing the regeneration of ions from the starting electrolyte.

[0074] In this reaction, the reaction variables for copper extraction can be expressed as follows:

number

[0075] Furthermore, the response variable for power consumption can be expressed as:

number

[0076] Example 5: Example 5 demonstrates the effect of current on the oxidant in copper conversion. Example 5 also demonstrates the conversion of peroxodisulfate ions (SO) to oxidants in the analyte. 2- ) and trivalent iron ions (Fe 3+ ) and is compared to analytes commonly used in the prior art that do not use both of these electrolytes.

[0077] 6A and 6B show the extent of copper conversion under various conditions in the electrochemical cell of FIG. 5 with a working volume of 80 mL. In all cases, sulfate ions (SO 2- ) concentration was 2.5 M, and the current density was 150 mA cm -2 The reaction time was 1 hour, and the initial copper concentration was 3.0 g.

[0078] As a control, no current was initially applied to the cell to indicate no copper conversion. Applying a current promoted the generation of oxidizing agents and initiated copper extraction. Figure 6A shows the peroxodisulfate ion (SO) 2- Figure 6B shows the extraction of small amounts of copper in an analyte containing only peroxodisulfate ions (SO). 2- ) and trivalent iron ions (Fe 3+ ) and shows significantly better copper extraction in analytes containing both.

[0079] 1:0.2 ratio (sulfate ions (SO4 2- ) versus divalent iron ions (Fe 2+ )) gives better results than a ratio of 1:0.6.

[0080] Example 6: Example 6 investigated the effect of reaction time on the extent of copper conversion.

[0081] FIG. 7 is a plot of copper conversion versus reaction time for the electrochemical cell of FIG. 5. The reaction involved sulfate ions (SO4) at a concentration of 2.5 M. 2- ), 150mA cm -2A current density of 0.01, a working volume of 80 mL, and an initial 3.0 g of copper were included. The plot shows how the copper conversion increases with time, reaching a maximum conversion concentration (approximately 80%) after 3 hours of reaction time.

[0082] Example 7: Example 7 shows that the reaction time is 3+ ) on the degree of production.

[0083] Figure 8 shows the iron composition (Fe 2+ / Fe 3+ ) versus reaction time. This reaction was performed using 2.5M sulfate ions (SO4 2- ), 150mA cm -2 The current density was 0.01, the working volume was 80 mL, and an initial 3.0 g of copper was included. 2+ ) to iron (Fe 3+ Complete conversion to ferric ions (Fe) was achieved after 2 hours of reaction time. After 2 hours, copper extraction was reversed to ferric ions (Fe 3+ ) and peroxodisulfate ion (SO8 2- ) is caused by the presence of

[0084] Example 8: Example 8 investigated the effect of current and oxidant on the power consumption of an electrochemical cell.

[0085] Figure 9 is a plot of the power consumption measured over a time period (t) of 1 hour using the electrochemical cell of Figure 5 versus the composition of various leaching media. The reaction included sulfate ions (SO4) at a concentration of 2.5 M. 2- ), 150mA cm -2 The current density was 100 kJ / s, the working volume was 80 mL, and the initial copper was 3.0 g. The power consumption was 1:0.2 (sulfate ion (SO4)) with 2.5 M sulfuric acid (H2SO4) only. 2- ) versus divalent iron ions (Fe 2+ )) and a 1:0.6 (sulfate ion (SO4 2- ) versus divalent iron ions (Fe 2+ )) mixture was measured for analytes including

[0086] The plot shows the sulfate ion (SO4 2- ) versus divalent iron ions (Fe 2+ ) ratio is important.

[0087] Example 9: Example 9 investigated the effect of reaction time on the power consumption of an electrochemical cell.

[0088] Figure 10 is a plot of power consumption measured using the electrochemical cell of Figure 5 versus reaction time. The reaction involved the addition of sulfate ions (SO4) at a concentration of 2.5M. 2- ), 150mA cm -2 A current density of 0.1 g, a working volume of 80 mL, and an initial copper weight of 3.0 g were included.

[0089] [E-waste leaching]

[0090] Example 10: In this example, e-waste was collected from used computer units, primarily CPU mainboards and desktop screens. The e-waste material was first shredded into small particles, sieved to a particle size of 2.7±1.2 mm (<4 mm), and then mixed with aqua regia (a 3:1 mixture of HNO3 and HCl). For a typical analysis, 0.5 g of shredded e-waste was mixed with 12 mL of aqua regia. The e-waste and aqua regia mixture was then heated at 180°C for 30 minutes in a microwave system. Under these conditions, most metals dissolved into solution, and their concentrations were measured.

[0091] [Electrolytic cell equipment for leaching metals from e-waste]

[0092] Electrolytic cells for leaching metals from e-waste were constructed at two different scales (i.e., laboratory scale and pilot scale). The electrolytic cells consisted of a working electrode (anode) and a counter electrode (cathode). The anode was made of BDD material, while a stainless steel plate was used as the cathode. Experimental metal leaching from e-waste was performed using a 150 mA cm from a DC power source.-2 The experiment was carried out under a constant power density of .

[0093] Figure 11 shows a schematic diagram of a separate electrolytic cell according to the present invention. Two vessels (4, 5) (pre-charged with 100 mL of solution) were used to recirculate the solution through the anode and cathode compartments separately. Both vessels were charged with an oxidiser precursor solution (0.5-2.5 M H2SO4), and the e-waste sample was charged into the analyte vessel (4). 50 mL min -1 A constant flow rate of 1000 s was delivered by a peristaltic pump (8). The surface area of ​​each electrode was 10 cm 2 It was.

[0094] A pilot-scale electrolytic cell (7), as shown in Figure 12, was utilized to evaluate metal extraction in larger volumes and followed a similar configuration to the laboratory-scale electrolytic cell. The working electrode contained BDD and the counter electrode contained stainless steel, with each electrode measuring 100 cm 2 The surface area was 1000 m.

[0095] In a typical run, 50 g of e-waste was loaded into a vessel (referred to as a fixed bed) as shown in Figure 12. 1 L of an oxidant precursor solution (e.g., sulfate and ferrous ions) was loaded into the analyte and catholyte chambers. The outlet for the analyte stream containing the oxidant was connected to the bottom of the fixed bed, passed through the bed, and then recycled back to the analyte vessel.

[0096] The solution containing the leached metals from the pilot-scale electrolytic cell was further processed using an electrodeposition process to recover the metals. The apparatus for metal recovery is shown in Figure 13. The solution was recirculated through the electrolytic cell. The dissolved metal was deposited on the anode rod installed in the center of the cell and on 400 cm 2 The electrodeposition was carried out in an annular cell (11) with a cathode sheet having a surface area of ​​1000 .mu.m. The experiment was carried out at a constant current density and flow rate, i.e., 50 mA cm. -2 and 5.5 L / min.

[0097] Example 11: The concentrations, oxidation strength, and pH values ​​of peroxodisulfate and metals obtained during metal recovery from e-waste were measured and are described in the following paragraphs. Peroxodisulfate concentrations were measured using a ThermoFisher ion chromatograph. Dissolved metals were measured using inductively coupled plasma (ICP). Oxidation strength and pH were measured using a Mettler Toledo multiparameter sensor probe.

[0098] Approach 1: Electro-oxidation in unseparated cells for in situ leaching of metals.

[0099] In the first approach, the electrooxidation step was carried out in a laboratory-scale electrolytic cell with a non-isolated structure. In-situ leaching was carried out by generating peroxodisulfate oxidant from sulfate oxidation and by leaching Fe in the analyte container. 2+ and the immediate consumption of the peroxodisulfate oxidant to oxidize a pre-loaded solid metal sample (e.g., e-waste). The main reactions that can occur in an electrolytic cell using this approach are shown in Figure 15.

[0100] Approach 1 was tested in the absence of a solid metal sample to observe the rate of peroxodisulfate production from sulfate oxidation at the BDD anode. After 2 hours of operation, only small amounts of peroxodisulfate (approximately 7% yield) were measured in solution. While not wishing to be bound by theory, the low production of peroxodisulfate in this configuration suggests that there is a competing reaction that prevents the accumulation of peroxodisulfate in solution. In the non-separated cell configuration, peroxodisulfate ions migrated to the cathode and were reduced back to sulfate ions.

[0101] A copper wire (representing an e-waste sample) was then introduced into the system to observe the copper leaching rate in a non-separated cell configuration. Approximately 30% copper leaching was achieved after 3 hours of operation. Most of the leached copper was deposited on the cathode side, suggesting that dissolved copper was electrodeposited at the cathode.

[0102] Approach 2: Electro-oxidation in separate cells for in situ leaching of metals.

[0103] In the second approach, the electro-oxidation step was carried out in an electrolytic cell with an anode and cathode compartments separated by a monovalent selective anion exchange membrane. - and Br - Only monovalent ions such as Fe were allowed to pass through the membrane. 3+ ) was then sequestered in the analyte compartment and therefore not reduced. At the end of the process, the analyte solution contains dissolved metals that can be further recovered in a separate unit run. The main reactions that can occur in an electrolytic cell using this approach are shown in Figure 16.

[0104] The use of the membrane made a significant difference in the production of peroxodisulfate. In the absence of a solid metal sample, a significant amount of peroxodisulfate (approximately 48% yield) was measured in the analyte solution, suggesting that the membrane eliminated any competing reactions that prevented peroxodisulfate generation. The membrane may have prevented the reduction of peroxodisulfate on the cathode side. When a copper wire was introduced into the system, approximately 50% copper extraction was achieved after 3 hours of operation. A clear blue analyte solution and a colorless catholyte solution were visually observed at the end of the process, suggesting that the dissolved copper was sequestered in the analyte solution.

[0105] Based on these two approaches, the separated cell structure shows better performance in oxidant generation and copper leaching. The following results were collected from the experiments in the separated cell structure.

[0106] [result]

[0107] Figure 17 shows the profiles of peroxodisulfate concentration and yield as a function of reaction time. Peroxodisulfate production was observed in a microflow electrolysis cell at 150 mA cm, respectively. -2 and 50 mL min -1 The experiments were carried out at various initial sulfuric acid concentrations (i.e., 0.5, 1, and 2.5 M) using a current density and flow rate of 1000 kJ / s. As can be seen in Figure 17, the concentration of peroxodisulfate shows an increasing trend as the reaction time increases from 0.25 to 1 h at all initial sulfuric acid concentrations investigated. The same trend can be observed at various initial sulfuric acid concentrations at each reaction time, where higher initial sulfuric acid concentrations produce larger amounts of peroxodisulfate. This trend highlights the importance of the initial sulfuric acid concentration in influencing peroxodisulfate production and may have important implications for optimizing electrochemical processes for metal extraction.

[0108] The production of peroxodisulfate was also tested using anion-exchange and cation-exchange membranes. There was no significant difference in the production of peroxodisulfate, suggesting that both membranes can prevent the migration of peroxodisulfate to the cathode side.

[0109] Figure 18 shows the profiles of ferric iron and peroxodisulfate concentrations as a function of reaction time at various initial sulfuric acid and ferrous iron concentrations. As shown in Figure 18, at a constant initial ferrous iron concentration of 0.5 M, the rate of ferric iron production was independent of the initial sulfuric acid concentration, i.e., complete conversion of ferrous iron to ferric iron was achieved within 2 hours. The generation of peroxodisulfate was observed to gradually increase after 1 hour of reaction time, from 0.05 to 0.35 M at 3 hours of reaction time.

[0110] Without wishing to be bound by theory, the mechanism of peroxodisulfate formation is believed to be likely initiated by the electro-oxidation of sulfate ions to form peroxodisulfate (shown in equation (3)).

number

[0111] Peroxodisulfate is a stable oxidizing agent that subsequently oxidizes ferrous iron to ferric iron. The reaction of peroxodisulfate with ferrous iron proceeds at a stoichiometric ratio of 1, i.e., 1 mole of peroxodisulfate reacts with 1 mole of ferrous iron to produce 1 mole of ferric iron, as shown in equation (4).

number

[0112] Based on Figure 18, complete conversion of ferrous iron to ferric iron at initial ferrous iron concentrations of 0.25M and 0.5M was achieved after 0.5 and 1 hour of reaction time, respectively. This result is consistent with the observations made with respect to Figure 13, where approximately 0.3M and approximately 0.6M peroxodisulfate were produced after 0.5 and 1 hour of reaction time. The peroxodisulfate was consumed to promote the oxidation of ferrous iron to ferric iron. As a result, the generation of peroxodisulfate in solution proceeds only when complete conversion of ferrous iron to ferric iron, which occurs after 1 hour of reaction time, is achieved.

[0113] The rate of ferric iron production was significantly faster at low initial ferrous iron concentrations (e.g., 0.25 M), where complete conversion of ferrous iron to ferric iron was achieved after 45 minutes of reaction time (as shown in Figure 19B). Consistent with what was observed at an initial ferrous iron concentration of 0.5 M, the initial sulfuric acid concentration did not affect the rate of ferric iron production.

[0114] Figure 19 shows various SO4 2- / Fe 2+ 19A shows the profiles of copper extraction and ferrous ion concentration as a function of reaction time in ratio. As shown in FIG. 19A, the amount of extracted copper was 100% of the total SO4 2- / Fe 2+The ratio showed a steady increase as a function of reaction time. Only a relatively small amount of copper (about 15%) was extracted, and the SO4 2- / Fe 2+ The ratio had no apparent effect on copper extraction over a 1-hour reaction time. In the absence of ferrous ions, approximately 50% of the copper was extracted after 3 hours of reaction time. The addition of ferrous ions to the leaching process, i.e., 1 / 0.1 SO4 2- / Fe 2+ The ratio increased the copper extraction rate to about 75%, while the SO4 2- / Fe 2+ A further increase in the ratio to 1 / 0.5 resulted in about 57% copper extraction.

[0115] Various SO4 2- / Fe 2+ The profile of the production of ferric ions as a function of reaction time, in terms of ratio, is shown in Figure 19B. It is clear that the initial concentration of ferrous ions influences the rate of production of ferric ions. At low concentrations of ferrous ions (e.g., SO4 2- / Fe 2+ = 1 / 0.05 and 1 / 0.1), complete conversion of ferrous iron to ferric iron occurred within 1 hour. 2- / Fe 2+ A further increase in the ratio to 1 / 0.5 resulted in a slower conversion of ferrous iron to ferric iron, i.e., approximately 70-80% conversion of ferrous iron was achieved after 3 hours.

[0116] Low ferrous iron concentration (SO4 2- / Fe 2+ At ratios of 1 / 0.05 and 1 / 0.1, peroxodisulfate was consumed over 1 hour to oxidize ferrous iron and leach copper, resulting in a small amount (about 15%) of extracted copper. Once complete conversion of ferrous iron was achieved (after 1 hour), copper leaching was then assisted by ferric iron and peroxodisulfate, resulting in a significant increase in copper leaching rate, from about 10% to 75% as the reaction time increased from 1 hour to 3 hours. High concentrations of ferrous ions (e.g., SO4 2- / Fe 2+When the SO4 ratio (=1 / 0.4 and 1 / 0.5) was added, peroxodisulfate was used up in the oxidation of ferrous iron to ferric iron. Under these conditions, copper leaching was probably assisted only by ferric iron. 2- / Fe 2+ =1 / 0.5, the amount of copper extracted is the same as in the absence of ferrous ions (SO4 2- / Fe 2+ =1 / 0), i.e., 50% copper extraction with a single oxidant was achieved after 3 hours of reaction time. The trend shown in Figure 19 is consistent with the amount extracted with SO4 2- / Fe 2+ This suggests that a ratio of 0.1 / 1 is the optimal oxidant ratio at which copper extraction can be maximized. This confirms the results presented in Figures 13 and 14, where oxidant generation is initiated by the production of peroxodisulfate, which can then oxidize any species in solution.

[0117] [Extraction of metals from waste electrical and electronic equipment on a laboratory scale]

[0118] Figure 20 shows the profile of metal concentration in the post-leaching solution as a function of reaction time. Leaching was carried out in a laboratory-scale electrolytic cell at 150 mA cm -2 current density, SO4 2- / Fe 2+ = 1 / 0.1, and flow rate = 50 mL min -1 The elements measured in the solution included Cu, Al, Fe, Zn, Ni, and Pb.

[0119] As shown in Figure 20, all metal concentrations increased during the first hour of reaction, gradually decreased after 2 hours, and then remained constant up to 4 hours of reaction time. Cu and Fe were the most abundant elements measured in the post-leaching solution (reaching approximately 10 g / L and approximately 7 g / L, respectively, after 4 hours of leaching). The concentrations of other elements, including Zn, Ni, Al, and Ni, ranged from 0.1 to 0.5 g / L, while the concentration of Pb in the solution was less than 0.005 g / L.

[0120] Considering that Cu and Fe are the most abundant elements in e-waste, it is not surprising that the concentrations of Cu and Fe in the leaching solution are relatively higher than those of other elements. However, in a comparison of the extraction of Zn versus Al, the Zn concentration was higher than that of Al (0.8 g / L vs. 0.2 g / L), while the Zn content in e-waste was lower than that of Al (1.7% vs. 5.7%). This trend may suggest that the leaching rates of metals in e-waste are not uniform, i.e., Zn exhibits a higher leaching rate than Al.

[0121] [Pilot-scale extraction of metals from waste electrical and electronic equipment]

[0122] The concentrations of metals in the analyte solution collected from the pilot-scale electrolytic cell are shown in Figure 21. As shown, all metal concentrations increased during the first hour of reaction time, gradually decreased after two hours, and then remained constant until four hours of reaction time. Cu and Fe were the most abundant elements measured in the post-leaching solution (reaching approximately 15 g / L and approximately 3 g / L, respectively, after four hours of leaching). The concentrations of other elements, including Zn, Al, and Ni, ranged from 0.3 to 2.2 g / L, while the concentration of Pb was less than 0.01 g / L. Overall, the elements measured in the post-leaching solution from the pilot-scale electrolytic cell followed the same distribution as that observed in the corresponding laboratory-scale electrolytic cell.

[0123] ORP is an important parameter in the leaching process because it indicates whether a solution has oxidizing or reducing power. Positive ORP values ​​suggest an oxidizing solution, while negative values ​​suggest a reducing solution. Figure 25 shows ORP as a function of reaction time for leaching processes carried out in laboratory-scale and pilot-scale electrolytic cells. In both cases, ORP showed an increase from approximately 400 to 1000 mV over the 4-hour reaction time, suggesting that the electrolytic cells produced an oxidizing solution for metal leaching.

[0124] Figure 22 shows the metal concentration profiles in the post-leaching solution for three batches of e-waste. The run time and e-waste input for each batch were 4 hours and 50 g, respectively, and 2.5 M sulfuric acid was used as the oxidant precursor for batch 1. At the end of each batch's processing, the residual solids were collected, and the solution was reused for metal extraction in batches 2 and 3.

[0125] Figure 24 shows the profile of copper recovery as a function of reaction time. Complete copper recovery (about 99%) was achieved in approximately 90 minutes of run time.

[0126] Cu was the most abundant element extracted in all batches, with approximately 15 g / L of copper extracted from each batch. Other extracted metals, including Fe, Al, Fe, Zn, Ni, and Pb, all followed the same trend as Cu. Approximately 1.3 g / L of Fe and Zn, approximately 0.35 g / L of Al, and approximately 0.16 g / L of Ni were leached in each batch. The metal concentration profiles presented in Figure 22 demonstrate that the electrolytic cell according to the present invention maintains the same performance in extracting metals for various input amounts of e-waste. This further demonstrates that the leach solution can be recycled multiple times with the same oxidizing power. Figure 26 shows the ORP profile for the leaching process under the conditions presented in Figure 19. The ORP showed an increase from about 400 to 1000 mV after 4 hours of reaction time in Batch 1, and this value leveled off at about 1000 mV when the solution was used to leach metals from the subsequent two batches of e-waste. The ORP profile shown in Figure 26 clearly demonstrates the ability of the electrolytic cell according to the present invention to regenerate oxidant for metal leaching, thus offering a promising green process for metal extraction from e-waste.

[0127] Figure 23 shows a commercially available electrolytic cell design for use with the method of the present invention. Multiple e-waste beds (43) can be installed in parallel with the analyte (4) output stream. E-waste extraction can proceed by passing the analyte stream through one e-waste bed for a period of time until most of the metals have been extracted. Once this is achieved, the analyte output stream can then be switched to a second e-waste bed for further metal extraction. While the electrolytic cell extracts metals from the second e-waste bed, the retentate from the e-waste bed can be collected for further processing. This step can be repeated for metal extraction from the e-waste bed. Following this approach, the electrolytic cell can also be used for continuous metal extraction using the same oxidant solution.

[0128] [Ore leaching]

[0129] Example 12: In this example, chalcopyrite was obtained and crushed to a D90 of 45 μm. The ore phases of the chalcopyrite were measured using an X-ray diffractometer (XRD) manufactured by Rigaku Smartlab. The target was CuKα, the tube current was 40 mA, the tube voltage was 40 kV, and the scan range 2θ was 15–80°. The sample composition was characterized using X-ray fluorescence (XRF) generated by a Rigaku Supermini200™.

[0130] Analysis was carried out to understand the effect of different temperatures on copper recovery from chalcopyrite. Different temperatures, 25°C, 35°C, and 45°C, were investigated. The effect of ultrasound and agitation speed on copper recovery was also carried out.

[0131] [Electrolytic cell apparatus for metal leaching from chalcopyrite]

[0132] An electrolytic cell for metal leaching from ore was constructed on a laboratory scale. The cell consisted of a working electrode (anode) and a counter electrode (cathode), with an anion exchange membrane separating the electrode compartments. The anode was made of BDD material, while a platinum plate served as the cathode. Experimental metal leaching from chalcopyrite ore was performed using a DC power supply with a current of 25–100 mA cm. -2 The test was carried out under a current density of .

[0133] Two vessels (pre-charged with 100 mL of solution) were used to recirculate the solution through the anode and cathode compartments separately. Both vessels contained an initial sulfuric acid concentration of 2.5 M, SO4 2- / Fe 2+ In a typical run, 2 g of chalcopyrite ore sample was added to the analyte vessel. A mechanical stirrer was used with stirring speeds of 400-1000 r / min to enhance mass transfer. The BDD anode was 2.25 cm 2 The platinum cathode is 1.5cm 2 It was.

[0134] The metal concentrations in the leach digestion solution were determined using an Aquaculture Photometer. A 3 M KCl electrode was used as the reference electrode. The electrolysis cell was placed in an ultrasonic water bath (Unisonics-FXP10™) with a power of 500 W and a frequency of 40 kHz.

[0135] Approach 1: Effect of temperature on electro-oxidation in separate cells for in situ leaching of metals.

[0136] In the first approach, electro-oxidation steps combining various temperatures were carried out in a laboratory-scale electrolytic cell with a separate structure to investigate the effect of temperature on copper conversion. In-situ leaching was performed to investigate the generation of peroxodisulfate oxidant from sulfate oxidation, as well as the generation of Fe in the analyte vessel. 2+and the immediate consumption of peroxodisulfate oxidant to oxidize a pre-charged solid metal sample (e.g., chalcopyrite ore).

[0137] Approach 1 was tested over a range of temperatures (25°C, 35°C, and 45°C) to observe the efficiency of copper conversion as a result of temperature. After 48 hours of operation, the efficiency of copper conversion during the electrolysis process was significantly affected by temperature. Figure 27 shows that increasing the temperature significantly improves the copper conversion efficiency over time. At 45°C, the copper conversion efficiency reaches approximately 84.75% after 48 hours, while at 25°C and 35°C, the conversion efficiencies are 25.79% and 35.69%, respectively.

[0138] In terms of surface reaction rate, higher temperatures help to accelerate the reaction of reactive oxygen species (ROS) and peroxodisulfate (PDS) with chalcopyrite.

[0139] In terms of oxidant generation, higher temperature conditions result in the production of hydroxyl radicals ( OH) and sulfate radicals (SO4 ·- The presence of more radical oxidants is essential to accelerate the leaching process.

[0140] Approach 2: Effect of mixing on electro-oxidation in separate cells for in situ leaching of metals.

[0141] In the second approach, the electro-oxidation step combined with various stirring speeds was carried out in a laboratory-scale electrolyzer with a separated structure to investigate the effect of stirring speed on copper conversion.

[0142] Approach 2 was tested over a range of stirring speeds (400-1000 r / min) to observe the efficiency of copper conversion as a result of stirring speed. After 48 hours of operation, the efficiency of copper conversion during the electrolysis process was significantly affected by stirring speed. Figure 28 shows that increasing the stirring speed significantly improves copper conversion efficiency over time. Conversion increased to 52.50% after 48 hours at 700 r / min, compared to 24.90% with no stirring at all.

[0143] By incorporating agitation into the reaction system, the mass transfer of the ore is significantly enhanced, allowing the ore to reach the ROS reaction zone on the BDD. This promotes the participation of ROS in the leaching process. A significant increase in leaching rate is attributed to the participation of ROS, as shown in the data.

[0144] Furthermore, Figure 29 confirms that varying the agitation speed between 400 and 1000 rpm does not significantly affect the leaching conversion, suggesting that agitation speeds above 400 rpm are sufficient to eliminate mass transfer limitations between the ore particles and the oxidant for the radical oxidant. This indicates a pivotal shift in the limiting step from mass transfer to ROS generation when a certain agitation speed is introduced into the system.

[0145] Approach 3: Effect of UV on electro-oxidation in separate cells for in situ leaching of metals.

[0146] In the third approach, an electro-oxidation step combined with the use of UV was carried out in a laboratory-scale electrolyzer with isolated structure to investigate the effect of UV on the copper conversion rate.

[0147] Approach 3 was tested using UV to observe the efficiency of copper conversion rate as a result of placing the electrolysis cell in an ultrasonic water bath at a power of 500 W and a frequency of 40 KHz. After 48 hours of operation, the efficiency of copper conversion during the electrolysis process was significantly affected.

[0148] The effectiveness of copper conversion from chalcopyrite in an electroleaching system is significantly enhanced by ultrasound, which primarily focuses on enhancing mass transfer within the system. Figure 28 outlines the clear enhancement in conversion rate when this mechanism is employed. At 24 hours, the conversion rate rises to 80.73%, compared to 24.90% using electrooxidation alone.

[0149] Example 13: Analysis was conducted to understand the recovery of leached metals from two high quality nickel ores, Nickel Ore 1 and Nickel Ore 2. Nickel Ore 1 and Nickel Ore 2 both had particle sizes smaller than 3.35 mm. Leached metals were investigated, including nickel, iron, copper, cobalt, manganese, and chromium.

[0150] [Electrolytic cell equipment for metal leaching from nickel ore]

[0151] An electrolytic cell for metal leaching from ore was constructed on a laboratory scale. The electrolytic cell consisted of a working electrode (anode) and a counter electrode (cathode), with an anion exchange membrane separating the electrode compartments. The anode was made of BDD material, while a stainless steel plate served as the cathode. Experimental metal leaching from nickel ore was performed using a 150 mA cm from a DC power supply. -2 The test was carried out under a current density of .

[0152] Two vessels (pre-charged with 120 mL of solution) were used to recirculate the solution through the anode and cathode compartments separately. Both vessels contained an initial sulfuric acid concentration of 2.5 M, SO4 2- / Fe 2+ = 1 / 0.1 oxidant precursor solution was added. In a typical run, a 5g nickel ore sample was added to the analyte vessel.

[0153] [Metal extraction from nickel ore on a laboratory scale]

[0154] Figures 30 and 31 show the profiles of metal concentrations in the post-leaching solution as a function of reaction time for both Nickel Ore 1 and Nickel Ore 2, respectively. Leaching was carried out in a laboratory-scale electrolytic cell at 150 mA cm -2 current density, SO4 2- / Fe 2+ = 1 / 0.1, and flow rate = 50 mL min -1 The elements measured in the solution included Ni, Cu, Co, Fe, Mn, and Cr.

[0155] As shown in Figure 30, the concentrations of all metals from Nickel Ore 1 increased over the 10-hour reaction. Ni and Fe were the most abundant elements measured in the post-leaching solution (reaching approximately 140 mg / L and approximately 1050 mg / L, respectively, after 4 hours of leaching). The concentrations of other elements, including Co, Co, Mn, and Cr, ranged from 2 to 20 mg / L.

[0156] Figure 31 shows that all metal concentrations from Nickel Ore 2 increased over the 10 hour reaction, however, again Ni and Fe were the most abundant elements in the post-leaching solution. In both cases, Ni and Fe would be expected to be the most abundant elements in the raw ore.

[0157] Example 14: Analysis was performed to understand the recovery of leached metals from vanadium-bearing ores. Leached metals were investigated, including vanadium, iron, chromium, nickel, cobalt, and manganese. Ore characterization was performed using aqua regia microwave digestion at 180°C, 3 bar, for 30 minutes.

[0158] [Electrolytic cell equipment for metal leaching from vanadium ore]

[0159] An electrolytic cell for metal leaching from ore was constructed on a laboratory scale. The electrolytic cell consisted of a working electrode (anode) and a counter electrode (cathode), with an anion exchange membrane separating the electrode compartments. The anode was made of BDD material, while a stainless steel plate was used as the cathode. Experimental metal leaching from vanadium ore was performed using a 150 mA cm from a DC power supply. -2 The test was carried out under a current density of .

[0160] Two vessels (pre-charged with 120 mL of solution) were used to recirculate the solution through the anode and cathode compartments separately. Both vessels contained an initial sulfuric acid concentration of 2.5 M, SO4 2- / Fe 2+ An oxidant precursor solution with a r = 1 / 0.1 was added. A 10% solids concentration was used. Elements measured in the leached solution included V, Fe, Cr, Co, Mn, and Ni.

[0161] [Laboratory-scale extraction of vanadium from ore]

[0162] Figure 32 shows the improvement in metal recovery after 2 hours of treatment in an electrolytic cell with analyte solution versus a standard acid leach using only 2.5M H2SO4. Significant improvements in metal recovery are evident, primarily for V, Fe, and Cr (1000% improvement in recovery for V, 820% for each of Fe and Cr).

[0163] Example 15: The above findings can be extended to achieve improved metal recovery in industrial-scale recovery of metals from ores, landfill residues, sludge, tailings, slag, ash, filter dust from incinerators, blanks, e-waste, or other metal-containing wastes. A specific example is the recovery of metals from low-grade ores or tailings in the mining sector. These hydrometallurgical processes primarily include heap leaching, in-situ leaching, and tank leaching.

[0164] Approach 1: Use of analyte solutions to accelerate metal recovery in heap leaching.

[0165] In this example, the analyte solution may be generated on-site and supplied to replace reagents, such as acid sulfates or alkali carbonates, in heap leaching of ore and tailings at a mine. In heap leaching, mined ores, such as precious metals, copper, nickel, and uranium, are crushed and placed on a leach pad lined with impermeable plastic or clay. The heap is washed with a leach solution to dissolve the metals within the ore, which are then recovered. The analyte solution may be used as the leach solution in place of standard acid or alkali reagents for washing tailings or low-grade ore heaps, thus leaching the metals into a pregnant liquor. The pregnant liquor containing the dissolved metals is then processed for metal recovery by standard methods. Figure 33 shows a standard process for heap leaching of copper from low-grade ore. An improvement over the standard process is increased efficiency of metal recovery, for example by using analyte solutions instead of acid sulfate, as depicted for example in the case of copper in Figure 28 and for V, Fe, Cr, Co, Mn, Ni in Figure 32.

[0166] Approach 2: Use of analyte solutions to accelerate metal recovery in tank or vat leaching.

[0167] In this example, the analyte solution may replace standard leaching reagents (e.g., sulfuric acid) in a tank or vat leaching process. Tank and vat leaching involve placing ores, or other solid materials containing metals, usually after crushing and sorting, into large tanks, where they are then either submerged in a leach solution (in the case of vat leaching) or crushed and mixed with water to form a slurry before the leach reagent is added (in the case of tank leaching).

[0168] In some cases, the tanks are fitted with agitators and baffles to keep the solids suspended and therefore improve the efficiency of metal extraction. Stirred tanks are closed cylindrical vessels, either vertical or horizontal, with power-driven paddles or agitators on a vertical or horizontal axis. They are equipped with a facility at the bottom for withdrawing the leach solution at the end of the run. In some designs, a horizontal drum is the extraction vessel, and the solids and liquid are tumbled inside by the rotation of the drum on rollers. They are operated on a batch basis, each being a single leaching stage. They can also be used in series for multi-stage operations.

[0169] For the leaching of finely divided solids, pachuca tanks were used, which are widely used in the metallurgical industry. These tanks are constructed of wood, metal, or concrete and are lined with the appropriate material depending on the characteristics of the leachate. Agitation is achieved by airlift: bubbles rising through a central tube cause an upward flow of liquid and suspended matter within the tube, thus circulating the mixture. Conventional mechanical agitation is also used for this purpose.

[0170] Once the desired leaching is achieved, agitation is stopped, the solids are allowed to settle, and the supernatant liquid is decanted by siphoning from the top of the tank or by withdrawing it through a drain placed at an appropriate height on the side of the tank. As the solids settle, forming a compressible sludge, the remaining solution becomes more abundant, and typically the last traces of solute are recovered in a countercurrent fashion. In a new embodiment, the analyte solution acts as an electrolyte, improving the efficiency of metal recovery relative to the use of acid sulfate, as depicted, for example, in Figure 28 for copper and Figure 32 for V, Fe, Cr, Co, Mn, and Ni. Further improvements are achieved through optimization, including temperature, agitation speed, and the use of UV light, as shown in Figures 27 and 28.

[0171] Approach 3: Use of analyte solutions to accelerate metal recovery in countercurrent separation.

[0172] In a continuing variation of the tank and vat process, countercurrent leaching is used to more efficiently extract metals from materials such as ores, concentrates, slags, muds, chips, and metal wastes (and other terms used in the industry to describe wastes containing metals). While the principles of solvent strength, contact time, surface area, temperature, and flow rate all remain key components of the process, there are many variables in this approach.

[0173] Important factors in CCE extraction include what the target metal is for extraction and what materials are contained within it. In addition, how the target is prepared will also be a factor in the successful operation of many CCE processes. For example, lithium must be calcined and crushed prior to initiating the leaching process, while ores such as copper and nickel must be crushed to a uniform size for extraction.

[0174] While there are many forms of CCE, ranging from very simple siphoning processes in tanks containing the ore to very sophisticated agitation systems and tank arrangements known as shank systems (Figure 34), the concept of CCE remains the same. As shown in Figure 34, the solids or slurry move countercurrently to the liquid flow (individually, in the direction of the arrows from (a) to (h)), with emphasis on effective solvent selection.

[0175] Sometimes the pressure drop of the liquid flow due to gravity is high or the volatility of the solvent is high. In such circumstances, the liquid is pumped through a fixed bed in a vessel called a diffuser. The main advantage of these units is that they prevent evaporation loss of the solvent when operated above the boiling point of the solvent. The use of a leaching tank is suitable for this operation, although it is not appropriate when the ore or target is in very fine form. In such circumstances, the solids can be filtered and leached in a filter press by pumping the solvent through the press cake.

[0176] The Rotorcell Extractor, depicted in Figure 33, is an improvement over the Shank system, in which the leaching tank is continuously moved, allowing for continuous loading and unloading of solids. It consists of an annular shell divided into several cells, each fitted with a hinged screen bottom to support the solids. This shell rotates slowly above a stationary compartmented tank. As the rotor rotates, each cell passes in turn under a prepared solids feeder and then under a series of sprayers, which periodically immerse the contents of each cell in solvent for leaching. When leaching is expected to be complete by the end of one rotation, the leached solids from each cell are automatically dumped into the stationary compartment below, from which they are continuously transported. The solvent sprayed into each solid-filled cell percolates down through the solids and support screen into the appropriate compartment of the lower tank, from which it is pumped to the next sprayer. The leaching is countercurrent, with the strongest solution coming from the cell filled with virgin solids. An improvement over the standard process is the increased efficiency of metal recovery through the use of an analyte solution. The improvement over sulfuric acid is depicted for copper in Figure 30.

[0177] While the invention has been described in connection with specific embodiments thereof, it will be understood that it is capable of further modifications. This application is intended to cover all such variations or adaptations of the invention generally in accordance with the principles of the invention, including such departures from the present disclosure as are known or customary in the art to which the invention pertains and as may conform to the essential features described above.

[0178] Since the present invention can be embodied in several forms without departing from the spirit of the essential features of the present invention, it should be understood that the above-described embodiments do not limit the present invention unless otherwise specified, but rather should be broadly interpreted within the spirit and scope of the present invention as defined in the appended claims. The described embodiments are to be considered in all respects only as illustrative and not restrictive.

[0179] Various modifications and equivalent arrangements are intended to be included within the spirit and scope of the invention and the appended claims. The specific embodiments are, therefore, to be understood to be illustrative of the many ways in which the principles of the invention may be practiced.

[0180] When used herein, the terms "comprises / comprising" and "includes / comprising" are to be interpreted as specifying the presence of stated features, integers, steps, or components, but not as excluding the presence or addition of one or more other features, integers, steps, components, or groups thereof. Thus, unless the context clearly requires otherwise, throughout the specification and claims, words such as "comprises," "comprising," "includes," "including," and the like are to be interpreted in an inclusive sense, i.e., "including, but not limited to," as opposed to an exclusive or exhaustive sense.

[0181] When a Markush group or other grouping is used herein, all individual members of the group, and all possible combinations and subcombinations of the members of the group, are intended to be included individually in the disclosure. All combinations of the components described or exemplified herein may be used to practice the invention, unless otherwise specified.

[0182] Whenever a range is given herein, e.g., a temperature range, a time range, or a composition or concentration range, all individual values ​​included in the given range, as well as all intermediate ranges and subranges, are intended to be included in the disclosure, and it will be understood that any subrange or individual value within a range or subrange included herein may be excluded from the claims herein.

[0183] One skilled in the art will understand that materials and methods other than those specifically exemplified may be employed in the practice of the present invention without resorting to undue experimentation. All art-known functional equivalents of any such materials and methods are intended to be encompassed by the present invention. The terms and expressions employed are used as terms of description and not of limitation, and there is no intention to use such terms and expressions to exclude any equivalents of the features shown and described, or portions thereof, and it is recognized that various modifications are possible within the scope of the invention as claimed. Therefore, while the present invention has been specifically disclosed by examples, preferred embodiments, and optional features, it should be understood that modifications and variations of the concepts disclosed herein may be practiced by those skilled in the art, and that such modifications and variations are considered to be within the scope of the present invention, as defined by the appended claims. [Explanation of symbols]

[0184] [Parts List] The drawings and specification herein should be read with reference to the following item numbers appearing in the drawings and text:

[0185] 1...Anode, 2...Cathode, 3...Membrane, 4...Analyte, 5...Catholite, 6...DC Power Supply, 7...Electrolyzer, 8...Peristaltic Pump, 9...Electro-oxidation Anode Cell, 10...Electro-reduction Cathode Cell, 11...Annular Cell, 12...Cell Effluent, 13...Cell Feed, 14...Electrolyte Pump, 15...Circulation Tank, 16...Electro-oxidation, 17...Metal Leaching, 18...Metal Recovery, 19...Recovered Metal, 20...Spent Solution, 21...Particulate E-waste, 22...Metal Precipitation, 23...Metal-Rich Solution, 24...Hydrogen Generation, 25...Oxidant, 26...Solution Application, 27...Leaching Solution (Lixivant) Solution), 28...Ore, 29...Crushing and Agglomeration, 30...Heap, 31...Pregnant Pond, 32...Unused Solids, 33...Concentrated Extract, 34...Spent Solids, 35...Interstage Solution, 36...Unused Solvent Sprayer, 37...Solid Feed, 38...Solid Discharge, 39...Extract, 40...Solvent / Solution Spray Pump, 41...Sprayed Solvent / Solution, 42...Solids, 23...E-Waste Bed

Claims

1. 1. A method for producing an oxidant solution using an electrochemical cell having an anode and a cathode, comprising: (i) supplying a raw electrolyte solution to a reaction zone between the anode and the cathode, wherein the raw electrolyte solution contains sulfate ions (SO ４ ２－ ) and divalent iron ions (Fe ２＋ ) and (ii) In the operation cycle, the raw electrolyte solution is electrolyzed to produce peroxodisulfate (S ２ O ８ ２－ ) and trivalent iron ions (Fe ３＋ ), and (iii) providing the oxidizing acid solution; A method comprising:

2. The sulfate ions (SO ４ ２－ 2. The method of claim 1, wherein the concentration of ) is between 0.1 molar and 5 molar.

3. The raw electrolyte (Fe ２＋ ) in the divalent iron ion (Fe ２＋ 3. The method of claim 1, wherein the concentration of ) is between 0.1 and 0.5 molar.

4. The SO ４ ２－ : The Fe ２＋ 10. The method of any one of the preceding claims, wherein the ratio of is between 1:0.05 and 1:0.

5.

5. The method of claim 1 , wherein the anode comprises boron-doped diamond.

6. 10. A method for producing an oxidant solution using an electrochemical cell according to any one of the preceding claims, comprising an anode half-cell having an anode and a cathode half-cell having a cathode, the method comprising: (i) supplying the raw electrolyte solution flowing into and out of a fluid path formed between the anode and the cathode to the anode half-cell, wherein the raw electrolyte solution contains sulfate ions (SO ４ ２－ ) and divalent iron ions (Fe ２＋ ) and (ii) In the operation cycle, the raw electrolyte solution is electrolyzed to produce peroxodisulfate (S ２ O ８ ２－ ) and trivalent iron ions (Fe ３＋ forming an analyte solution comprising: (iii) dispensing the analyte solution from the anode half-cell; A method comprising:

7. The method of claim 6 , wherein a separator is positioned between the anode half-cell and the cathode half-cell.

8. 10. A method for producing an oxidant solution using an electrochemical cell according to any one of the preceding claims, comprising the steps of: (iii) the trivalent iron ion (Fe ３＋ ) is the divalent iron ion (Fe ２＋ supplying said oxidizing acid solution to the metal-containing waste so that said metal-containing waste is reduced to (iv) the peroxodisulfate (S ２ O ８ ２－ ) to the divalent iron ions (Fe ２＋ ) to contact the iron (III) ions (Fe ３＋ ) and the sulfate ion (SO ４ ２－ ) and and a further step of regenerating the oxidant solution by

9. 10. A system for generating an oxidizer solution according to any one of the preceding claims, comprising: (a) an anode and a cathode defining a reaction area of ​​an electrochemical cell; (b) an inlet port and a flow controller for passing an aqueous starting electrolyte between the electrodes, the starting electrolyte containing sulfate ions (SO ４ ２－ ) and divalent iron ions (Fe ２＋ ) wherein the sulfate ions and the divalent iron ions are peroxodisulfate (S ２ O ８ ２－ ) and trivalent iron ions (Fe ３＋ the inlet port and the flow controller are selected for electrochemical generation of an oxidant comprising (c) current means for supplying a current for electrolyzing the aqueous raw electrolyte solution to produce the oxidant solution containing the oxidant in the reaction region; (d) an outlet port for delivering the oxidant solution from the electrochemical cell; A system comprising:

10. 9. A system for on-site generation of oxidizer solution according to any one of claims 1 to 8, comprising: (a) an anode half-cell and a cathode half-cell; (b) an inlet port and flow controller for the passage of aqueous feed electrolyte through the anode half-cell, wherein the feed electrolyte contains sulfate ions (SO ４ ２－ ) and divalent iron ions (Fe ２＋ ) wherein the sulfate ions and the divalent iron ions are peroxodisulfate (S ２ O ８ ２－ ) and trivalent iron ions (Fe ３＋ the inlet port and the flow controller are selected for electrochemical generation of an oxidant comprising (c) current means for supplying a current for electrolyzing the aqueous feedstock electrolyte to produce an oxidant solution containing the oxidant in the anode half-cell; (d) an outlet port for supplying the oxidant solution from the electrochemical cell from the anode half-cell; A system comprising:

11. The system for generating an oxidant solution according to claim 10 , wherein a separator is positioned between the anode half-cell and the cathode half-cell.

12. 12. The system of any one of claims 9 to 11, wherein the oxidant solution is used for metal extraction, preferably for the extraction of one or more metals selected from the group comprising Cu, Co, Ni, V, Cr, Mn, Fe, Au, Ag, Pt, Pd, Rh, Ir, Ru, Te, Ga, Se, Ta, Ge, Pb, Cd, In, and Sb.

13. The system of claim 12, wherein the metals are extracted from e-waste or minerals.

14. 1. A method for leaching metals from metal-bearing waste, comprising: providing an oxidizer solution produced by the method of any one of claims 1 to 7; contacting the oxidant solution with the metal-containing waste; A method comprising:

15. 15. The method of claim 14, wherein the metal to be leached is selected from the group consisting of Cu, Co, Ni, V, Cr, Mn, and Fe.

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