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
... /
32
31\ NH3 29-
ACI 01 FICATION PRECIPITATION ... H2 0 2
, \33
WAX/SOLVENT
r--35
"
o
S
URANYL PEROXIDE
4,302,427
1
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
2
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;
(c) contacting said U+4 loaded acid strip solution with
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
4
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
3
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
6
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
5
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
7
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
8
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
9
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
10
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;
(c) contacting said U+4 loaded acid strip solution
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
5
30
25
12
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;
(c) separating the resulting organic solution containing
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.
* * * * *
11
(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.
50
55
60
65
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