4,302,427 Recovery of uranium from wet-process phosphoric acid

Patent Number/Link:

Un.ited States Patent [19]

Berry et at

[11]

[45]

4,302,427

Nov. 24, 1981

[54] RECOVERY OF URANIUM FROM

WET-PROCESS PHOSPHORIC ACID

OTHER PUBLICATIONS

Caropreso et a!', "Hydrogen Peroxide Precipitation of

Uranium ... ", Trans. Soc. Mining Eng. AIME 254(4),

pp. 281-284 (1973).

[75] Inventors: William W. Berry, Lakeland, Fla.;

Angus V. Henrickson, Golden, Colo.

[73] Assignee: International Minerals & Chemical

Corporation, Northbrook, Ill.

Appl. No.: 22,079

[57] ABSTRACT

14 Claims, 1 Drawing Figure

Galkin et a!., Technology of Uranium, AEC-tr-6638, pp.

170-187, 1964.

Harrington et aI., Eds., Uranium Production Technology,

Van Nostrand Company, Inc., Princeton, N.J., 1959,

pp. 157-159.

Hurst, "Recovery of Uranium from Wet-Process Phosphoric

Acid" I & EC Process Design and Development,

vol. II, No.1, pp. 122-128 (1972).

Shabir, U.S. Bureau of Mines Report of Investigations-

RI 7931 (1974).

Hurst, "Recovering Uranium from Wet-Process Phosphoric

Acid", Chem. Eng. (Jan. 3, 1977) pp. 56-57.

Hurst, ORNL-TM-2522 (1960).

Hurst, ORNL-2952 (1960).

Bunus, "Synergistic Extraction Of Uranium ... ", J.

Inorg. NucL Chem., vol. 40, pp. 117-121 (1977).

Primary Examiner-Deborah L. Kyle

Attorney, Agent, or Firm-Schuyler, Banner, Birch,

McKie & Beckett

Uranium values are recovered as uranyl peroxide from

wet process phosphoric acid by a solvent extractionprecipitation

process. The preferred form of this process

comprises a first solvent extraction with DEPATOPO

followed by reductive stripping of the extractant

with Fe+ + - containing phosphoric acid. After reoxidation,

the uranium-containing aqueous stripping solution

is extracted again with DEPA-TOPO and the pregnant

organic is then stripped with a dilute ammonium carbonate

solution. The resulting ammonium uranyl tricarbonate

solution is then acidified, with special kerosene

treatment to prevent wax formation, and the acidified

solution is reacted with H202 to precipitate a uranyl

peroxide compound.

Filed: Mar. 19, 1979

References Cited

U.S. PATENT DOCUMENTS

2,819,145 111958 McCollough et al. 423/10

2,859,094 1111958 Schmitt et al. 423/10

3,052,513 9/1962' Crouse 423/10

3,052,514 9/1962 Schmitt 423/10

3,100,682 8/1963 Kelmers 423/8

3,711,591 111973 Hurst et al. 423/10

3,737,513 6/1973 Wiewiorowski 423/8

3,835,214 9/1974 Hurst et al. 423/10

3,966,872 6/1976 Sundar : 423/9

3,966,873 6/1976 Elikan et al. 423/10

4,002,716 111977 Sundar 423/10

4,024,215 5/1977 Caropreso et al. 423/260

4,087,512 5/1978 Reese et al. 423/321 R

4,105,741 8/1978 Wiewiorowski 423/10

Int. Cl.3 COlG 43/01; BOlD 11/04

U.S. Cl. 423/10; 210/634;

423/16; 423/253; 423/260

Field of Search 423/8, 10, 11, 260,

423/261,16,253;210/21

[56]

[21]

[22]

[51]

[52]

[58]

u.s. Patent Nov. 24, 1981 4,302,427

PHOSPHORIC ACID

H2 02

1---;., /2

3",

PURl FICATION

4-....

,., ,., 7'\ I .. ACID RAFFINATE

6\ ~ PRIMARY

OXIDATION +6· SOLVENT EXTRACTANT U EXTRACTION -

\.5 '8

H3P04 (Fe++)

O2 \.10 + ,IV ,

12'\

REDUCTIVE

1-11

OXIDATION -- +4 STRIPPING /13 U ..

14/

15'\ .. SECONDARY EXTRACTANT SOLVENT - U+6

EXTRACTION '16

.. ACID RAFFINATE

H2O 1--18

(NH 4}2C03

26- H2SO4

21~

1--22

,. ,r /19 " \24

DILUTE -~25

SCRUBBING ... ALKALINE

ACID

Jo STRIPPING

SCRUBBING

H2 SO4 23\

' ~

28- ' ,

RGANIC ,..-27

34\ (NH4 )2 S04 OLVENT

30--1· ,r r

... /

31\ NH3 29-

ACI 01 FICATION PRECIPITATION ... H2 0 2

, \33

WAX/SOLVENT

r--35

URANYL PEROXIDE

4,302,427

RECOVERY OF URANIUM FROM WET-PROCESS

PHOSPHORIC ACID

The mining of phosphate rock such as that found in 5

Florida and many countries of the world, e.g., Morocco

has as its prime objective the production of phosphatecontaining

fertilizer. In one widely used process, the

phosphate values are recovered from the rock by digestion

with sulfuric acid to produce a phosphoric acid 10

solution (called wet-process phosphoric acid) and an

insoluble calcium sulfate (gypsum). Phosphate rock

may contain significant quantities of uranium, e.g., on

the order offrom about 0.1 to 0.5 pounds of uranium per

ton of phosphate rock mined and more generally within 15

the range of about 0.2 to about 0.4 pounds per ton.

During the digestion step the uranium values are solubilized

resulting in a uranium concentration (expressed as

U30S) in the wet-process phosphoric acid of from about

0.05 to about 0.3 grams per liter and more generally 20

from about 0.15 to about 0.25 grams per liter.

Attempts to recover uranium values from wet-process

phosphoric acid have centered on the use of solvent

extraction processes in which the uranium values are 5

transferred to an organic phase, stripped from the or- 2

ganic phase and subsequently recovered as a uranium

precipitate. The uranium-free wet-process phosphoric

acid is then processed conventionally to form various

phosphate-containing fertilizer products. 30

The uranium solvent extraction process which has

generated the most commercial interest is the so-called

reductive stripping extraction. process developed by

Oak Ridge National Laboratories (ORNL). See Hurst,

U.S. Pat. No. 3,711,591; Hurst, et al, "Recovery of 35

Uranium from Wet-Process Phosphoric Acid", 1& EC

Process Design and Development, Vol. II, p. 122-128,

January 1972; Hurst et aI, "Recovering Uranium from

Wet-Process Phosphoric Acid", Chemical Engineering,

Jan. 3, 1977, p. 56-57; See also, Hurst et aI, ORNL-TM- 40

2522 (1969) and Hurst et aI, ORNL-2925 (1960). The

ORNL process as described in the Hurst '591 patent

employs a synergistic extraction mixture of di(2-ethylhexyl)

phosphoricacid (DEPA) and trioctylphosphine

oxide (TOPO) dissolved in an organic diluent. This 45

extraction mixture is known to have a high affmity for

uranium in the hexavalent oxidation state.

In the process described in the Hurst patent, wetprocess

acid is first subjected to oxidation conditions to

convert the uranium to the hexavalent state and then 50

contacted in a first extraction cycle with DEPATOPO.

The pregnant organic phase is then treated with

an aqueous phosphoric acid stripping solution containing

ferrous ions in an amount sufficient to reduce the

uranium values in the mixture to the quadrivalent state. 55

In the quadrivalent form the uranium values have substantially

no affinity for the organic phase and are transferred

into the aqueous strip solution. After reoxidation

of uranium values to the hexavalent state, the acidic

strip solution is fed to a second extraction cycle where 60

it is contacted with a more dilute DEPA-TOPO extractant

to retransfer the concentrated uranium values

into the organic phase. After scrubbing to remove entrained

aqueous materials, the pregnant organic extractant

is stripped with an ammonium carbonate solution 65

which causes transfer of uranium values to the aqueous

phase and simultaneous precipitation of insoluble ammonium

uranyl tricarbonate (AUT). The resulting

AUT slurry is then filtered, washed and calcined to

produce a dry uranium concentrate.

One of the problems associated with the ORNL process

is metallic contamination resulting from the coextraction

and coprecipitation of metals such as iron present

in the orginal acid. For nuclear fuel applications

high amounts of iron contamination are intolerable.

Sundar, U.S. Pat. No. 4,002,716 attempts to obviate the

problem of iron contamination by stripping the uranium

values from the secondary cycle organic extractant

with a dilute ammonium carbonate solution containing

sulfide ions. Under these conditions the AUT complex

remains soluble and iron values are said to precipitate as

iron sulfides. After filtration of the iron sulfides, the

AUT complex is treated with additional quantities of

ammonia and carbon dioxide to raise the ammonium

carbonate concentration to a level at which AUT precipitation

occurs. This AUT product can then be

washed and calcined according to the ORNL process.

Applicants have surprisingly discovered that uranium

having reduced metallic contamination can be more

efficiently recovered from wet-process phosphoric acid

by modifying the ORNL process to provide a dilute

carbonate strip solution which can be subsequently

acidified and reacted with H202 to precipitate uranium

peroxide.

The present invention provides a process for recovering

uranium from wet-process phosphoric acid containing

hexavalent uranium values which comprises the

steps of: (a) contacting said acid with an organic extractant

comprising a mixture of di(2-ethylhexyl) phosphoric

acid and trioctylphosphine oxide in a phosphoric

acid-immiscible organic solvent and separating the resulting

uranium loaded primary extractant from the lean

acid; (b) contacting said uranium loaded primary extractant

with a phosphoric acid strip solution containing

dissolved Fe+ + and separating the resulting U+4

loaded acid strip solution from said organic extractant;

an oxidizing agent to convert the uranium values to the

U+6 form; (d) contacting the resulting U+610aded acid

strip solution with a second portion of said organic

extractant and separating the resulting uranium loaded

secondary extractant from the lean acid strip solution;

(e) contacting said U+6 loaded secondary extractant

with a dilute aqueous ammonium carbonate solution

and separating the resulting aqueous ammonium uranyl

tricarbonate solution from said organic extractant; (0

separating any iron or other impurity-containing precipitates

that may form in step (e) from said aqueous

ammonium uranyl tricarbonate solution; (g) contacting

said aqueous ammonium uranyl tricarbonate solution

with an acid to form an aqueous acidic solution having

a pH of about 2 and removing from said aqueous acidic

solution the C02 formed as the carbonate ions are destroyed;

(h) contacting either the aqueous ammonium

uranyl tricarbonate solution or the aqueous acidic solution

of step (g) with a water-immiscible, organic solvent

for the acidified form di(2-ethylhexyl) phosphoric acid

and separating the resulting organic solution from said

aqueous acidic solution; and (i) contacting said aqueous

acidic solution with hydrogen peroxide at a pH in the

range of from about 3.5 to about 4.5 to precipitate a

uranyl peroxide compound.

The present invention also provides a process for

producing uranium peroxide from an aqueous ammonium

uranyl tricarbonate solution obtained by stripping

uranium values from an organic extractant containing

4,302,427

mates. These humates are by-products of decayed vegetable,

animal, or other types of organic matter contained

in the phosphate rock as mined which transfer to the

acid during the digestion step. The feed acid generally

also contains various metallic impurities, traces of silica

and gypsum solids that have crystallized after filtration.

The presence of these impurities may result in significant

difficulties during the uranium solvent extraction

process, primarily due to the buildup of solids at the

interface during phase separation in the extraction process.

While the present invention contemplates the recovery

of uranium from so-called brown phosphoric

acid, it is preferred to subject the brown feed acid to a

purification step shown generally at 2 which removes a

15 major portion ofpotentially interfering contaminants.

While any of the known phosphoric acid purification

processes may be employed for the preliminary cleanup

step, the preferred method is that described in commonly

assigned U.S. application Ser. No. 22,083 entitled

"Purification of Phosphoric Acid" filed on even date

herewith in the names of Allen and Berry, which is

hereby incorporated by reference. The purification

method described in that application comprises, in a

preferred form, cooling the feed acid, mixing an activated

clay such as bentonite with the cooled acid, adding

a flocculating agent to the acid-clay mixture to

cause impurity sedimentation and thereby producing a

partially clarified acid, and feeding the partially clarified

acid to an activated carbon column for removal of

the remaining impurities. This process produces a lowhumic,

low solids phosphoric acid (clean acid). When

the purification process is operated in the aforementioned

manner, the spent carbon material can be readily

and efficiently regenerated by the process described in

commonly assigned U.S. application Ser. No. 22,082

entitled "Regeneration of Activated Carbon" filed on

even date herewith in the names of Allen, Berry and

Leibfried, .which is hereby incorporated by reference.

The clean acid 3 from the purification step 2 is then

treated with an oxidizing agent 4 in oxidizer 5 to convert

any U+4 present to the U+6 form. Various oxidizing

agents such as sodium chlorate, air and the like, may

be employed although the preferred oxidizing agent is

hydrogen peroxide. The oxidizing agent should be

added to the clean acid in an amount sufficient which

insures that any uranium present as U+4 is converted

U+6. Typically, oxidizing agent additions offrom about

0.01% to about 0.1% (100% H202) and preferably from

about 0.01% to about 0.03% by weight are effective to

achieve this result. The progress of uranium oxidation

in the clean acid can be controlled by monitoring the

conversion of Fe+ + present in the solution to Fe+ + +.

Typically, wet-process acid may contain up to about 10

to 12 grams per liter of total iron (as Fe). The oxidizing

55 agent (e.g., a 50% aqueous solution of hydrogen peroxide)

can be added until the Fe+ + concentration is reduced

to a value of less than about 0.1 grams per liter

which insures that any U+4 originally present is converted

to U+6. The oxidation step is preferably accomplished

in a mixing tank and suitable mixing times can be

up to about 5 minutes or more preferably from about 5

to about 15 minutes, although longer periods of time are

acceptable.

If the acid has not been cooled as a part of the preferred

purification step described above, it should be

cooled at some point prior to the primary solvent extraction.

Commercially produced acid comes from the

gypsum filter at a temperature of about 140°-150° F. In

di(2-ethylhexyl) phosphoric acid with an aqueous ammonium

carbonate solution, said process comprising: (a)

contacting said aqueous ammonium uranyl tricarbonate

solution with an acid to form an aqueous acidic solution

having a pH of about 2 and removing from said aqueous 5

acidic solution the C02 formed as the carbonate ions are

destroyed; (b) contacting either the aqueous ammonium

uranyl tricarbonate solution or the aqueous acidic solution

of step (a) with a water-immiscible, organic solvent

for the acidified form di(2-ethylhexyl) phosphoric acid 10

and separating the resulting organic solution from said

aqueous acidic solution; and (c) contacting said aqueous

acidic solution with hydrogen peroxide at a pH in the

range of about 3.5 to 4.5 to precipitate uranium peroxide.

The present invention further provides a process for

removing alkaline-soluble ammonium complexes of

di(2-ethyhexyl) phosphoric acid from aqueous alkaline

uranium-containing solutions which are to be acidified

in subsequent processing to recover uranium values, 20

said process comprising: (a) adding a water-immiscible

organic solvent for the acidified form of di(2-ethylhexyl)

phosphoric acid to said alkaline solution just

prior to or during acidification; and (b) separating the

resulting di(2-ethylhexyl) phosphoric acid containing 25

organic solution from said aqueous solution.

Precipitation of uranium from an acid medium rather

than from the alkaline carbonate medium of the prior

art allows for significant control of iron precipitation

which in tum permits wider latitude in extractant con- 30

centrations in the second cycle of the reductive stripping

process. The ORNL process presently employs a

0.5 M DEPA concentration in the first cycle and a 0.3

M DEPA concentration in the second cycle to suppress

coextraction of iron. In the process of the present inven- 35

tion, however, 0.5 M DEPA can be employed for both

cycles resulting in increased extraction and handling

efficiency. The uranium peroxide precipitate has the

additional advantage of being convertible to a commercial

uranium concentrate by simple drying, thus obviat- 40

ing the need for expensive calcining. Moreover, the

uranium peroxide precipitate may be shipped directly to

uranium converters as a peroxide slurry.

While processes for treating various uranium-containing

solutions to produce uranium peroxide precipi- 45

tates are known in the prior art, see, for example, Shabbir

et aI, "Hydrogen Peroxide Precipitation of Uranium,"

Bureau of Mines Report of Investigations RI7931

(1974); Caropreso, U.S. Pat. No. 4,024,215, this

process has not been used to produce uranium peroxide 50

from alkaline strip solutions generated in the DEPATOPO

extraction of uranium from wet-process phosphoric

acid.

FIG. 1 is a block flow diagram showing the preferred

form of the process of the present invention.

The process of the present invention will now be

described with reference to FIG. 1. Feed to the process

of the present invention via line 1 comprises wet-process

phosphoric acid obtained by the sulfuric acid digestion

of phosphate rock such as that found in Florida. 60

This commercial wet-process phosphoric acid from the

gypsum filter usually runs approximately 28-30% by

weight P20S.

The first step of the process of the present invention

may comprise phosphoric acid purification. Commer- 65

cially produced wet-process phosphoric acid made

from Florida phosphate rock is generally brown in

color as a result of contamination with so-called hu4,302,427

30%, and a ferrous ion concentration of at about 25 to

about 45 grams and preferably about 35 to about 40

grams per liter as Fe+2. Loaded strip solution 12 (Le.,

containing uranium) should contain an excess ferrous

concentration of about 14 to about 17 grams per liter,

more preferably about 15 to about 17 grams per liter.

The strip acid can be conveniently prepared by adding

ground metallic iron at a suitable rate, e.g., about 0.2 to

about 0.3 pounds per gallon to the acid in a preparatory

step. (The specific amount added depends, to some

extent, on the amount of iron present in the original

phosphoric acid). A convenient source of phosphoric

acid for use in the primary cycle stripping is formed by

taking a slip stream of clean acid from the initial purification

step described above, and either concentrating it

or adding a small amount of more concentrated acid, if

required, in order to produce the desired P20S concentration

in the strip acid. The stripping step is preferably

carried out in the organic continuous mode at an organic

to acid ratio of about 0.7 to about 1. The barren

organic 13 is recycled back to the primary extraction

unit.

The loaded strip solution 12 is then subjected to an

oxidation step 14 with an oxidizing agent of the type

described above, preferably with hydrogen peroxide. In

this step the acid soluble V+4values are converted back

to the U+6 form.

The oxidized strip acid 15 is then contacted again

with a DEPA-TOPO extractant 16 in a second countercurrent

extraction system in a manner similar to that

used in the primary system. The major portion of the

lean strip acid 18 is returned to the primary extraction

section where it is mixed with the main acid feed

stream, although some of the lean strip acid may be

35 recycled.

As described above, the DEPA-TOPO concentrations

in the primary extraction cycle conventionally are

about 0.5 molar DEPA and 0.125 molar TOPO. According

to the preferred ORNL process the secondary

extraction cycle is carried out at reduced extractant

concentrations of 0.3 M DEPA and 0.075 M TOPO to

suppress the coextraction of soluble iron values from

the strip acid. As will be described in furtller detail

below, the process of the present invention avoids the

problem of iron contamination of the uranium product

by precipitating uranium in an acidic medium. Accordingly,

it is possible and advantageous to perform the

secondary extraction cycle of the process of the present

invention at DEPA concentrations of 0.5 M (Le., at the

same level employed for the primary cycle) although

DEPA concentration within the range of from about

0.5 to about 0.3 are preferred. The ability to upgrade the

secondary extraction concentrations provides a distinct

advantage in the overall process in that only one source

of DEPA-TOPO mixing need be employed and handling

problems are therefore significantly reduced. In

addition, the use of the same DEPA-TOPO concentrations

in both the primary and secondary cycle provide

for the added flexibility of bleeding primary and secondary

loaded organics back and forth between the two

cycles to achieve better control of the process. For

example, if a buildup of iron contamination occurs in

the primary cycle, the extractant from this loop can be

bled through the secondary cycle with effects precipitation

and removal of iron from the system in the laterdescribed

dilute carbonate stripping step.

The pregnant organic 18 from the secondary extraction

unit is preferably scrubbed with water in scrubber

order to maxImIze the coefficients of extraction for

uranium in the subsequent operations, the acid should

be cooled to a temperature in the range of from about

100· F. to 130· F. Cooling much below 100· F. requires

considerable additional equipment thus resulting in ad- 5

ditional capital cost. Preferred is a temperature of about

120· F. This cooling step can precede or follow the

above-described oxidation step.

In the next step ofthe process of the present invention

the cooled, clarified, oxidized acid 6 is fed to a primary 10

solvent extraction unit 7 in which it is contacted countercurrently

with an immiscible organic extractant 8 to

cause transfer of the uranium values into the organic

phase. In practice, the extraction is carried out in a

number of sequential extraction stages, each comprising 15

a mixer-settler arrangement. In the preferred embodiment

all the extraction stages have all aqueous-continuous

phase except the last stage which has an organic

continuous phase to minimize acid entrainment. The

lean aqueous acid streams may be returned to the phos- 20

phoric acid plant after suitable treatment to. remove

entrained organics.

The extractant employed in the process of the present

invention is a mixture of di(2-ethylhexyl) phosphoric

acid (DEPA) and trioctylphosphine oxide (TOPO) 25

dissolved in an organic diluent such as kerosene.

Contact of the uranium-bearing acid solution with this

immiscible extractant mixture results in the conversion

of uranyl ions to a V02++-DEPA complex. Typically,

the extractant contains about 0.1 to 1 mol per liter of 30

DEPA and about 0.025 to about 0.25 mol per liter of

TOPO. The feed acid to extractant ratio by volume is

generally in the range of about 0.1 to 10. Contact times

generally range from about 1 to 5 minutes, preferably

from about 2 to about 3 minutes.

Even when preliminary acid purification is employed,

solid impurities (crud) may build up at the phase

interface in the settlers (especially in the first extractiori

stage). While some buildup of this interfacial crud is

tolerable, it is preferable to effect continuous removal of 40

this crud layer by the process of commonly assigned

V.S. patent application Ser. No. 22,218 entitled "Improved

Method For Solvent Extraction of Metallic

Mineral Values From Acidic Solutions" filed on even

date herewith in the names of Allen and Berry, which is 45

hereby incorporated by reference. In this application a

process is described wherein a dispersion of air bubbles

is introduced into the mixer which causes the crud to

float to the surface in the settler where it is continuously

removed for example, by skimming. 50

The next step of the process of the present invention

isreductive stripping. In this step the pregnant organic

9, which now contains from about 0.2 to about 0.6, and

more generally from about 0.3 to about 0.5 grams per

liter of uranium values in the hexavalent state, is con- 55

tacted with a Fe+2-containing phosphoric acid stripping

solution 10 in a stripping vessel 11 to cause transfer

of the uranium values into the aqueous phase. During

the reductive stripping ferrous ion is oxidized to ferric

ion and the DEPA-complexed uranyl ion is reduced to 60

the V +4 ion. Since the DEPA has very little affinity for

the quadrivalent uranium species, the V+4 ion concentrates

in the aqueous stream. Typical loadings in the

stripping acid are from about 10 to about 12 grams per

liter (as V). 65

Stripping solution 10 (substantially uranium free)

should have a P20S concentration of about 28% to 32%

P20S by weight, and preferably about 29% to about

4,302,427

19 to remove any entrained phosphoric acid which

could increase ammonia consumption in later processing.

The scrubbed secondary organic 20 is then subjected

to a carbonate stripping step in vessel 21. In this

stripping step, the uranium values (U+6) are stripped 5

from the pregnant secondary cycle organic with a dilute

ammonium carbonate solution 22 which results in the

formation of a soluble ammonium uranyl tricarbonate

(AUT) complex. Preferably the ammonium carbonate is

produced in a separate system and fed to the alkaline 10

stripping vessel as an aqueous solution. In general, an

ammonium carbonate equivalent concentration of from

about 0.25 M to about 1.0 M may be employed as lopg

as conditions are controlled to avoid precipitation of the

AUT in the stripping system. Preferred are ammonium 15

carbonate concentrations under 0.5 M with the most

preferred concentrations falling in the range of about

0.3 M to 0.4 M. As the concentration of the ammonium

carbonate increases, the loading of uranium, of course,

drops off. The stripped organic 23 is then contacted 20

with acid 24 (e.g., H2S04) in a mixer 25 to reconvert it

from the NH4-form to the acid-form which can be recycled

to the secondary extraction cycle via line 26.

It is important to closely control the pH and temperature

conditions during the above-described carbonate 25

stripping step. Applicants have found that a pH in the

narrow range of about 8.2 to about 8.5 is desirable. If the

stripping is carried out at a pH much above 8.5 to 9,

increased amounts of iron are extracted. If, on the other

hand, the ammonium carbonate concentration drops 30

much below about 8.2 uranium extraction falls off significantly.

The alkaline stripping step is also preferably

carried out at a temperature in the range of about 1000

to 1250 F. Most preferred, is a temperature of about

1150 F. Warm stripping serves to suppress the formation 35

of emulsions which tend to entrain carbonate materials

in the organic phase. When the carbonate-containing

organic is converted back to the acid form in the acid

scrub vessel, the uranium values in the carbonate solution

are reextracted by the organic causing the overall 40

efficiency of the system to drop.

The dilute carbonate aqueous strip solution 27 is then

acidified with an acid 28 to a pH of about 2 in a stirred

tank reactor 29 equipped with an air sparger to assist in

removal of liberated C02. Any mineral acid which will 45

destroy the carbonate ions without introducing an insoluble

anion into the system can be employed. Suitable

acids include sulfuric acid, nitric acid, hydrochloric

acid and phosphoric acid. Sulfuric acid is preferred.

Destruction of the carbonate ion and reduction ofpH to 50

about 2 is an essential prerequisite to precipitation via

the peroxide route. Because the ammonium-form of

DEPA is slightly soluble in alkaline solution, the aqueous

strip solution from the dilute carbonate stripping

stage contains quantities of ammonium-DEPA. Acidu- 55

lation of such a stripping solution, however, causes the

precipitation of highly insoluble waxy derivatives of

DEPA. The precipitation of the waxy acid form of

DEPA presents a serious obstacle to the utilization of

acid precipitation of uranium via the peroxide route. 60

Applicants have unexpectedly discovered that the advantages

inherent in the DEPA-TOPO extraction process

can be coupled with the advantages inherent in

acid precipitation by the peroxide route by treating the

strip solution either before or during acidulation with an 65

effective amount of an organic solvent 30 for DEPA

waxes such as kerosene. Other known organic hydrocarbon

solvents such as Amsco 480, a highly refined

petroleum based solvent, may also be employed in this

step. In practice, amounts up to about 5 to 10% by

volume of the DEPA solvent based on the acid solution

are mixed with the acidified solution and sent to a separating

stage where settling and air flotation coupled

with skimming are effective to remove the organic

phase which contains the dissolved DEPA waxy materials.

Hydrogen peroxide 33 is then added to the clarified

acidic solution 31 which contains U02++. The pH is

then adjusted with ammonia, via 32, to approximately

3.5 to 4.5, in the precipitator 34, to produce a uranyl

peroxide product 35. In association with the precipitation

step, a reaction mixture is fed to a settler from

which the uranium-containing sludge is withdrawn and

washed to remove soluble ammonium salts (e.g.,

(NH4)2S04) contamination. The sludge is then centrifuged

to remove water, and dried, for example, at about

1100 C. to produce a uranium concentrate suitable for

direct utilization by uranium converters. Unlike the

alkaline precipitated AUT processes of the prior art,

calcination to remove C02 from the product of the

present invention is not required. Moreover, the peroxide

precipitate may be shipped to a converter in the

form of a slurry which further eliminates processing

steps.

The following example is intended to illustrate more

fully the nature of the present invention without acting

as a limitation on its scope.

EXAMPLE

Brown phosphoric acid from a conventional wet

process phosphoric acid plant containing approximately

27.9% P20S and 0.129 grams per liter ofU, at a temperature

of about 1400 F., was introduced to a purification

unit at the rate of about 10 gallons per minute. The acid

was cooled in a heat exchanger to 1190 F. A bentonite

clay was added to the cooled acid in stirred-tank mixer

at the rate of about 0.3% by weight of the acid. Flocculant,

specifically Nalco 7873, was added at the rate of 15

ppm by weight in a flocculation tank. This material was

overflowed from the flocculation tank to a clarifier

where the solids were permitted to settle. In this clarification

step a major portion of the suspended solids and

acid color were removed. In this example the solids in

the brown phosphoric acid were 3.12% by volume and

the acid was a dark brown color. The partially clarified

acid contained 0.14% solids by volume and 58% of the

color bodies had been removed, as measured on a spectrophotometer.

This partially clarified acid was then fed to the inlet

of a carbon column system at the rate of 8 gallons per

minute. The carbon column system was operated in a

series upflow expanded bed manner, utilizing five columns

approximately 2.5 feet in diameter with a settled

carbon bed depth of about 6 feet. The acid leaving the

column (clean acid) was light green in color and overall

color body removal was approximately 92% as measured

on a spectrophotometer.

Hydrogen peroxide was then added to this clean acid

and the Fe+ + was lowered from the original 1.1

gramslliter to 0.07 gramlliter (as Fe+ +2). The H202

(35% concentration by weight) was added at the rate of

2.2 pounds (of 100% peroxide) per ton of 100% P20S, as

a liquid in a stirred tank. The oxidized material then

overflowed to a hold tank for completion of the oxidation

and from the hold tank to a surge tank.

4,302,427

Oxidized clean acid was then pumped to the solvent

extraction system at the rate of 8.25 gallons per minute.

This acid contained uranium in the U+6 state in the

amount of 0.129 grams per liter (measured as U). This

acid was then contacted with 0.5 M DEPA-0.125 M 5

TOPO in a 4 stage countercurrent solvent extraction

mixer-settler system. The organic flow rate was about

4.1 gallons per minute. All stages, with the exception of

the last mixer-settler unit, were run aqueous continuous.

The last stage was run organic continuous to minimize 10

organic entrainment, and in the settler portion of the

last stage the U+610aded primary extractant was separated

from the lean acid. The lean acid was further

treated to remove entrained extractant to minimize

organic losses and then returned to the conventional 15

wet process phosphoric acid plant. The uranium concentration

in the lean acid (raffinate) was 0.0038

grams/liter (as U). This represents an extraction efficiency

of 97%.

The U+6 loaded primary extractant was pumped to a 20

3 stage countercurrent stripping system where it was

contacted with a phosphoric acid strip solution containing

46 gramslliter of dissolved iron as Fe+ +z and 29%

PzOs. The strip acid flow rate was 558 ml per minute.

The loaded primary extractant flow was 4.8 gallons per 25

minute.

The U+6 loaded acid strip solution containing 8

gramslliter of uranium (measured as U) was then stored

in a surge tank. The primary extractant was reduced

from 0.285 gramslliter ofU to 0.0018 gramslliter result- 30

ing in a stripping efficiency of,99.4%. The extract was

then recycled to the primary extraction mixer-settler.

About 2.3 days accumulation of the U+4 loaded acid

strip solution was collected in the surge tank.

Fifty gallons of the U+4 loaded acid strip solution 35

was contacted on a batch~basis with 35% hydrogen

peroxide solution (1500 ml of the 35% solution). This

converts the Fe+z to Fe+3 and the uranium values to

the U+6 form. The contact was permitted for about 3

hours. The resulting U+6 loaded acid strip solution 40

containing 8.4 grams per liter of uranium values (calculated

as U) was fed to a 4 stage mixer-settler system at

the rate of 25 mllminute. It was contacted with 0.4 M

DEPA, 0.1 M TOPO which was fed to the system at the

rate·of 37.5 mllminute. In the last mixer settler a U+6 45

loaded secondary extractant was separated from the

lean acid strip solution. The lean strip acid was returned

to the primary extraction circuit and it contained 0.082

grams per liter of uranium values (as U). The extraction

efficiency was 99%. 50

The U+610aded secondary extractant (37.5 mllmin.)

was· contacted with a spent water solution to remove

entrained phosphoric acid. The water washed extractant

was then contacted with 0.3 M ammonium carbonate

solution (15 mllmin). The resulting aqueous ammo- 55

nium uranyl tricarbonate solution was then separated

from the organic extractant. .The uranyl tricarbonate

solution contained 21.0 gramslliter of uranium values

(measured as U). The uranium in the organic extractant

was reduced from 5.75 gramslliterof uranium values 60

(measured as U) to 0.107 gramslliter, resulting in a

stripping efficiency of approximately 98%. The organic

extractant was then mixed with 5 mllminute of 20%

HZS04 for the purpose of regenerating the organic extractant.

The extractant was then recycled to· the sec- 65

ondary extraction circuit. The uranyl tricarbonate solution

was thenstored for use in the uranium precipitation

circuit. During this test the pH in the stripping system

ranged from 8.7 to 9.1. This resulted in the undesirable

precipitation ofiron in the mixer-settlers. In subsequent

runs, it was determined that operation of the system

within the pH range of about 8.2 to about 8.5 would

limit the amount of this undesirable precipitate yet

maintain acceptable uranium stripping efficiencies. During

the secondary extraction the contact with the aqueous

ammonium carbonate was at a temperature of from

100° F. to 120° F.

The separated aqueous ammonium uranyl tricarbonate

solution was contacted with 5% sulfuric acid to

reduce the pH to 2.0. During this acidification air was

sparged into the solution to aid in eoz removal, which

eoz is formed as the carbonate ions are destroyed. A

small amount of kerosene (or a suitable organic solvent

for acidified DEPA) was added either prior to or during

acidification for the purpose of dissolving DEPA

waxes which formed during the acidification stage. The

kerosene was used in amount of 5-10% by volume of

the amount of solution being treated. If the DEPA

waxes were not dissolved they would build up on equipment

surfaces and could result in emulsion problems.

The kerosene containing DEPA waxes was then

removed from the aqueous acidic solution. The aqueous

acid solution was then mixed with 30% HzOz at the rate

of 0.24 g (100% HzOz) per gram of uranium values

(measured as U). The pH was then adjusted with ammonia

to 3.5-4.0. The amount of ammonia used was approximately

0.11 gram NH3 per gram of U. Retention

time for this precipitation reaction was approximately

70 minutes although shorter times in other runs were

acceptable.

The precipitation uranium peroxide prepared was

then separated from the liquid, washed, then dried and

analyzed. The uranium precipitation efficiency was

99.9%. The precipitate was dried at 100° C., 200° e. and

550° C., then analyzed. The analyses (based on U) were

68.2%, 79.8% and 83.3% respectively.

While certain specific embodiments of the invention

have been described with paticularity herein, it will be

recognized that various modifications therefore will

occur to those skilled in the art. Therefore, the scope of

the invention is to be limited solely by the scope of the

appended claims.

We claim:

1. A process for recovering uranium from wet-process

. phosphoric acid containing hexavalent uranium

values comprising:

(a) contacting said acid with an organic extractant

comprising a mixture of di(2-ethylhexyl) phosphoric

acid and trioctylphosphine oxide in a phosphoric

acidimmiscible organic solvent and separating

the resulting uranium loaded primary extractant

from the lean acid;

(b) contacting said uranium loaded primary extractant

with a phosphoric acid strip solution containing

dissolved Fe++ and separating the resulting U+4

loaded acid strip solution from said organic extractant;

with an oxidizing agent to convert the uranium

values to the U+6 form;

(d) contacting the resulting U+6 loaded acid strip

solution with a second portion of said organic extractant

and separating the resulting U+6 loaded

secondary extractant from the lean acid strip solution;

4,302,427

9. The process of claim 1 wherein said loaded secondary

extractant is contacted with said dilute aqueous

ammonium carbonate solution at a pH in the range of

about 8.2 to about 8.5.

10. The process of claim 1 wherein said loaded secondary

extractant is contacted with said dilute aqueous

ammonium carbonate solution at a temperature of about

1000 F. to about 1250 F.

11. A process for producing uranium peroxide from

an aqueous ammonium uranyl tricarbonate solution

obtained by stripping uranium values from an organic

extractant containing di(2-ethylhexyl) phosphoric acid

with an aqueous ammonium carbonate solution, said

process comprising:

(a) contacting said aqueous ammonium uranyl tricarbonate

solution with a water-immiscible, organic

solvent for the acidified form di(2-ethylhexyl)

phosphoric acid either prior to or during the acidification

of step (b);

(b) contacting said aqueous ammonium uranyl tricarbonate

solution with an acid in the presence of said

organic solvent to form an aqueous acidic solution

having a pH of about 2 and removing from said

aqueous acidic solution the C02 formed as the

carbonate ions are destroyed;

dissolved DEPA from said aqueous acidic solution;

and

(d) contacting said aqueous acidic solution with hydrogen

peroxide at a pH in the range of about 3.5 to

4.5 to precipitate uranium peroxide.

12. The process of claim 1 or 11 additionally comprising

the step of drying said uranium peroxide precipitate.

13. A process for removing alkaline-soluble ammonium

complexes of di(2-ethylhexyl) phosphoric acid

from aqueous alkaline uranium-containing solutions

which are to be acidified in subsequent processing to

recover uranium values, said process comprising:

(a) adding a water-immiscible organic solvent for the

acidified form of di(2-ethylhexyl) phosphoric acid

to said alkaline solution just prior to or during

acidification; and

(b) separating the resulting di(2-ethylhexyl) phosphoric

acid containing organic solution from said

aqueous solution.

14. The process of claim 1, 11 or 13 wherein said

water-immiscible organic solvent for the acidified form

of di(2-ethylhexyl) phosphoric acid is kerosene.

* * * * *

(e) contacting said U+6 10aded secondary extractant

with a dilute aqueous ammonium carbonate solution

and separating the resulting aqueous ammonium

uranyl tricarbonate solution from said organic

extractant;

(0 separating any iron or other impurity-containing

precipitates that may form in step (e) from said

aqueous ammonium uranyl tricarbonate solution;

(g) contacting the aqueous ammonium uranyl tricarbonate

solution with a water-immiscible organic 10

solvent for the acidified form di(2-ethylhexyl)

phosphoric acid either prior to or during the acidification

of step (h);

(h) contacting said aqueous ammonium uranyl tricarbonate

solution with an acid in the presence of said 15

organic solvent to form an aqueous acidic solution

having a pH of about 2 and removing from said

aqueous acidic solution the C02 formed as the

carbonate ions are destroyed;

(i) separating the resulting organic solution contain- 20

ing dissolved DEPA from said aqueous acidic solution;

and

U) contacting said aqueous acidic solution with hydrogen

peroxide at a pH in the range of about 3.5 to

4.5 to precipitate uranyl peroxide compound.

2. The process of claim 1 wherein said wet process

phosphoric acid containing hexavalent uranium values

is produced by contacting wet process phosphoric acid

with an oxidizing agent for converting uranium values

to the hexavalent state.

3. The process of claim 1 additionally comprising the

step of purifying said wet process phosphoric acid prior

to step (a) to remove potentially interfering contaminants.

4. The process of claim 1 wherein said phosphoric 35

acid strip solution contains from about 25 to about 45

grams of dissolved Fe+ + per liter.

5. The process of claim 1 wherein said U+4 loaded

acid strip solution contains about 14 to about 17 grams

of Fe+ + per liter. 40

6. The process of claim 1 wherein said phosphoric

acid strip solution contains from about 28% to about

32% by weight of P20S.

7. The process of claim 1 wherein said dilute aqueous

ammonium carbonate solution is from about 0.25 M to 45

about 1.0 M in ammonium carbonate.

8. The process of claim 7 wherein said dilute aqueous

ammonium carbonate solution is produced in a separate

operation and fed to the system as an aqueous solution.

UNITED STATES PATENT AND TRADEMARK OFFICE

CERTIFICATE OF CORRECTION

PATENT NO. 4,302,427

DATED November 24, 1981

INVENTOR(S) William W. Berry et al

It is certified that error appears in the above-identified patent and that said Letters Patent

are hereby corrected as shown below:

Column 1, line 41, "ORNL-2925" should read -- ORNL-2952

Column 2, line 6, "orginal" should read -- original --

Column 6, line 64, "with" should read -- which

~igncd and ~calcd {his

Fint Da'I 0 f JIIII. I'll

ISEALI

GERALD J. MOSSINGHOFF

Commissioner of Patents and TfDdema,ks