Separation process for the recovery of uranium from wet-process phosphoric acid

#Wet-process phosphoric acid contains a significant amount of uranium, typically about 0.1 to 0.2 g/l. Apart from being economically favourable, the recovery of uranium as a by-product of phosphate fertilizer production supplies part of the world's nuclear fuel resource and reduces the environmental contamination. In order to recover the valuable source of uranium, a so-called "D.sub.2 T-D.sub.2 T process" has been developed by a research team at the Oak Ridge National Laboratory in the United States to separate it from a phosphoric acid solution prior to further treatment to produce a fertilizer product. (See U.S. Pat. No. 3,711,591 by F. J. Hurst et al. titled "Reductive Stripping Process for the Recovery of Uranium from Wet-process Phosphoric Acid". Also, see the article in Ind. Eng. Chem. Process Des. Develop., 11, 122 (1972) by F. J. Hurst et al. titled "Recovery of Uranium from Wet-process Phosphoric Acid").

#The D.sub.2 T-D.sub.2 T process comprises two steps. A synergistic extractant combination consisting of di(2-ethylhexyl) phosphoric acid (D2EHPA) and trioctylphosphine oxide (TOPO) in kerosene is used in this process to extract uranium in the +6 oxidation state in the first step. In the second step of the process, the concentrated acidic aqueous stream is again oxidated and extracted with the same synergistic organic extractant, and finally stripped with ammonium carbonate solution to precipitate ammonium uranyl tricarbonate, (NH.sub.4).sub.4 UO.sub.2 (CO.sub.3).sub.3, (AUT).

#This process typically suffers from several disadvantages: (1) The TOPO synergistic agent is expensive and hence the total chemical reagent cost is higher. (2) In the second purification part of D.sub.2 T-D.sub.2 T process, the use of ammonium carbonate as the stripping reagent results in the conversion of the D2EHPA to a highly hydrated ammonium salt, ##STR1## When the stripped solvent is recycled to the extraction step, the extracted water and NH.sub.4.sup.+ transfer to the aqueous phosphoric phase, dilute and contaminate the wet-process phosphoric acid, which is unacceptable to the phosphoric acid plant to produce feed stocks and various phosphate chemicals. These in turn have some other disadvantageous effects such as unclear phase separation, ammonium salt precipitation [Fe.sub.3 NH.sub.4 H.sub.8 (PO.sub.4).sub.6.6H.sub.2 O etc.] and cause problems during operation.

#The invention will be described in detail with reference to FIG. 1 which is a schematic diagram illustrating the process of the present invention.

#FIGS. 2 and 3 are material balance diagrams for the first and second cycles, respectively.

#After pretreatment for removing suspended solid and organic matter from wet-process phosphoric acid, the uranium in the feed acid is oxidized to the +6 oxidation state by adding a suitable amount of H.sub.2 O.sub.2. which is dependent on the EMF or Fe.sup.2+ concentration in the feed acid, sometimes this step is even unnecessary. Referring to FIG. 1, feed acid from line 1 enters the extraction section 2 (5 to 8 stages). In the extraction section, the feed acid is countercurrently mixed and settled with the organic extractant from line 3 which reacts with the uranyl ions to form a complex soluble in the solvent. The composition of the feed acid varies with different batches, it generally ranges from 4 to 6M in H.sub.3 PO.sub.4 and contains 60 to 120 parts per million (ppm) by weight of uranium. The organic extractant used in this invention consists essentially of di(2-ethylhexyl) phosphoric acid (D2EHPA) and dibutyl butyl phosphonate (DBBP) dissolved in kerosene. In the first cycle said organic extractant is 0.5.about.0.7M D2EHPA-0.1.about.0.14M DBBP-kerosene (The optimum D2EHPA:DBBP mole ratio is 5:1). The ratio of the feed acid to organic extractant can be adjusted from 2:1 to 4 :1 which is dependent on the phosphoric acid concentration and uranium concentration of the aqueous phase. The uranium extraction coefficients decrease as the temperature of the acid is increased. The operation temperature in the extraction section is better kept under room temperature. The organic extractant, containing complexed uranium, passes through line 4 to the reductive stripping section 5. To a portion of the raffinate from the extraction section 2 is added scrap iron (Fe.degree.) to reduce ferric ions and to bring the ferrous ion concentration up to 25 to 50 grams per liter and passes through lines 6 and 7 to mix with the organic extractant. The uranyl ion is reduced to the quadravalent U.sup.4+ ion and therefore stripped to the aqueous stream due to the fact that the U.sup.4+ ion is not complexed by D2EHPA-DBBP. The volumetric flow ratios of the uranium-loaded organic extractant stream 4 to strip liquor stream 7 is 35/1 to 50/1. Intrastage recycle of the aqueous phase from the settlers to the mixers is provided to provide an aqueous/organic phase ratio of about 2 in the mixers-settlers and an operation temperature of 45.degree..about.55.degree. C. to ensure efficient mixing and stripping of the uranium. The organic extractant leaving the reductive stripping section is then recycled through line 3 to extraction section 2. The product solution from the first concentration cycle contains about 6 to 15 g/l uranium and about 25 to 50 g/l iron.

#The quadravalent U.sup.4+ ion in the first cycle product solution is oxidized by H.sub.2 O.sub.2 to the uranyl ion before entering extraction section 9 (3 to 6 stages) through line 8. The amount of oxidant added is dependent upon the Fe.sup.2+ concentration in the acid. In the extraction section, the oxidized acid is contacted with the organic extractant from line 10 containing 0.3.about.0.5M D2EHPA-0.06.about.0.1M DBBP-kerosene. The volumetric flow ratio of aqueous phase to organic phase is about 1 and the concentration of the organic extractant is lower than that used in the first cycle. The main purpose is to obtain a high purity product by increasing the saturation capacity of the extractant for uranium and minimizing the extraction of impurities. The operation temperature of the extraction section is preferably kept at room temperature. More than 99% of the uranium is extracted into the organic phase. In addition to uranium, the extract contains 0.5.about.1.5 g/l phosphate and 0.10.about.0.20 g/l iron. Organic phase from line 11 and water from line 13 enter three-stage water scrubbing section. More than 98% of the phosphate in the organic phase is removed under operation at an organic/aqueous ratio of 2/1 to 5/1. The uranium-loaded organic extractant stream 14 then enters the ammonium carbonate stripping section 15. Almost all of the uranium is removed from the organic phase with ammonium carbonate solution in two stages. Make-up ammonia and carbon dioxide are bubbled into the aqueous phase of the settler in the first stage to maintain the concentration of ammonium carbonate at about 2 to 2.5 M. Dilute ammonium carbonate solution of about 0.5 M is fed at a slow flow rate to the second stripping stage to scrub the very small amount of entrained uranium from the organic phase and to compensate for the loss of water from the stripping system, which occurs on conversion of the D2EHPA to the hydrated ammonium salt. Calcination of the AUT precipitate for two hours at 500.degree..about.800.degree. C. yields a product containing more than 99% of U.sub.3 O.sub.8.

#The stripped organic extractant leaving the ammonium carbonate stripping section 15 contains 7.2.about.8.3 g/l NH.sub.4.sup.+ and 65.about.80 g/l H.sub.2 O. The NH.sub.4.sup.+ and H.sub.2 O in the organic phase are scrubbed with the sulfuric solution from line 18 at an organic/aqueous ratio of 2/1 to 5/1. The concentration of the feed sulfuric acid solution in the two-stage scrubbing section is preferably 6N to 16 N. The sulfuric acid solution leaving the second scrubbing stage is recycled to the first stage. The scrubbing efficiency is less than 99% when the H.sup.+ concentration of the recycled sulfuric acid solution gradually decreases to below 0.6N. The sulfuric acid scrubbing section should be designed so that it can be by-passed and replaced with fresh solution when necessary. Operating temperatures of 50.degree..about.60.degree. C. are preferred for efficient phase separation. After sulfuric acid scrubbing, organic extractant contains about 100 ppm of SO.sub.4.sup.2-. The concentration of SO.sub.4.sup.2- is reduced to below 10 ppm when scrubbed with water from line 21. No precipitate is found in the extraction section when the recycled organic extractant is scrubbed with sulfuric acid solution and water. Also, phase separation is improved in the second cycle extraction step, dilution and contamination of the wet-process phosphoric acid are prevented. As described above, the present process for the recovery of uranium from wet-process phosphoric acid has the following advantages :