Note : Les descriptions sont présentées dans la langue officielle dans laquelle elles ont été soumises.
FIELD OF THE INVENTION
This invention is directed to a novel hydro-
metallurgical process for the recovery of copper and
optionally, iron, from sulfide copper concentrates
containing common copper sulfides such as chalcopyrite,
bornite and chalcocite.
BACKGROUND OF THE INVENTION
Present copper recovery processes utilize
copper smelters which inherently create a significant
amount of air pollution and hence ecological harm to the
environment, In recent years, there has been strong
emphasis to minimize ecological damaye to the environ-
ment and as a consequence certain processes have been
developed for the treatment of copper concentrates to
recover copper without attendant environmental problems.
A hydrometallurgical process for the electro-
lytic recovery of selected base metals (especially
copper, and optionally nickel) from sulfide ore concen-
trates concurrently with the extraction of metallic iron
in commercially usable quantities had previously been
developed by the applicant and is now the subject of
United States Patent No. 4,159,232, issued June 26,
1979. The process disclosed and claimed in that patent
utilizes at least a primary and a secondary bank of
sequentially disposed electrolytic cells, the cells of
each bank being electrically connected in parallel.
Each of the cells has separate anode and cathode com-
partments, the compartments being separated from other
compartments in a bank by a permeable dividing element
capable of passing electrolyte between the compartments.
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A first supply of anode solution is continuously ~ith-
drawn from the anode compartments of the electrolytic
cells in each bank of cells, the anode solution being an
aqueous electrolyte includiny in solution hydrochloric
acid and a soluble metal chloride. The anode solution
is transported to at least one leaching vessel to con-
tinuously leach a supply of ore concentrate to one
leaching vessel to reduce ferric ions in solution to
their lowest valence state (ferrous). The leaching
vessel generates a liquid-solid slurry output including
solid residue, partially leached concentrate and leach-
ing solution. The solids are separated from the liquid
in the resulting slurry output by using a suitable
solids-liquid separator. The solution from the solids-
liquid separator is returned to the cathode compartmentsof both the primary and secondary banks of cells, with a
first preselected portion of the liquid being returned
to the cathode compartments of the secondary bank of
cells. Base metal is precipitated at the cathode in a
non-adherent form to provide a slurry with a cathode
solution. Base metal deposits are continuously with-
drawn from the slurry comprising cathode solution and
precipitated base metal obtained from the cathode com-
partments of the bank of cells. The separated cathode
solution obtained from the base metal slurry is returned
to the cathode compartments of both the primary and
secondary bank of cells with a second pre-selected
portion of thc separated cathode solution being returned
to the cathode compartments of the secondary bank of
cells. The amounts of the first and second pre-selected
portions are small enough to allow for the establishment
in the cathode compartments of base metal impoverished
solution areas. The amounts of remaining separated
leaching liquid and separated cathode solution returned
to the cathode compartments of the primary bank of cells
are large enough to avoid the development of base metal
impoverished areas next to the cathode electrodes of the
primary bank of cells. sase metal impoverished cathode
solution is continuously withdrawn from the base metal
impoverished solution areas of the cathode compartments
of the secondary bank of cells. Hydrogen gas is cathod-
ically produced from the catholyte of the electrolytic
cells in each bank of cells. The base metal impover-
ished cathode solution is evaporated and crystallized to
yield hydrated ferrous chloride. The hydrated ferrous
chloride is reduced by the hydrogen gas at a selected
elevated temperature to produce metallic iron. Copper
is obtained from the cathode compartments of both the
primary and secondary banks of cells.
SUMMARY OF THE INVENTION
The invention is directed to a process for
recovery of copper and optionally iron. The process is
not intended for the recovery of nickel or nickel
minerals.
A significant advantage of the process is that
no solids are tolerated in the electrowinning cells.
Further, no solution flows from the cathode compartments
to the anode compartments of the electrowinning cells.
The process is also somewhat simplified compared to pre-
vious processes in that alkali metal chlorides such as
4B9
sodium chloride have been eliminated fro~ the process
solutions. The hydrochloric acid concentration has been
increased.
The invention is directed to a hydrometal-
lurgical process for the recovery of copper fromsulphide copper concentrate containing copper and iron
involving the use of electrowinning cells having anolyte
and catholyte chambers, wherein the efficiency of the
electrowinning cells is increased by preventing any
mixing of the anolyte and the catholyte in the electro-
~inning cells. The mixing of the anolyte and the
catholyte is prevented by a membrane which is substan-
tially impervious to copper chloride and iron chloride
complexes but is pervious to hydrogen ions. The anolyte
chambers of the cells are operated at hydrogen chloride
levels above about 1 mole H+ to thereby maintain the
copper and iron in solution as copper chloride and iron
chloride complexes so that they do not pass through the
membrane. The anolyte and catholyte are substantially
2~ free from alkali metal chlorides such as sodium
chloride. No significant amount of solids is present in
the electrowinning cells.
The process for the recovery of copper
utilizes electrowinning cells and comprises leaching
copper concentrate in first and second stages, the
leaching in the first stage being done with a partially
reduced leach liquor obtained rom the second stage;
separating the concentrate obtained from the first stage
in a solid liquid separator, the partially leached
solids being directed to the second leaching stage, the
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124~4B~
liquid from the separator being directed to the catho-
lyte compartment of an electrowinning cell; leaching the
solids in the second leaching stage with regenerated
leaching liquid obtained from the anolyte compartment of
the electrowinning cell, the anolyte and catholyte
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compartments being separated by a membrane which is
impermeable to copper chloride and iron chloride
complexes but is permeable to hydrogen ions; separating
the concentrate obtained from the second stage in a
second solid-liquid separator and filtering and washing
the solids residue from the separator.
DRAWINGS
In the drawings;
FIGURE l represents a schematic flow sheet
illustrating the copper and iron recovery process.
FIGURE 2 represents a graphical depiction of a
batch test demonstrating copper concentration relative
to time utilizing the process.
FIGURE 3 represents a flow arrangement wherein
hydrogen required for iron reduction is generated elec-
trolytically.
FIG~RE 4 represents an alternative arrangement
wherein hydrogen is produced as a separate operation.
FIGURE 5 represents the results of electro-
oxidation tests carried out at various cell voltages
utilizing the process.
FIGURE 6 represents the results obtained by
removing water from a stripped catholyte solution.
DETAILED DESCRIPTION OF A PREFERRED
EMBODIMENT OF THE INVENTION
Referring to FIGURE l, the preferred arrange-
ment for leaching of the concentrate involves a two
stage, countercurrent flow of solids and liquids.
Finely ground copper concentrate is introduced
into the first stage leach 2 wherein it is partially
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leached with partially reduced leach liquor obtained
from the second stage leach 6. Leaching is carried out
at temperatures greater than 80C but below the boiling
temperature of the liquid (approximately 106C).
The first stage leaeh 2 is followed by a
solid-liquid separation device 4 from which the under-
flow 7 is eomprised of partially leached concentrate and
the overflow 11 is comprised of a clear solution enrich-
ed in copper. In order to achieve rapid settling rates
of the solids as well as an overElow solution which is
free of solids, it has been found to be necessary to use
organic settling aids (flocculants). It has been found
that Percol E10 at a rate of 1 mg/litre is a satisfac-
tory floeeulant on a eost effeetiveness basis.
The partially leached coneentrate of the
underflow 7 is further leached in a second leaching ves-
sel 6 with regenerated leaching li~uid obtained from the
anode compartments 16 of electrowinning cells.
The overflow 13 from the second stage leach 4
goes to a solid-liquid separation deviee 8 from whieh
the underflow 9 eonsists of a thickened slurry of eom-
pletely leaehed eoneentrate residue and the overflow is
the partially reduced leach liquor used for the first
stage leach 2.
The thickened leach residue of underflow 9 is
filtered in residue filter 34 and washed with a suitable
solution such as a barren solution containing 15% by
weight HCl to remove entrained leaeh liquors leaving a
final residue suitable for disposal or for reeovery of
sulfur or eontained traee eonstituents.
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The relative proportion oE copper leached in
each of the first and second stage leaches 2 and 4 can
be varied. The only significant parameters in this re-
gard are that the second stage leach residue must be
S completely leached and the first stage leach overflow
solution must be free of ferric and cupric ions.
A balanced circuit may be achieved by adding
sufficient new concentrate to the first leach 2 to match
the available ferric and cupric ions in the regenerated
solution going to the second stage leach from the turbo-
aerator 10. Alternatively to generating all the oxi-
dized species (ferric and cupric) in the turboaerator
10, air may be introduced into the second stage leach 6
so that leaching of the copper minerals and regeneration
of leach liquor are occurring at the same time.
The enriched solution from the Eirst stage
leach thickener 4 is directed to the cathode compart-
ments of electrolytic cells (electrowinning cells). The
anode 16 and cathode 12 compartments of the electrowin-
ning cells (which compartments are shown as separateunits in the flow sheet of Figure 1) are separated by a
solution impermeable membrane so that anolyte and
catholyte are prevented from mixing. Preventing any
mixing of solutions is essential for achieving maximum
electrowinning efficiency. A commercially available
membrane material known as NAFION'~ (a perfluorosulfonic
acid polymer which may be supported by a fluorocarbon
polymer fabric) has been found to be suitable for this
purpose. In the present system, passage of elements
through this membrane has been found to be limited to
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small mobile cations such as hydrogen ions or sodium
ions. It has been found that when a cell has a catho-
lyte fully loaded with copper and an anolyte free of
copper but fully converted to the ferric state, no
inefficiency is observed due to the transport of these
ions through the membrane.
United States Patent No. 4,159,232 indicated
the need for chloride salt such as sodium chloride in
order to maintain adequate conductivity in the solution.
With the present electrolytic cell employing NAFION~
membranes it has been found that sodium chloride, or a
like salt, is not necessary and indeed improved results
can be achieved by increasing the hydrochloric acid
concentration and operating with no sodium chloride
present. This is demonstrated by the results of the
following set of electrowinning tests carried out with
solution generated by leaching of chalcopyrite
concentrates.
TABLE 1
HCl Cu Cu Deposi-
Concentration Range Cell tion Rate Power
Molar qpl Voltaqe q/m2/min. kwh/k~ Cu
0.65 28-2 1.56.35 0.691
(+44 gpl Na)
2.5 46-27 1.5511.08 0.693
2.5 22-10 1.559.25 ~.750
2.5 10-3 1.555.16 1.146
2.5 44-9.5 2.012.27 1.245
1.75 28-7 1.526.99 0.950
1.75 31-6 2.0211.94 1.300
~LZ~69L~39
The influence of high hydrochloric acid con-
centration on the copper deposition rate is demonstrated
to be very beneficial. At the same time, the elimina-
tion of sodium chloride from the process streams greatly
simplifies the process since the tendency for metal
salts (eg. CuCl) to crystallize from solution during
leaching or subsequent transport of solutions is
decreased and one less salt has to be dealt with in the
subsequent rejection of iron from copper depleted
solutions.
It iS apparent from the tabulated results in
Table 1 that the power required to win copper from solu-
tion increases as the concentration of copper decreases.
It is therefore undesirable to remove all the copper
from solution since excessive power costs would be in-
curred. In addition, the results as shown in Figure 2
(batch test) and Table 1 reveal that the copper deposi-
tion rate decreases as the concentration decreases.
Operating at copper concentrations of less than 10 grams
per litre results in a disproportionate increase in the
required electrode area per unit of copper recovered at
these low copper concentrations. Counterbalancing this,
however, is the requirement that the portion of the
depleted catholyte which goes to the crystallizer 26 for
ferrous chloride tetrahydrate crystallization must be
completely free of copper. Any copper remaining in this
stream would represent a loss of copper and would
contaminate the iron product.
A balance must therefore be struck. There are
a number of possible cell configurations and operating
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conditions and costs that can be used to achieve the
above objectives. Two pos.sible alternative configura-
tions are shown in Figures 3 and ~. Figure 3 (Alterna-
tive A) shows a configuration ~herein all copper is
electrowon from solution and hydrogen required for iron
reduction is generated electrolytically. Two types of
cells are illustrated. One type can have cathode com-
partments with sloplng bottoms for copper removal. The
second type is used to produce hydrogen and ferric ion.
The hydrogen is drawn off the cathode compartments o
the second cell type.
Figure ~ (Alternative B) shows hydrogen being
produced in a separate operation. Ferric oxidation is
carried by means of turboaeration. Only one type of
cell is required since every cell is producing copper.
This means a decreased number of electrowinning cells.
The hydrogen required for iron reduction is taken from a
hydrogen plant. Such a plan-t may use natural gas as
feedstock. It should be apparent and understood that in
~O each case where an anode or cathode is shown, a multi-
plicity of cells appropriate for the plant throughout is
is being considered.
The copper is electrocrystallized at the cath-
ode in the form of a loosely adherent dendritic growth.
The copper growth fa].ls from the cathode or is removed
by periodic shaking. The size to which the copper grows
before falling from the cathode is dependent on the
nature of the cathode material as shown in Table II.
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TABLE II
Influence on Cathode Material on Copper Product
MESH CUMl1LATIVE WEIGHT AND RETAINED
TYLER SERIES Stainless Steel
Titanium Copper (Type 317)
4 7.4 0.8
8 13.1 2.8
14 37.4 17.0
28 0.1 ~65.8 43.9
48 3.4 84.7 69.2
100 15.2 93.4 84.4
200 35.6 97.2 92.5
325 49.3
The variation in copper product size obtained
with the various cathode materials leads to alternate
methods of copper product removal from the circuit. The
coarse material present with a copper cathode is unde-
sirable in that it indicates a tendency for continued
growth and this ultimately could result in damage to the
~0 cell membrane. The product from the stainless steel
cathode settles rapidly and must be filtered directly.
The product from a titanium cathode can be pumped to a
thickener prior to being filtered.
The catholyte solution which has been
sufficiently depleted in copper goes to a catholyte
scrubber 20 wherein the solution is used to absorb
hydrochloric acid from gases produced in the iron
reduction stage 22. Addition of hydrochloric acid to
this stream is beneficial since it is directed to the
anode compartment of the electrolytic cells and as
89
previously discussed, increased acid concentration is
beneficial to electrowinning efficiency.
In the anode compartments, Eerrous iron is
oxidized to the ferric state to balance the electrical
requirement for copper reduction at the cathode. Only
one-third of the iron is oxidized during copper reduc-
tion. The rest of the iron must be oxidized in a sepa-
rate stage. This oxidation can be carried out in the
anode compartment of additional electrowinning cells
while hydrogen is produced at the cathode as shown in
Figure 3. Figure 3 shows these cells being maintained
at a potential of 2 volts. The results of electro-
oxidation tests carried out at various cell voltages are
shown in Figure 5. It is apparent that as the cell
voltage increases the cell current and therefore the
rate of ferric production also increase. The production
of chlorine at the anode was not detected until less
than 10 grams per litre of ferrous remained in solution.
The current efficiency for ferric production is there-
fore 100% until this concentration is reached and thepower consumption for ferric production will be directly
proportional to the cell voltage. Decreasing the cell
voltage to minimize power consumption also decreases the
ferric production rate and therefore increases the
number of cells required. The optimum cell voltage will
therefore be an economic decision based on a balance
between both power costs and the cost of an increased
number of cells.
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Alternatively, the remaining ferrous may be
oxidized to the ferric state through the use of turbo-
aeration as illustrated in ~igure 4. Turboaeration is
the oxidation of ferrous to ferric brought about by
blowing air through the heated anolyte discharge
solution while agitating the solution vigorously.
That portion of the catholyte which has heen
completely stripped of copper in the copper stripper 24
is sent to the crystallizer 26 for the crystallization
of FeC12.4H2O. While the presence of sodium
chloride in solution, such as in United States Patent
No. 4,159,232, results in a complex two-stage crystal-
lization procedure, the solution of this process
requires only the removal of water for crystallization
to occur.
Figure 6 shows results obtained by removing
water from a stripped catholyte solution. When approxi-
mately 30% of the volume has been removed as vapour,
crystallization begins. Continued removal of vapour
results in an increased proportion of the iron becoming
crystallized, In actual practice the quantity of
solution which goes to the crystallizer 26 is greater
than that required for removal of the appropriate amount
of iron. The advantage of this is that it provides an
outlet for water from the circuit without th0 loss of
dissolved constituents. A recycle stream from the
crystallizer 26 carries the excess iron back to joln the
solution going to the anode compartment 16. The recycle
stream is enriched in dissolved species such as HCl.
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The ferrous chloride tetrahydrate crystals
from the crystallizer 26 are washed in a crystal wash 28
and dried in a crystal drier 30 to the dihydrate state
prior to being hydrogen reduced in an iron reductor 22.
The hydrogen is obtained from a suitable hydrogen source
32.
Hydrogen reduction has been carried out on
pelletized crystals produced from the process solutions.
The reduction of pellets was found to be complete after
three to four hours reduction at 600C. The reduced
pellets are friable in nature and are readily pulverized
and then briquetted. Alternatively, the solution from
the copper stripper 24 can be oxidized to produce ferric
oxide and recover the hydrogen chloride, such treatment
of ferrous chloride liquors being known art in the
steelmaking industry.
As will be apparent to those skilled in the
art in the light of the foregoing disclosure, many
alterations and modifications are possible in the
practice of this invention without depar~ing from the
spirit or scope thereof. Accordingly, the scope of the
invention is ~o be construed in accordance with the
substance defined by the following claims.
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