Note: Descriptions are shown in the official language in which they were submitted.
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TITLE OF THE INVENTION
[0001] Electrochemical process for the recovery of metallic iron and
chlorine values from iron-rich metal chloride wastes.
FIELD OF THE INVENTION
[0002] The present invention relates to an electrochemical process for the
recovery of metallic iron and chlorine values from iron-rich metal chloride
wastes.
More specifically, the present invention is concerned with an electrochemical
process for the recovery of metallic iron and chlorine values from iron-rich
metal
chloride wastes such as carbo-chlorination wastes, spent acid leaching
liquors,
pickling liquors, or any other iron-rich metal chloride liquor or solution.
BACKGROUND OF THE INVENTION
[0003] In the chemical industries, chlorine gas (CI2) is one of the most
widely used inorganic chemicals. For example, polyurethanes, halogenated
hydrocarbons and white titanium dioxide pigment are commonly manufactured in
processes using chlorine gas.
[0004] In the latter case of white titanium dioxide pigment manufacture,
feedstock is chlorinated with chlorine gas. Chlorinated species are reduced to
waste by-products such as: hydrogen chloride (HClgas), hydrochloric acid
(HCIaq) or
inorganic metal chlorides (e.g., FeCI3, FeCI2, MgC12).
[0005] In particular, when titanium tetrachloride (TiCI4) is prepared by the
carbo-chlorination of titaniferous ores feedstock (e.g., weathered ilmenite,
titanium
slag or synthetic rutiles), significant amounts of iron and metal chlorides
species
are generated as by-products. These by-products may comprise either ferrous or
ferric chlorides or a combination thereof, depending on the reaction
conditions of
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the chlorinator. The actual by-products are in fact more complex as these
consist
of a chlorination waste which is essentially made of a blend of particulate
iron
chlorides contaminated with unreacted titanium feedstocks, petroleum coke,
silica
and silicates, and other metal chlorides. The approximate chemical composition
of
the metal chlorides collected from the cyclones of chlorinators operating with
titanium slag only is presented in Table 1 below.
Table 1- Average composition ranges of the metal chlorides in an as-
received chlorinator dust, expressed as anhydrous salts (wt.%)
Metal chlorides Formula Percentage
Iron II chloride FeCI2 30-70
Aluminum III chloride AICI3 5-15
Magnesium II chloride M CIz 5-20
Manganese II chloride MnCl2 4-15
Sodium chloride NaCi 1-8
Vanadium (IV) oxychloride VOCI2 1-6
Chromium III chloride CrC13 0.5-6
Titanium III chloride TiCI3 0.1-3
[0006] The formation of these chlorinator wastes has severe economic and
environmental implications on the overall process because the wastes must be
processed for disposal. Usually, by-product iron chlorides are dumped in large
scale deep wells or at sea landfills or simply discharged into wastewater
stream.
Such discarding involves both environmental issues and a complete loss of the
economic value of the chlorine species. Despite being environmentally unsound,
these practices are still extensively used at many plant locations, worldwide.
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[0007] Although attempts have been made to commercialize these by-
metallic chloride products as flocculating agent in the treatment of
wastewater or
as etching agent in pickling baths, these attempts are hampered by the low
market
value of these by-products. In addition, since the by-products are usually in
the
form of aqueous solutions, transportation charges are prohibitive.
[0008] For these reasons, there has been extensive research on chlorine
recycling and various attempts have been made over the past forty years in the
titanium dioxide pigment industry to recover the chlorine values from iron
chlorides.
[0009] In addition, since the introduction in 1998 of the upgrading of
titanium slag by high pressure hydrochloric acid leaching, an increasing
interest
has arose in recovering chlorinated metal values from the spent acid. At
present
the spent acid is pyro-hydrolysed to regenerate an azeotropic solution of
hydrochloric acid leaving behind inert metals oxides that are landfilled as
mining
residues. The average composition ranges of a spent acid is presented in Table
2
below.
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Table 2 - Average composition ranges of spent acid
Cations or Concentration
chemicals (c/g.dm3)
HCI (free) 40-70
Fe(total) 30-60
Fe(I1) 20-45
Mg(II) 10-30
AI(III) 4-12
Fe(III) 4-12
Ca( I I ) 0.5-2
V( I 11) 0.5-2
Mn( I I ) 0.5-3
Cr( I 11) 0.3-2
Ti IV 0.1-1
[0010] Until today, there is an absence of a satisfactory industrial process
for recovering elemental chlorine from iron chlorides. The main prior art
route for
recovering chlorine from spent chlorides is the thermo chemical oxidation of
iron
chlorides in an excess of oxygen.
[0011] Thus, several attempts have centered around the oxidation of iron
chlorides during which the following chemical reactions are involved:
2 FeCI2(s) + 3/2 02(g) --> Fe203(s)+ 2 C12(g)
2 FeC13(s) + 3/2 02(g) 4 Fe203(s)+ 3 CI2(g)
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[0012] However, until today it has proved very difficult to develop a
satisfactory industrial process incorporating the reaction exemplified in the
previous equations. Many efforts have been made to overcome the attendant
difficulties by conducting the reaction in the gaseous phase such as indicated
by
Harris et al.'. Harris suggested that ferric chloride can be treated with
oxygen in a
fluidized-bed reactor in the vapor phase. The process produces chlorine gas,
which can be recycled to an ilmenite or rutile chlorination process, and iron
oxide
by-product rather than soluble chloride wastes.
[0013] GB Patent 1,407,034 2 discloses oxidation of gaseous ferrous
chloride with oxygen in excess at temperatures sufficiently high to avoid
condensation of the ferrous chloride.
[0014] US Patent 3,865,9203 to RZM Ltd., discloses a process consisting in
preheating ferrous chloride at 980 C to 1110 C and then oxidizing it by
passing
pure oxygen to form a mixture of iron chlorides, iron oxide, oxygen and
chlorine,
which mixture is thereafter cooled and the residual iron chloride converted to
iron
oxide and chlorine.
[0015] The main issues with the full oxidation of either FeCI2 or FeCI3 to
iron
oxides and chlorine is that thermodynamics requires low temperatures, i.e.,
usually
below 400 C, to shift the equilibrium in favor of the oxidation of the ferric
chloride.
However it appears that, at low temperatures imposed by thermodynamics, the
reaction kinetics becomes too slow whereas at higher temperatures, where the
reaction proceeds at a practical rate, the reaction is far from complete.
[0016] It was subsequently found that the utilization of a catalyst such as
iron oxide accelerates the reaction at lower temperatures. Thus the use of an
iron
oxide fluidized bed reactor was proposed to lower the reaction temperatures.
Actually, US Patent 2,954,2744 to Columbia Southern Chemical Corp. proposed to
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oxidize ferrous iron chloride by means of air or oxygen at temperatures from
400 C
to 1000 C in a fluidized bed of iron chloride and optionally iron oxide.
Later, in US
Patent 3,793,4445 to E.1 DuPont de Nemours the oxidation of gaseous iron
chloride was performed by passing a mixture of the iron chloride and oxygen
through several superposed zones subdivided by walls and in the presence of
recycled inert solid particles (e.g., silica sand). During this process,
ferrous chloride
(FeC12) is continuously oxidized, first to ferric chloride (FeC13) and then to
ferric
oxide (Fe203) in one stage. Afterwards, in US Patent 4,144,3166 to E.1 DuPont
de
Nemours, Reeves and Hack improved the process by carrying out the
dechlorination reaction in a recirculating-fluidized-bed reactor for example
of the
type suggested in US Patent 4,282,1857
.
[0017] However, an additional problem arises during thermal oxidation, that
is, the deposition of a solid, dense and hard iron oxide scale (Fe203). This
scale
has a severe tendency to accumulate and adhere strongly on the reactor walls
and
associated equipment, causing problems in the efficient operation and
maintenance of the reactor. Actually, it has been demonstrated that oxide
scale
occurs above bed level to such an extent that the outlet may become completely
clogged in a short time and the operation must be frequently stopped for
removing
the scale leading to expensive shutdowns. Moreover, serious problems were
encountered in increasing the size of the fluid bed reactor towards an
industrial
scale for this reaction.
[0018] Other proposals consisted in operating the oxidation process at
lower temperatures using a molten salt bath of NaCI to form a salt complex or
eutectic with the iron (NaCI-FeC13) compound; or conducting the oxidation
under a
pressure sufficient to effect the liquefaction of the ferric chloride.
However, these
methods generally require the use of complicated apparatus and the exercise of
very careful controls over operating conditions. Furthermore, difficulties
seem to be
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encountered in the removal of by-product iron oxide from the reactor and in
the
sticking of the particulate bed material.
[0019] Another drawback of the thermal oxidation process in general seems
to be the poor quality of the gaseous chlorine produced, namely about 75 vol%
CIZ
because it is largely contaminated with ferric chloride and other volatile
impurities
and also strongly diluted with unreacted oxygen (11 vol.% 02) and carbon
dioxide
(7.5 vol.% C02). Hence it exhibits a relatively poor commercial value. In
addition,
immediate recycling to the chlorinator as well as efforts to concentrate the
dilute
chlorine, involve great additional expenses.
[0020] Moreover, efficient chlorine recovery by thermal oxidation requires
essentially pure ferrous chloride as feedstock. However, the mechanical
separation of the particulate ferrous chloride from the major contaminants
(i.e.,
coke) in chlorinator dust is a hard task. In fact, if thermal oxidation of
impure
ferrous chloride is carried out at temperatures in excess of 800 C, the coke
present in the dust is burned up, thereby producing hot spots in the reactor,
which
leads to the sintering of the iron oxide accompanied by a build-up of the
oxide on
the walls, which in turn leads to clogging within a short time.
[0021] After the unsuccessful pilot and pre-commercial trials made by E.I.
Du Pont de Nemours for thermal oxidation, other titanium dioxide pigment
producers investigated this technology such as SCM Chemicals Ltd.8, Kronos
Titan
GmbH9 and recently Tioxide10.
[0022] Another route, namely the electrolytic route, was considered for
recovery of both chlorine and iron values.
[0023] It appears from the prior art that work has been done on the
electrodeposition of iron metal from iron-containing solutions since the
second half
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of the eighteenth century. In fact, various processes for electrowinning,
electroplating, or electrorefining iron metal are known. Usually, the aim of
these
processes is to prepare an electrolytic iron with a high purity and to a
lesser extent
pure iron powders. Usually, the most common electrolytes were based on iron
sulphate and to a lesser extent with iron chlorides.
[0024] Most of the known electrochemical processes were originally
designed to electrodeposit iron at the cathode while the anodic reaction
usually
consisted in the anodic dissolution of a soluble anode made of impure iron. In
such
processes, the use of consumable-type anodes seems to have generally allowed
avoiding an undesirable evolution of corrosive nascent oxygen or hazardous
chlorine gas.
[0025] On the anode side, chlorine recovery by electrolysis from brines or
by-produced hydrochloric acid is well-documented technology with many plants
operating worldwide with a discrete number of electrolytic processes. However
an
industrial scale electrochemical process that combines the two principles of
recovering directly both iron and chlorine from waste iron-containing
chlorides
does not seem to exist.
[0026] The first well-documented attempt apparently dates back to 1928
with the patents of LEVY10. The inventor disclosed a simple electrochemical
process for recovering both nascent chlorine and pure electrolytic iron from a
solution of pure ferrous chloride. The electrolyser was divided with a
diaphragm as
separator made of porous unglazed clay to prevent the mixing of products. The
electrolysis was conducted at 90-100 C under a current density of 110 - 270
A.m"2
with an average cell voltage of 2.3-3.0 V. The Faradaic current efficiency was
90-
100%. The anolyte was a concentrated chloride solution (e.g., CaCI2, NaCI)
while
the catholyte was an aqueous solution containing 20 wt.% FeCI2. The anode was
carbon-based while the cathode was a thin plate, mandrel or other suitable
object.
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[0027] More recently, in 1990, OGASAWARA et al. from Osaka Titanium Co.
Ltd (now Toho)12 disclosed in a patent application an electrolytic process to
produce iron and chlorine through the electrolysis of an iron chloride-
containing
aqueous solution (an effluent resulting from the pickling of steel or from the
process of producing titanium tetrachloride or nonferrous titanium ore) by the
use
of anion and cation exchange membranes in conjunction with a three-compartment
electrolyser. In this process as exemplified in Ogasawara, the catholyte,
which is
made of high purity ferrous chloride and constantly adjusted to a pH of 3 to 5
with
ammonia, and the anolyte made of sodium chloride, recirculate in loop inside
their
respective compartments, while the iron-rich chloride-containing solution to
be
electrolysed circulates through the central compartment, that is, the gap
existing
between the two ion-exchange membranes. The cathode used is preferably iron
but may also be stainless steel, titanium or titanium alloy, and the anode
used is
made of insoluble graphite. According to the inventors, this 3-compartment
process apparently allows, in contrast to that using a two-compartment
electrolytic
process, to avoid polluting the resulting electro-crystallized iron by
embedded
impurities such as metal oxides. In addition, maintaining the catholyte pH
between
3 and 5 allows avoiding hydrogen evolution at the cathode.
[0028] However, in such process, there appears a high ohmic drop due to
(i) the additive resistivities of the ion exchange membranes and (ii) the
associated
gap existing between the two separators. In addition, the utilization of a
graphite
anode combined with a sodium chloride brine anolyte seems to cause a high
overpotential for the reaction of chlorine evolution. Both the high ohmic drop
and
the anodic overvoltage contribute to the cell potential. This therefore leads
to a
high specific energy consumption for both chlorine and iron recovery, which is
not
compatible with a viable commercial process.
[0029] Therefore remains a need for an efficient and economical process to
recover both iron metal and chlorine gas from iron-rich metal chloride wastes.
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[0030] The present description refers to a number of documents, the
content of which is herein incorporated by reference in their entirety.
SUMMARY OF THE INVENTION
[0031] The present invention generally relates to an electrochemical
process for the recovery of metallic iron and chlorine gas from iron-rich
metal
chloride wastes.
[0032] More specifically, an aspect of the present invention relates to an
electrochemical process for the recovery of metallic iron and chlorine gas
from an
iron-rich metal chloride solution comprising the following steps:
a) providing an iron-rich metal chloride solution;
b) electrolysing the iron-rich metal chloride solution in an electrolyser
comprising a cathodic compartment equipped with a cathode having a
hydrogen overpotential higher than that of iron and containing a catholyte
having a pH below about 2, an anodic compartment equipped with an
anode and containing an anolyte, and a separator allowing for anion
passage, the electrolysing step comprising circulating the iron-rich metal
chloride solution in a non-anodic compartment of the electrolyser, thereby
causing iron to be electrodeposited at the cathode and chlorine gas to
evolve at the anode, and leaving an iron-depleted solution; and
c) separately recovering the electrodeposited iron and the chlorine gas.
[0033] In a specific embodiment, step (a) of providing an iron-rich metal
chloride solution includes the following steps:
al) leaching a solid carbo-chlorination waste with a hot aqueous
solution, thereby forming an aqueous slurry; and
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a2) subjecting the aqueous slurry to a separation of solids, thereby
forming an insoluble cake and isolating an iron-rich metal chloride
solution.
[0034] In another specific embodiment, the pH of the catholyte is adjusted
to range between about 0.3 and about 1.8, preferably between about 0.6 and
about 1.5, more preferably between about 0.6 and about 1.1, most preferably
between about 0.9 and about 1.1.
[0035] In another specific embodiment, the cathode has an overvoltage, at
200 A.m"2, greater than about 425 mV in 0.5 mol.dm"3 HCI at 25 C.
[0036] In another specific embodiment, the cathode is constructed from or
coated with a material selected from the group consisting of titanium,
titanium
alloy, zirconium, zirconium alloy, zinc, zinc alloy, cadmium, cadmium alloy,
tin, tin
alloy, copper, copper alloy, lead, lead alloy, niobium, niobium alloy, gold,
gold
alloy, mercury and metallic amalgam with mercury.
[0037] Another aspect of the present invention relates to a process for the
recovery of metallic iron and chlorine gas from an iron-rich metal chloride
solution,
which process comprises:
a) providing an iron-rich metal chloride solution;
b) electrolysing the iron-rich metal chloride solution in a two-
compartment electrolyser comprising a cathodic compartment equipped
with a cathode having a hydrogen overpotential higher than that of iron,
and an anodic compartment equipped with an anode and containing an
anolyte, the cathodic and anodic compartments being separated by an
anion-exchange membrane, the electrolysing step comprising circulating
the iron-rich metal chloride solution, adjusted to a pH below 2, as a
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catholyte in the cathodic compartment of the electrolyser, thereby causing
iron to be electrodeposited at the cathode and chlorine gas to evolve at
the anode, and leaving an iron-depleted solution; and
c) separately recovering the electrodeposited iron and the chlorine gas.
[0038] Other objects, advantages and features of the present invention will
become more apparent upon reading of the following non-restrictive description
of
specific embodiments thereof, given by way of example only with reference to
the
accompanying drawings.
BRIEF DESCRIPTION OF THE DRAWINGS
[0039] In the appended drawings:
[0040] Figure 1 is a flow-sheet diagram illustrating the various steps of the
entire electrochemical process according to a first embodiment of the present
invention, based on a two-compartment electrolyser and performing electrolysis
with a pH-adjusted iron-rich metal chloride solution;
[0041] Figure 2 is a flow-sheet diagram illustrating the various steps of the
entire electrochemical process according to a second embodiment of the present
invention, based on a two-compartment electrolyser and performing electrolysis
with a pH-adjusted iron-rich metal chloride solution from which the vanadium
has
been removed by precipitation prior to its introduction in the cathodic
compartment;
[0042] Figure 3 is a flow-sheet diagram illustrating the various steps of the
entire electrochemical process according to a third embodiment of the present
invention, using a three-compartment electrolyser and performing electrolysis
with
a non-adjusted iron-rich metal chloride solution;
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[0043] Figure 4 is a schematic illustration of a two-compartment electrolyser
used in some embodiments of the present invention with major electrochemical
reactions occurring at each electrode;
[0044] Figure 5 is a schematic illustration of a three-compartment
electrolyser used in some embodiments of the present invention with major
electrochemical reactions occurring at each electrode;
[0045] Figure 6 is a photograph obtained by a scanning electron
microscope (SEM) showing an overview of a co-deposition of iron and vanadium,
as obtained in Example 2a;
[0046] Figure 7 is a photograph obtained by a scanning electron
microscope (SEM) showing a detail view of a co-deposition of iron and vanadium
pentoxide, as obtained in Example 2a;
[0047] Figure 8 is a photograph showing a smooth iron electrodeposit with
a small amount of vanadium, as obtained in Example 2b;
[0048] Figure 9 is a photograph showing an electrodeposited thin plate of
iron metal, as obtained in Example 5;
[0049] Figure 10 is a photograph showing an iron metal deposit plate, as
obtained in Example 6;
[0050] Figure 11 is a graphical illustration showing the polarization curves
as obtained in Example 8 (selection of a cathode material);
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[0051] Figure 12 is a graphical illustration showing the polarization curves
as obtained in Example 9 (selection of an anion exchange membrane); and
[0052] Figure 13 is a graphical illustration showing the polarization curves
as obtained in Example 10 (selection of an anolyte).
DESCRIPTION OF ILLUSTRATIVE EMBODIMENTS
[0053] Various feedstocks may be used in a process according to the
present invention, including, but not limited to, carbo-chlorination wastes,
for
example from carbo-chlorination of titaniferous ores, spent acid leaching
liquors,
pickling liquors or any other iron-rich metal chloride liquor or solution.
Thus the
feedstock may be solid, anhydrous, in slurry form or in solution.
[0054] As used herein, the term "electrolyser" generally designates a two-
compartment or three-compartment electrolyser. All electrolysers used in the
process of the present invention at least comprise an anodic compartment and a
cathodic compartment, separated by at least one ion exchange membrane.
[0055] As used herein when referring to an electrolyser, the term "non-
anodic compartment" designates the cathodic compartment of a two-compartment
electrolyser and/or the central compartment of a three-compartment
electrolyser.
For more clarity, it does not designate the cathodic compartment of a three-
compartment electrolyser.
[0056] As used herein, the term overpotential (also known as overvoltage)
generally designates the difference between the electrical potential of an
electrode
under the passage of current and the thermodynamic value of the electrode
potential in the absence of electrolysis for the same experimental conditions.
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[0057] As used herein when referring to a cathode, the term "hydrogen
overpotential" designates an overpotential associated with the liberation of
hydrogen gas at the cathode. A cathode having high hydrogen overpotential
minimizes hydrogen evolution during electrolysis, and thus facilitates iron
electrodeposition. Known and non-limiting examples of materials having high
hydrogen overpotential are given, for example, in Cardarelli13 and in US
Patent
5,911,869 to Exxon Research and Engineering and Co.14. Advantageously, the
cathode material also allows stripping of the iron metal deposit. Non limiting
examples of suitable cathode materials include titanium (of commercial or
higher
purity), titanium alloy (for example titanium palladium ASTM grade 7),
zirconium (of
commercial or higher purity), zirconium alloy, zinc (of commercial or higher
purity),
zinc alloy, cadmium (of commercial or higher purity), cadmium alloy, tin (of
commercial or higher purity), tin alloy, copper (of commercial or higher
purity),
copper alloy, lead (of commercial or higher purity), lead alloy, niobium (of
commercial or higher purity), niobium alloy, gold (of commercial or higher
purity),
gold alloy, mercury or metallic amalgam with mercury.
[0058] It is to be understood that a cathode having high hydrogen
overpotential may consist of a bulk of a material having high hydrogen
overpotential or may simply be coated with such a material.
[0059] As used herein when qualifying a cathode, the expression "having a
hydrogen overpotential higher than that of iron" means that, in absolute
value, the
cathode has an overvoltage, at 200 A.m"2, greater than about 425 mV in 0.5
mol.dm"3 HCI at 25 C.
[0060] It is to be understood that the relevance of performing some optional
steps of the process according to the present invention depends on the
presence
in the feedstock of given elements to be recovered. For example, not all
feedstocks possibly useable in a process according to the present invention
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contain vanadium. Of course, a vanadium-separation step is only relevant if
vanadium is present in the feedstock.
[0061] As used herein, the expression "vanadium-separation step"
essentially designates a step wherein vanadium is separated from iron. Thus it
may correspond to, but it is not necessarily a step wherein vanadium gets
recovered as a substantially pure vanadium compound.
[0062] In an embodiment wherein the feedstock is in a solid and/or
anhydrous form, the process generally first consists in leaching the
feedstock,
such as an anhydrous chlorinator dust by-produced during carbo-chlorination of
titania-rich feedstocks (e.g., weathered ilmenite, titanium slag, natural and
synthetic rutiles), with either one of: hot acidic process water, hot diluted
hydrochloric acid, hot spent acid coming from the high pressure acid leaching
of
titanium slags or even from spent liquors by-produced during the pickling of
steel.
After complete dissolution of all metal chlorides, the resulting slurry is
filtered to
separate the remaining insoluble solids comprising unreacted titania slag,
silica
and silicates, titanium dioxide fines and coke fractions from soluble metal
chlorides
in the form of an iron-rich metal chloride liquor or solution. The filter cake
obtained
is carefully washed with a minimum of acidic water, dewatered, dried and
eventually sent back to the carbo-chlorination plant or discarded and
landfilled
(depending on its titanium and coke values and content of silica), while the
wash
water may be reused in the first leaching step.
[0063] In another embodiment, wherein the feedstock is in the form of a
slurry, the leaching may help dissolve the soluble solids before a solid-
liquid
separation, for example by filtration.
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[0064] In still another embodiment, wherein the feedstock is in a clear
aqueous liquid form, i.e. that of an iron-rich metal chloride solution, the
leaching
step is of no particular interest.
[0065] Afterwards, three main process variants can be used for recovering
both chlorine and metal values from the iron-rich metal chloride solution,
based on
the same general principle of simultaneous recovery of metal iron and chlorine
values from an iron-rich metal chloride solution by electrolysis, using a
catholyte
adjusted to a pH below 2 and a cathode having a hydrogen overpotential higher
than that of iron.
[0066] In a particular embodiment of the process according to the present
invention, as illustrated in Figure 1, the pH of the iron-rich metal chloride
solution
is first adjusted to between about 0.6 and about 1.8, with alkaline reagents
such
as, but not limited to, magnesia or ammonium hydroxide or a mixture thereof,
after
which the solution is ready for electrolysis.
[0067] Still in reference to Figure 1, the electrolytic stage consists in
circulating the pH-adjusted iron-rich metal chloride solution inside the
cathodic
compartment of an electrolyser. The iron-rich metal chloride solution thus
acts as
catholyte. The electrolyser consists of two compartments separated by an anion-
exchange membrane (as illustrated in Figure 4). The cathodic compartment
comprises a cathode made of titanium or titanium alloy (usually ASTM grade 7),
while the anodic compartment has a dimensionally stable anode for the
evolution
of chlorine (DSAT""-C12). The anolyte that circulates in loop in the anodic
compartment is made of a mixture of about 20 wt.% hydrochloric acid and about
17 wt.% magnesium chloride with about 10,000 ppm of ferric iron (Fe3+) as
corrosion inhibitor.
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[0068] During electrolysis, at the above-mentioned pH ranging between
about 0.6 and about 1.8, iron metal deposits at the cathode along with
precipitated
crystals of vanadium pentoxide. The precipitation of vanadium pentoxide
results
from the consumption of hydrogen cations at the cathode and local increase of
the
pH above the precipitation point of hydrated vanadium pentoxide. On the other
hand, chloride anions migrate through the permeable anion exchange membrane
towards the anodic compartment and discharge as chlorine gas at the surface of
the anode according to the following electrochemical reactions:
FeZ+(aq) + 2e" -- Fe (s) (cathode, -)
2CI"(aq) -- C12(g) + 2e' (anode, +)
[0069] The overall reaction therefore being:
FeCI2 -- Fe(s) + C12(g)
[0070] Side-reactions may also occur, such as the evolution of oxygen at
the anode:
2H20(I) -' 02(g) + 4H+(aq) + 4e ,
hydrogen evolution at the cathode:
2H'(aq) + 2e" -- H2(g),
along with the reduction of traces of ferric cations:
Fe3+(aq) + e" -+ Fe2+(aq).
[0071] On the cathode side, these undesired side reactions are minimized
by maintaining the pH of the catholyte below pH of about 2 and by using a
cathode
material having a high overpotential for the discharge of hydrogen so as to
prevent
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hydrogen evolution. More specifically, the cathode materials used in the
process
according to the present invention have hydrogen overpotential higher (in
absolute
value) than that of iron in given electrolysis conditions. Preferably, the pH
of the
catholyte is maintained between about 0.6 and about 1.8, more preferably
between
about 0.6 and about 1.5, still more preferably between about 0.6 and about
1.1,
and most preferably between about 0.9 and 1.1. In addition, using an inert
atmosphere of nitrogen above the cathodic compartment may help preventing the
oxidation of the ferrous cations.
[0072] On the anode side, the utilization of a dimensionally stable anode for
chlorine evolution may impede the evolution of oxygen gas, thereby ensuring
the
production of a high purity chlorine gas.
[0073] The electrolysis is usually conducted between about 40 C and about
110 C under a gaivanostatic control. The overall current density is comprised
between about 200 and about 2000 A/m2 with a cell voltage ranging from about
1.2
to about 3.5 V per cell. In this specific embodiment, the faradaic efficiency
is
usually greater than about 90% and the average specific energy consumption is
between about 2.1 and about 6.2 kWh per kg of iron and between about 1.1 and
about 3.5 kWh per kilogram of chlorine gas.
[0074] The wet chlorine gas evolved is recovered by conventional methods.
For example, as shown in Figure 1, it may be recovered by suction, cooled by
passing it through a graphite heat exchanger, and dried by passing it through
a
mist eliminator and several concentrated sulfuric acid spray-towers
(scrubbing).
Finally the dry and cold chlorine gas may be compressed and liquefied, thus
being
ready to be transported or stored on-site for future use.
[0075] The thick plates of electrodeposited iron metal are mechanically
stripped from the titanium cathode. The plates are then immersed into a hot
lye of
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concentrated sodium hydroxide (50 wt.% NaOH) to selectively dissolve the
vanadium oxides; traces of oxydiser, such as, but not limited to, potassium
chlorate, are added to convert all the vanadium into pentavalent vanadium and
pure iron metal is separately recovered. Ammonia along with ammonium chloride
(NH4CI) and/or ammonium hydroxide are then added to the remaining liquor in
order to precipitate all the vanadium as ammonium metavanadate (NH4VO3). Thus
in such specific embodiment, a vanadium-separation step occurs after the
electrolysis step.
[0076] Sulfuric acid is added to the spent iron-free electrolyte, or iron-
depleted solution, exiting the electrolyser, for removing calcium as insoluble
calcium sulfate dihydrate (CaSO4.2H20) and entraining optional traces of
radioactivity, mostly as radium sulfate.
[0077] The remaining spent magnesium- and aluminum-rich brine is then
pyro-hydrolysed to yield refractory spinel beads, pellets or granules ready to
be
used in the manufacture of refractories or proppants, while recovering
azeotropic
hydrochloric acid.
[0078] It is to be understood that changing the pH of the catholyte in the
process of Figure 1, for example to 0.3 to 0.5, would allow vanadium not to
precipitate along with iron codeposition but to remain in the iron-rich,
becoming the
iron-depleted solution, thus performing a vanadium separation step during
electrolysis. This is however not a preferred embodiment in a process using a
two-
compartment electrolyser since the iron obtained may be, although slightly,
contaminated by vanadium pentoxide and the Faradaic efficiency may drop.
[0079] In another particular embodiment of the process according to the
present invention, as generally illustrated in Figure 2, the exact vanadium
content
of the iron-rich metal chloride solution is determined by a conventional
method and
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21
a stoichiometric amount of potassium chlorate (KCI03) is introduced to oxidize
all
the vanadium into vanadium (V) (not shown). A corresponding amount of iron
(III)
chloride is then added and the pH of the solution is adjusted to between about
0.5
and about 3 with alkaline reagents such as for instance magnesia or ammonium
oxide, hydroxide or a mixture thereof. This precipitates together vanadium (V)
and
chromium (VI), entrained by co-precipitation with the ferric hydroxide
(Fe(OH)3).
The gelatinous vanadium-rich precipitate is then removed from the slurry by a
known technique of either decantation, centrifugation or filtration. The so-
obtained
vanadium-rich precipitate, for example in the form of a filter cake, is then
dissolved
in a minimum amount of concentrated solution of sodium hydroxide and oxidised
with traces of oxydiser. The remaining ferric and chromic hydroxides are
discarded
and the vanadium is selectively precipitated as ammonium metavanadate
(NH4VO3) by addition of ammonium hydroxide (NH4OH) and/or ammonium chloride
(NH4CI), and recovered.
[0080] The clear filtrate or supernatant from the vanadium separation step
is pH-adjusted at a pH below 2, preferably between about 0.6 and about 1.8 and
thus ready for electrolysis, in the form of a vanadium-depleted and pH
adjusted
iron-rich metal chloride solution (not shown).
[0081] Still in reference to Figure 2, the electrolysis consists in
circulating
the vanadium-depleted and pH-adjusted iron-rich metal chloride solution inside
the
cathodic compartment of an electrolyser. The iron-rich metal chloride solution
thus
acts as catholyte. Similarly to Figure 1, the electrolyser consists of a cell
divided by
an anion-exchange membrane (as illustrated in Figure 4). The cathodic
compartment has a cathode made of titanium metal or a titanium alloy (usually
ASTM grade 7). The anodic compartment has a dimensionally stable anode for the
evolution of chlorine (DSAT"^-CI2). The anolyte that circulates in loop is
made of a
mixture of about 20 wt.% hydrochloric acid and about 17 wt.% magnesium
chloride
with about 10,000 ppm of ferric iron (Fe3+) as corrosion inhibitor. During
electrolysis, pure iron metal is deposited at the cathode, while chloride
anions
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22
migrate through the permeable anion exchange membrane to the anodic
compartment and discharge as chlorine gas at the surface of the anode
according
to the following electrochemical reactions:
Fe2+(aq) + 2e -- Fe (s) (cathode, -)
2CI"(aq) -+ CI2(g) + 2e" (anode, +)
[0082] The overall reaction being:
FeCI2 -- Fe(s) + C12(g).
[0083] Again, side-reactions may also occur, such as the evolution of
oxygen at the anode:
2H20(I) -' 02(g) + 4H+(aq) + 4e",
hydrogen evolution at the cathode:
2H+(aq) + 2e" -4 Hz(g),
along with the reduction of traces of ferric cations:
Fe3+(aq) + e" -- Fe2+(aq).
[0084] Again, on the cathode side, these undesired side reactions are
minimized by maintaining the pH of the catholyte below 2 and by using a
cathode
material having high hydrogen overpotential. The cathode materials suitable
for
use in the process according to the present invention have a hydrogen
overpotential higher (in absolute value) than that of iron in given
electrolysis
conditions. Preferably, the pH of the catholyte is maintained between about
0.6
and about 1.8, more preferably between about 0.6 and about 1.5, still more
preferably between about 0.6 and about 1.1, and most preferably between about
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23
0.9 and 1.1. In addition, using an inert atmosphere of nitrogen above the
cathodic
compartment may help preventing the oxidation of the ferrous cations.
[0085] On the anode side, the utilization of a dimensionally stable anode for
chlorine evolution may impede the evolution of oxygen gas, thereby ensuring
the
production of a high purity chlorine gas.
[0086] In the embodiment of Figure 2, the electrolysis is usually conducted
between about 40 C and about 110 C under a galvanostatic control. The overall
current density is comprised between about 200 and about 2000 A/m2 with a cell
voltage ranging from about 1.9 to about 3.5 V per cell. In this specific
embodiment,
the faradaic efficiency is usually greater than 90% and the specific energy
consumption is usually between about 2 and about 3.7 kWh per kg of iron and
between about1.6 and about 3 kWh per kilogram of chlorine gas.
[0087] In this specific embodiment, the wet chlorine gas evolved is
recovered by suction, is cooled by passing it through a graphite heat
exchanger,
and dried by passing it through a mist eliminator and several concentrated
sulfuric
acid spray-towers (scrubbing). Finally the dry and cold chlorine gas is
compressed
and liquefied, thus being ready to be transported or stored on-site for future
re-
utilization.
[0088] The thick electrodeposited plates of pure iron metal are mechanically
stripped from the titanium cathode.
[0089] Concentrated sulfuric acid is added to the spent iron-free electrolyte,
or iron-depleted solution, exiting the electrolyser for removing calcium as
insoluble
calcium sulfate dihydrate (CaSO4.2H20) and entraining optional traces of
radioactivity, mostly as radium sulfate.
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24
[0090] The remaining spent magnesium- and aluminum-rich brine is then
pyrohydrolysed to yield refractory spinel beads, pellets or granules ready to
be
used in the manufacture of refractories or proppants while recovering
azeotropic
hydrochloric acid.
[0091] In another particular embodiment of the process according to the
present invention, as illustrated in Figure 3, the iron-rich metal chloride
solution is
sent without any prior treatment (such as pH adjustment) to the
electrochemical
plant. The electrolyser design used in this process (as illustrated in Figure
5) has
three compartments: (i) a cathodic compartment with a titanium plate cathode,
(ii)
an anodic compartment comprising a dimensionally stable anode for the
evolution
of chlorine, and (iii) a central compartment separated from the cathodic
compartment by a cation-exchange membrane and from the anodic compartment
by an anion exchange membrane. The catholyte circulating inside the cathodic
compartment is a saturated solution of ferrous chloride (about 350 g/L FeCI2)
with
magnesium chloride (about 220 g/L MgCI2), while the anolyte is made of about
20
wt.% hydrochloric acid and about 17 wt.% magnesium chloride with about 10,000
ppm of ferric iron (Fe3+) as corrosion inhibitor. The pH of the catholyte is
adjusted
below pH 2, preferably between about 0.6 and about 1.8, more preferably
between
about 0.6 and about 1.5, still more preferably between about 0.6 and about
1.1,
most preferably between about 0.9 and about 1.1. The iron-rich metal chloride
solution is passed through the central compartment continuously. During the
electrolysis (Figure 5), ferrous cations of the iron-rich metal chloride
solution
migrate through the cation exchange membrane and are reduced to pure iron
metal onto the titanium cathode while the chloride anions migrate through the
anion exchange membrane towards the dimensionally stable anode where they
are oxidized, thereby producing chlorine gas that evolves. The electrochemical
reactions involved are as follows:
Fe2+(aq) + 2e" -- Fe (s) (cathode, -)
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2CI"(aq) -- C12(g) + 2e" (anode, +)
[0092] The overall reaction being:
FeCI2 -- Fe(s) + CI2(g).
[0093] The electrolysis is conducted between about 40 and about 110 C
under gaivanostatic control with an overall current density comprised between
about 200 and about 2000 A/m2 with a cell voltage ranging from about 1.9 to
about
3.5 V per cell. In this embodiment, the faradaic efficiency is usually greater
than
about 90%.
[0094] In this embodiment, the pure and wet chlorine gas evolved is
recovered by suction, is cooled by passing it through a graphite heat
exchanger
and dried by passing it through a mist eliminator and several concentrated
sulfuric
acid spray-towers. Finally the dry and cold chlorine gas is compressed and
then
liquefied, thus being ready to be transported or stored on-site for future
utilization.
[0095] The thick plates of electrodeposited pure iron metal are mechanically
stripped from the titanium cathode.
[0096] Hydrogen peroxide (H202) is added to the iron-depleted solution
exiting the central compartment to oxidize all the traces of vanadium (IV, and
V) to
vanadium (V). Then magnesium oxide (MgO) is added to adjust the pH to about
1.8-2.2, which leads to the quantitative precipitation of hydrated vanadium
pentoxide (V205.250H20). The precipitate is removed by decantation, filtration
or
centrifugation, dried and calcined to yield flakes of vanadium pentoxide
(V205) (not
shown).
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26
[0097] Afterwards, sulfuric acid is added to the resulting iron and vanadium-
free brines for removing calcium as insoluble calcium sulfate dihydrate and
entraining traces of radioactivity, mostly as radium. The spent magnesium- and
aluminum-rich brine is then pyrohydrolysed to yield refractory spinel beads,
pellets
or granules ready to be used in the manufacture of refractories or proppants,
while
recovering azeotropic hydrochloric acid.
[0098] It is to be noted that the pH of the iron-rich metal chloride solution
may or may not be adjusted prior to electrolysis when using a three-
compartment
electrolyser. Such an adjustment could, for example, serve to effect a
vanadium
precipitation along with iron deposition, as above, although it is not a
preferred
embodiment here.
[0099] A number of parameters of the process according to the present
invention may be varied, as explained below.
[00100] Cathode materials suitable for use in the process of the present
invention (as bulk or coating materials) are materials having a high
overpotential
for the evolution of hydrogen, more specifically a hydrogen overpotential
higher
than that of iron in given electrolysis conditions. Advantageously, the
cathode
material also allows stripping of the iron metal deposit. Non limiting
examples of
suitable cathode materials include titanium (of commercial or higher purity),
titanium alloy (for example titanium palladium ASTM grade 7), zirconium (of
commercial or higher purity), zirconium alloy, zinc (of commercial or higher
purity),
zinc alloy, cadmium (of commercial or higher purity), cadmium alloy, tin (of
commercial or higher purity), tin alloy, copper (of commercial or higher
purity),
copper alloy, lead (of commercial or higher purity), lead alloy, niobium (of
commercial or higher purity), niobium alloy, gold (of commercial or higher
purity),
gold alloy, mercury or metallic amalgam with mercury.
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27
[0100] Anode materials suitable for use in the process of the present
invention include (as bulk or coating materials) (1) dimensionally stable
anodes for
the evolution of chlorine (DSATM-CI2) of the type [M/MXOy AOt] made of a
metallic
substrate or base metal M coated with a mixed metal oxides (MMO) as
electrocatalyst, wherein M is a refractory metal or an alloy with a valve
action
property such as titanium, titanium alloy, zirconium, zirconium alloy,
hafnium,
hafnium alloy, vanadium, vanadium alloy, niobium, niobium alloy, tantalum,
tantalum alloy, MXOy is a metallic oxide of a valve metal forming a thin and
impervious layer protecting the base metal such as Ti02, Zr02, Hf02, Nb02,
Nb205,
TaO2, and Ta205, and AZOt is an electrocatalytic metal oxide of a noble metal
or
more often an oxide of the platinum group metals (PGMs) such as Ru02, IrO2,
PtOx
and also sometimes a metallic oxide such as Sn02, Sb205, Bi203; (2) Bulk
electronically conductive ceramics such as: sub-stoichiometric titanium oxides
such as Magneli-Anderson phases with general formula Ti,O2n_1 (n is an integer
>=
3), conductive oxides with the spinel structure (AB204, wherein A = Fe(II),
Mn(II) or
Ni(II), and B = Al, Fe(III), Cr(III), Co(III)) or conductive oxides with the
perovskite
structure (AB03 , wherein A = Fe(II), Mn(II), Co(II) or Ni(II), and B =
Ti(IV)) or with
the pyrochlore structure AB207; or (3) carbon-based materials such as
graphite,
impervious graphite, or vitreous carbon.
[0101] The anolyte composition used in the process of the present invention
advantageously comprises hydrochloric acid, a salt such as MgCI2, NaCI, CaCI2
or
mixtures thereof and Fe(III) as corrosion inhibitor. For example, suitable
anolyte
compositions may vary in the following ranges: about 10 to about 37 wt.%
hydrochloric acid (preferably about 20%); about 1 to about 20 wt.% MgCI2,
NaCI,
KCI, LiCI, CaCI2 or mixtures thereof (preferably about 16%) with about 10 to
about
12,000 ppm wt. Fe(III) as corrosion inhibitor (preferably 8,000 to 10,000 ppm
wt).
[0102] In an embodiment of the present invention involving a three-
compartment electrolyser, the catholyte composition may vary in the following
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28
ranges: about 1 to about 450 g/L of iron (II) chloride (preferably about 335
g/L),
about 1 to about 350 g/L MgCl2 (preferably about 250 g/L), about 1 to about
350
g/L CaC12 (preferably about 250 g/L) or about 350 g/L of a mixture of MgClz
and
CaCI2 (preferably about 250 g/L); it may also further comprise 0 to about 10
g/L of
free HCI. In such embodiment, the catholyte pH generally ranges between about
0.6 and about 1.5, preferably about 0.6 to about 1.1, more preferably about
0.9 to
about 1.1.
[0103] The reaction temperature may range between about 40 and about
110 C, preferably between about 80 and about 95 C. Most preferably, the
operating temperature is about 85 C.
[0104] The volume flow rate of both anolyte and catholyte advantageously
ranges between about 0.1 and about 100 Umin, preferably between about 0.1 and
about 30 Umin. Most preferably, the volume flow rate is about 2 L/min.
[0105] The cathodic current density during electrolysis, to produce a
dendrite-free smooth deposit of iron, advantageously ranges between about 50
and about 1000 A/mz. Preferably in such case, the cathodic current density is
about 500 A/mZ.
[0106] The cathodic current density during electrolysis, to produce an iron
powder, advantageously ranges between about 3000 and about 5000 A/m2.
Preferably in such case, the cathodic current density is about 4000 A/m2.
[0107] Separators used in the process of the present invention may be
passive, such as a conventional diaphragm separator, or active such as ion
exchange membranes. Preferably, the separators used are ion exchange
membranes. Anion exchange membranes and cation exchange membranes used
in the process of the present invention are conventional membranes. Non-
limiting
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29
examples of suitable anion exchange membranes are presented in the Examples
below (Figurel2).
[0108] The interelectrode gap may also be varied, with a well-known impact
on the ohmic drop. It is advantageously about 6 mm.
[0109] The present invention is illustrated below in further details by way of
the following non-limiting examples.
EXAMPLE 1
[0110] Preparation of the iron-rich metal chloride solution and
separation of unreacted solids. A batch of 10 kilograms of anhydrous
chlorinator
dust, a by-product of carbo-chlorination of upgraded titania-rich slag (UGS)
was
provided by a titanium dioxide pigment producer. The material was first mixed
with
hot acidified water at 80 C that initially contained 10 g/L of free
hydrochloric acid
(HCI) in order to leach out all the soluble metal chlorides. After complete
dissolution of the soluble salts, the resulting warm and dense slurry was
filtered
under vacuum using large 240-mm inner diameter Buchner funnels (CoorsTek)
with a capacity of 4.5 liters each. The Buchners were installed ontop of a 10-
liter
Erlenmeyer vacuum flask (Kimax) connected to a vacuum pump. The filtration
media used were disks of ash-less filter paper No. 42 (Whatman). In order to
increase throughput, four of these Buchner-Erlenmeyer assemblies were operated
simultaneously in parallel.
[0111] The filter cakes thus obtained were carefully washed with a minimum
of hot and acidified deionised water, dewatered by acetone, placed into in a
stainless steel pan and then oven dried at 110 C overnight. From microscopic
examination and chemical analysis, the remaining insoluble solids comprised
mainly unreacted titanium slag, silica and silicates, precipitated fines of
titanium
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dioxide, and coke fractions. An example of the chemical composition of these
solids obtained after drying is given in Table 3 below.
Table 3 - Composition of insoluble solids after hot acidic water leaching, and
drying (wt.%)
Chemical component Formula Percentage
Carbon C 54.00
Titanium dioxide Ti02 21.07
Silica Si02 14.38
Iron sesquioxide Fe203 4.42
Sulfur S 1.44
Other metal oxides - 4.69
Total = 100.00
[0112] After filtration and washing completion, wash water and the four
filtrates totalized 18 L, which were collected into a large 5 US-gallons
cylindrical
tank made of polypropylene. The concentration of metal chlorides in this
initial
solution after leaching is presented in Table 4. Since the concentration of
iron (II)
chloride in the filtrate (i.e 83.7 g/L) was too low for performing the
electrolysis at a
cathodic current density sufficient to obtain a smooth deposit, the solution
was
further concentrated by evaporation into a large Erlenmeyer flask heated onto
a
hot plate (Corning). The evaporation was stopped when the volume of the
solution
was reduced by four (4.5 L). At that stage, the concentration of metal
chlorides
was greatly increased and reached 335 g/L for iron (II) chloride when sampled
at
80 C (see Table 4, concentrated solution). Hence, in order to prevent the
crystallization of ferrous chloride upon cooling at room temperature, the
solution
was immediately transferred into a 10-L jacketed glass reactor (Kimble-Contes)
heated by circulating hot water supplied by a heating bath (Lauda GmbH). The
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31
temperature of the solution was maintained at 80 C at all times. The solution
was
also acidified by adding minute amounts of concentrated hydrochloric acid to
maintain the concentration of free acid around 10 g/L. Actually, at a pH below
0.5,
the air oxidation of ferrous iron (Fez+) into ferric iron (Fe3+) is slowed
down.
Moreover, a blanket of nitrogen gas was also maintained above the solution for
the
same purpose of preventing oxidation, and small cm-size polypropylene balls
floating above the solution helped preventing an important water loss by
evaporation. The solution then prepared and stored was ready for the
subsequent
steps.
Table 4- Concentration of metal chlorides the iron-rich solutions (in g/L)
Metal chloride Formula Initial solution Concentrated After V
after leaching solution by precipitation and
evaporation pH-adjusted
(Example 1) (Example 1) (Examples 4 & 5)
Iron (II) chloride FeCI2 83.7 335 350(")
Magnesium (II) chloride MgCi2 19.7 79 200
Aluminum (III) chloride AICI3 20.3 81 70
Manganese (II) chloride MnCl2 13.4 53 35
Vanadium (V) oxychloride VOCIZ 5.7 22 0.1
Chromium (III) chloride CrCi3 2.4 9.5 0.4
Calcium (II) chloride CaC12 2.1 8.4 7.8
Free hydrochloric acid HCI 10 10 0.00
Density at 25 C kg/m3 1171 1259 1360
pH = 0.4 0.5 0.9
(*) some iron powder was added before increasing pH to convert remaining
traces of iron (III)
cations.
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EXAMPLE 2
[0113] Example 2a - Electrolysis of the initial concentrated iron-rich
metal chloride solution at pH 1.1). - The previous iron-rich metal chloride
concentrated solution from Example 1 was simply adjusted at a pH of 1.1 by
adding minute amount of magnesia and then circulated inside the cathodic
compartment of an electrolyser. The electrolyser consisted of a filter press
design
model MP cell from Electrocell AB (Sweden) with two compartments separated by
an anion-exchange membrane made of Excellion 1-200 (SnowPure LLC). The
geometric electrode and membrane surface area was 100 cm2 and the spacing
between each electrode and the separator was 6 mm.
[0114] The cathodic compartment comprised a cathode plate made of a
titanium palladium alloy (ASTM grade 7; Ti-0.15Pd) supplied by Titanium
Industries. Prior to electrolysis the cathode was chemically etched by
immersing it
into a fluoro-nitric acid mixture (70 vol% conc. HNO3, 20 vol.% conc. HF and
10
vol.% H20) and then rinsing it thoroughly with deionised water until no trace
of acid
remained.
[0115] The anodic compartment was equipped with a dimensionally stable
anode (DSAT"'-CI2) supplied by Magneto BV (Netherlands) made of a plate of a
titanium-palladium alloy substrate coated with a high loading of ruthenium
dioxide
(Ru02) acting as electrocatalyst for promoting the evolution of chlorine (Ti-
0.15Pd/Ru02). The anolyte that recirculated in loop consisted of an aqueous
solution of 20 wt.% hydrochloric acid with 17 wt.% magnesium chloride (MgCI2)
and 10,000 ppm of ferric iron (Fe3+) as corrosion inhibitor, the balance being
deionised water. The electrolysis was performed gaivanostatically at an
overall
current density of 500 A/m2. The operating temperature was 80 C and the volume
flow rate of both catholyte and anolyte was 1 Umin. At that current density,
the
measured overall cell voltage was 2.528 V. During electrolysis, pure iron
metal
deposited at the cathode. On the other hand, chloride anions migrated through
the
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33
permeable anion exchange membrane towards the anodic compartment and
discharged as chlorine gas at the surface of the anode according to the
following
electrochemical reactions:
Fe2+(aq) + 2e -+ Fe (s) (cathode, -)
2CI"(aq) -- C12(g) + 2e" (anode, +);
[0116] The overall electrochemical reaction being:
FeCI2 -- Fe(s) + C12(g)
[0117] After two hours of continuous electrolysis, the power was shut off
and the electrolyser was opened. The electrodeposited rough and blackened thin
plate was easily stripped from the titanium cathode by mechanical means. The
measured thickness was circa 0.126 mm and its mass was only 8.30 g. After
close
examination under the scanning electron microscope (SEM) it was in fact an
iron
metal electrodeposit with small, embedded grains of pure vanadium pentoxide
crystals (See Figures 6 and 7). After performing an ultimate chemical analysis
of
the bulk sample, it was made up of 68 wt.% iron and 32 wt. % vanadium
pentoxide
(V205). The codeposition of vanadium pentoxide was probably due to the fact
that
at the cathode surface, the hydronium cations (H+) were reduced to hydrogen
that
evolved, and hence locally this H+ depletion lead to an increase of pH, which
yielded a precipitation of vanadium pentoxide particles, embedded into the
iron
electrodeposit. From these experimental figures, the estimated faradaic
current
efficiency was 80% and the specific energy consumption at 500 A/m2 was 3.033
kWh per kg of deposit (iron + vanadium pentoxide) or 4.460 kWh per kg of pure
iron.
[0118] The wet chlorine gas evolved was recovered by suction using
downstream a peristaltic pump (Masterflex L/S Digital Pump) with Viton tubing.
The
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chlorine gas was first cooled by passing it through an empty washing
borosilicated
glass bottle immersed into a ice bath, the mist and moisture content were then
removed by passing the gas through several flasks filled with concentrated
sulfuric
acid (98 wt.% H2SO4), and finally the dry and cold chlorine gas was totally
absorbed into a saturated solution of potassium iodide (KI) liberating iodine
according to the following reaction:
CI2(gas)+3K+aq+3laq---3K+aq+l3aq+2Claq
[0119] After completion of the electrolysis, the free iodine was titrated by a
standardized solution of sodium thiosulfate (NaZS2O3) according to the
reaction:
4Na+aq + 2S2032"aq + K+aq + Ig aq -- 4Na+aq + S4062"aq + K+aq + 3I aq
[0120] Based on the titration, the anodic faradaic efficiency in chlorine was
established at 78%. The difference between the two current efficiencies (anode
and cathode) is most probably due to some hydrogen evolution at the cathode
and
some oxygen evolution at the anode. The anodic specific energy consumption at
500 A/m2 was hence 2.45 kWh per kilogram of pure chlorine gas (i.e., 7.652 kWh
per m3(NTP: 0 C, 101.325 kPa)).
[0121] Example 2b (Electrolysis of the initial concentrated iron-rich
metal chloride solution at pH 0.30). - As an alternative to Example 2a, the
iron-
rich metal chloride concentrated solution from Example 1 was adjusted at a
rather
low pH of 0.30, so as to prevent an increase of pH above the precipitation pH
of
vanadium pentoxide at the cathode surface, but not too low however, so as not
to
favour the evolution of hydrogen. This was done by adding and circulating
hydrochloric acid in the cathodic compartment of the electrolyser. The
electrolyser
was identical to that described in Example 2a but this time the electrolysis
was
performed galvanostatically at a current density of 1000 A/m2. At that current
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density and low pH, the measured cell voltage was 2.865 V. After one hour, a
bright and smooth electrodeposit was easily stripped from the titanium cathode
(see Figure 8). It had a mass of only 6.24 g. It was made of 99.88 wt.% iron
and
only 0.12 wt. % vanadium pentoxide (V205). From these experimental figures,
the
estimated faradaic current efficiency was 60% and the specific energy
consumption at 1000 A/m2 was 4.584 kWh per kg of iron.
[0122] The wet chlorine gas evolved was recovered by the same method as
that described in Example 2a.
EXAMPLE 3
[0123] Recovery of iron and vanadium from the iron-vanadium deposit
of Example 2a - The metallic deposit was ground into a pulverisette mill
(Fritsch)
and the resulting powder was treated under pressure with a caustic lye of
sodium
hydroxide (NaOH 50 wt.%) at 100 C for two hours into a 125 mL PTFE lined
digestion bomb (Parr Company). Upon cooling, the solution was filtrated to
recover
the insoluble iron metal fines. Then excess ammonium chloride (NH4CI) was
added
to the vanadium-rich liquor in order to precipitate the pure ammonium
metavanadate (NH4VO3). The pure ammonium metavanadate was later calcined
inside a porcelain boat in dry air at 400 C in a box furnace (Fisher Isotemp)
to give
off ammonia (NH3) and water vapor (H20), thereby yielding a red-orange powder
of vanadium pentoxide. The powder was then transferred into an Inconel
crucible
and melted at 700 C in air and the melt was cast onto a cool steel plate. The
resulting solidified black mass with a submetallic luster was then ground into
a two
disks vibratory cup mill with a hardmetal liner (Fritsch GmbH) using acetone
as
grinding aid and coolant. The product thus obtained was technical grade
vanadium
pentoxide powder.
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36
EXAMPLE 4
[0124] Removal of vanadium from the iron-rich metal chloride solution
from Example 1 prior to electrolysis - A stoechiometric amount of sodium
chlorate (NaCIO3) was added to the initial solution prepared in Example 1 to
oxidize all the vanadium cations (V4+, V5+) into pentavalent vanadium (V5+)
according to the reaction:
5VO2+ + CIO3 + 2H20 -+ 5VO2+ + 0.5CI2(g) + 4H+.
[0125] It is to be noted that the addition of sodium chlorate could also have
been done after concentration of the solution.
[0126] Afterwards, an equivalent amount of ferric chloride (FeCI3) was
introduced into the solution to enhance a co-precipitation of vanadium
pentoxide
and iron hydroxide. Such co-precipitation was used to promote complete
precipitation of vanadium. Indeed, should vanadium be the only species to
precipitate, the precipitation would stop at a vanadium concentration below
about
0.02 mol/L in the solution.
[0127] Red brown hydrated vanadium (V) pentoxide starts to precipitate at
about pH 1.8 while brown iron (III) hydroxide starts to precipitate at about
pH 2Ø
Thus, when both species are present, they co-precipitate at pH 1.8 - 2Ø In
the
present case, the pH of the solution was raised by careful addition of a
slurry of
slacked magnesia (Mg(OH)2) until the pH reached 2.0 but never above to avoid
the
precipitation of black mixed ferroso-ferric hydroxides. At that pH, the
complete co-
precipitation of hydrated vanadium pentoxide (V205=250H20) and iron (III)
hydroxide occurred in the form of a gelatinous red brown precipitate. The co-
precipitates were separated by filtration using a similar set-up to that
described in
Example 1.
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37
[0128] The resulting filtrate was then acidified again to adjust pH close to
0.5 and stored into the jacketed reactor until the next electrolysis step.
[0129] The red-brown gelatinous filter cake was removed from the filter
paper and digested into a warm caustic lye of sodium hydroxide (NaOH 50 wt.%).
Upon cooling, both solution and sludge were poured into 250 mL centrifugation
polypropylene bottles and centrifuged with a robust benchtop centrifuge (CL4
from
Thermo Electron) at 10,000 revolutions per minute. The insoluble and dense
gelatinous residue, mainly composed of iron hydroxide (Fe(OH)3), was separated
at the bottom, carefully washed with hot alkaline water (pH 10), centrifuged
again
and then discarded. Then excess ammonium chloride (NH4CI) was added to the
vanadium-rich supernatant in order to precipitate the pure ammonium
metavanadate (NH4VO3). The pure ammonium metavanadate was later calcined
inside a porcelain boat in dry air at 400 C in a box furnace (Fisher Isotemp)
to give
off ammonia (NH3) and water vapour (H20), thereby yielding a red-orange powder
of vanadium pentoxide. The powder was then transferred into an Inconel
crucible,
melted at 700 C in air and cast onto a cool steel plate. The solidified black
mass
with a submetallic luster was then ground into a two disks vibratory cup mill
with a
hardmetal liner (Fritsch GmbH) using acetone as grinding aid and coolant. The
product thus obtained was technical grade vanadium pentoxide powder containing
some chromium, iron and manganese as major impurities.
EXAMPLE 5
[0130] Electrolysis of the vanadium-free iron rich solution from
Example 4. - The iron-rich metal chloride solution from which vanadium was
removed during Example 4 was adjusted at a pH of 0.9 by adding minute amount
of magnesia and circulated inside the cathodic compartment of an electrolyser.
Its
composition prior to electrolysis is presented in Table 4, last column. The
electrolyser was identical to that described in examples 2a and 2b. The
electrolysis
was also performed galvanostatically at a current density of 200 A/m2. The
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38
operating temperature was 85 C and the volume flow rate of both catholyte and
anolyte was 1 L/min. At that current density, the measured cell voltage was
1.85 V.
After five hours of continuous electrolysis, the power was shut off and the
electrolyser was opened. The electrodeposited thin plate of iron metal was
easily
stripped from the titanium cathode by mechanical means. The thickness was
0.126 mm and its mass was 10.20 g (See Figure 9). It was a smooth and soft
material with some pitting probably due to attached hydrogen bubbles. From
these
experimental figures, the estimated faradaic current efficiency was 97.9% and
the
specific energy consumption at 200 A/mZ was only 1.87 kWh per kg of iron. The
purity of iron was 99.99 wt.% Fe with no traces of other metallic elements.
EXAMPLE 6
[0131] Electrolysis of the iron-rich metal chloride solution with a three
compartment electrolyser. - The iron-rich metal chloride concentrated solution
from Example 1 was simply adjusted at a pH of 1.1 by adding minute amount of
magnesia and then circulated inside the central compartment of an
electrolyser.
The electrolyser consisted of a filter press design model MP cell from
Electrocell
AB (Sweden) with three compartments separated by an anion-exchange
membrane (Excellion 1-100) and a cation exchange membrane (Excellion I-
200), both manufactured by SnowPure LLC. The geometric electrode and
membrane surface area was 100 cm2 and the spacing between each electrode
and the separator was 6 mm and also 6 mm between each membrane.
[0132] The cathodic compartment comprised a cathode plate made of a
titanium palladium alloy (ASTM grade 7; Ti-0.15Pd) supplied by Titanium
Industries. Prior to electrolysis the cathode was chemically etched by
immersing it
into a fluoro-nitric acid mixture (70 vol% conc. HNO3, 20 vol.% conc. HF and
10
vol.% H20) and then rinsing it thoroughly with deionised water until no trace
of acid
remained.
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39
[0133] The anodic compartment was equipped with a dimensionally stable
anode (DSATM) supplied by Magneto BV (Netherlands) made of a plate of a
titanium-palladium alloy substrate coated with a high loading of ruthenium
dioxide
(Ru02) acting as electrocatalyst for promoting the evolution of chlorine (Ti-
0.15Pd/Ru02).
[0134] The catholyte that circulated in loop within the cathodic compartment
was an aqueous solution of 350 g/L iron (II) chloride and 300 g/L magnesium
(II)
chloride adjusted at a pH of 1.1, while the anolyte that circulated in loop
within the
anodic compartment consisted of an aqueous solution of 20 wt.% hydrochloric
acid
with 17 wt.% magnesium chloride (MgCI2) and 10,000 ppm of ferric iron (Fe3+)
as
corrosion inhibitor the balance being deionised water.
[0135] The electrolysis was performed galvanostatically at a current density
of 500 A/m2. The operating temperature was 80 C and the volume flow rate of
both
catholyte, anolyte and initial solution was 1 L/min. At that current density,
the
measured overall cell voltage was 3.502 V. During electrolysis, ferrous
cations
from the iron-rich metal chloride solution crossed the Excellion 1-100 cation
exchange membrane, and pure iron metal deposited at the cathode. On the other
hand, chloride anions migrated through the permeable anion exchange membrane
towards the anodic compartment and discharged as chlorine gas at the surface
of
the anode.
[0136] After two hours of continuous electrolysis, the power was shut off
and the electrolyser was opened. The bright iron metal deposit plate was
easily
stripped from the titanium cathode by mechanical means. The measured thickness
was circa 0.126 mm and its mass was 10.04 g (See Figure 10). From these
experimental figures, the estimated faradaic current efficiency was 96.4% and
the
specific energy consumption at 500 A/m2 was 3.485 kWh per kg of iron. Chlorine
gas was recovered by means already described in Example 2a.
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[0137] Vanadium was also recovered by standard methods from the iron-
depleted solution exiting the central compartment as follows. A stoechiometric
amount of sodium chlorate (NaCIO3) was added to the iron-depleted solution to
oxidize all the vanadium cations (V4+, V5+) into pentavalent vanadium (V5+)
according to the reaction:
5VO2+ + CI03 + 2H20 -_+ 5VO2+ + 0.5CI2(g) + 4H+
[0138] Then the pH of the solution was raised by careful addition of a slurry
of slacked magnesia (Mg(OH)2) until the pH reached 2.0, but not above to avoid
the precipitation of black mixed ferroso-ferric hydroxides. At that pH, the
complete
precipitation of hydrated vanadium pentoxide (V205-250H20) occurred in the
form
of a gelatinous red brown precipitate. Since vanadium was the only species to
precipitate in this case, the precipitation would stop at a vanadium
concentration
below about 0.02 mol/L in the solution. Reconcentration of the solution
allowed to
recover more vanadium.
[0139] The red brown precipitate was separated by filtration using a similar
set-up to that described in Example 4. The red-brown gelatinous filter cake
was
removed from the filter paper and dried into an oven and later calcined inside
a
porcelain boat in dry air at 400 C in a box furnace (Fisher Isotemp) the water
vapour (H20), thereby yielding a red-orange powder of vanadium pentoxide. The
powder was then transferred into an Inconel crucible, melted at 700 C in air
and
cast onto a cool steel plate. The solidified black mass with a submetallic
luster was
then ground into a two disks vibratory cup mill with a hardmetal liner
(Fritsch
GmbH) using acetone as grinding aid and coolant. The product thus obtained was
technical grade vanadium pentoxide powder containing some chromium, iron and
manganese as major impurities.
[0140] Some results and characteristics of the electrolysis experiments
CA 02663652 2009-03-17
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41
conducted in Examples 2a, 2b, 5 and 6 are summarized in Table 5 below.
CA 02663652 2009-03-17
WO 2008/034212 PCT/CA2007/000026
42
~
0
~
OLC;~
c.S? E E M
c: 4= o V~ N l CO OM) O
t a C O(P (O (O
v) u) 0r- ao
C
Q T Y
co mNr M 'cr ct
a0
O p~ O ~ co
N U M Ce)
lL ~ C Ll LL C LL
o 0
C N L
'US U a) Cr)
c 0
~0 Q U .a 0 O~000 N_ 0 ~ O~ ~ 0
U O~ C~C~O (a~ V)QO)O~ d) N N a) N
+ l~ .. ~ i-.
(j U a) N N
Cc c: U- 7E U vU LL U LL U
O O O O lq: O O O
(a O~ 0 xLLCO Oa0 Co-n titt)
LL U G) 00 O~.I- CD tn a) C) a) O~
~ ~ ~ ~
0
> N (V M
U
Ln
O C =+ N
W s_TD cnE O
O
Q U U~ t~ ~ ~ N
F y C y C y'C y
,
O 0 C ~ ~ ~ O
~- C C C
rr~ Ea) C E moC Ec c E moc
O M M V o N t~
~
(O
2 G O (C ~~ O N ~ O O tC E G C
0) ~w ~ a~= o~~ o a~ E ~ 0 a) E~ o 0
-x
a) 0 =-x LL 3o-x LLso-m x ti 0
wa H~3a~E Ho3a~E E E
~
~ o
a~
E
o o co a$ oo
OX_-N O M ~ 00),
Q U ( ~ O -- O
O O O O U
O C
C X y ~ E
C
0 N 0 O
N~ p~ p O N^ tC t~) O
= 0 O C1 0 E`~
E QV~ O~ Q 0 Q Q~ O Q
` .
a) E coQCE -0 E> E ~cou wE
O. !a c O O c~ MO mO R C2cII
x xOZ.n> x~n K.o xc~ox
w w~0 oa) ~wco wca w>U)
m
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43
EXAMPLE 7
[0141] Removal of calcium from iron-depleted electrolyte. - After each
one of Examples 2a, 2b, 5 and 6, concentrated sulfuric acid was added to the
iron-
and possibly vanadium-depleted solution exiting the electrolyser for removing
calcium as insoluble calcium sulfate dihydrate (CaSO4.2H20) that precipitated.
The
precipitate was removed by filtration. The clear solution that contained only
magnesium and/or aluminium chlorides was ready for pyrohydrolysis.
EXAMPLE 8
[0142] Selection of the cathode material for conducting electrolysis in
Examples 2a, 2b, 5 and 6- The selection of cathode material was conducted with
an electrolyser and set-up identical to that used in Example 2a but with a
synthetic
catholyte circulating in loop and made of an aqueous solution of 350 g/L iron
(II)
chloride and 300 g/L magnesium (II) chloride adjusted at a pH of 1.1 while the
anolyte that circulated in loop consisted of an aqueous solution of 20 wt.%
hydrochloric acid with 17 wt.% magnesium chloride (MgCI2) and 10,000 ppm of
ferric iron (Fe3+) as corrosion inhibitor the balance being deionised water.
The
electrolysis was performed galvanostatically at 80 C during two hours. The
polarization curves, that is, the cell voltage vs. the current density were
recorded
for each cathode material. The materials tested were a titanium-palladium
alloy
ASTM grade 7(Ti-0.15Pd) from Titanium Industries, Zircadyne 702 from Wah
Chang, austenitic stainless steel AISI grade 316L, aluminum grade 6061 T6 and
pure copper. As expected, only titanium and zirconium allowed the easy
stripping
of the iron deposit. The polarization curves are presented in Figure11.
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44
EXAMPLE 9
[0143] Selection of the anion exchange membrane for conducting
electrolysis in examples 2a, 2b, 5 and 6 - The selection of the anion exchange
membrane was conducted with an electrolyser and set-up identical to that used
in
Example 2a. The synthetic catholyte circulating in loop in the cathodic
compartment was made of an aqueous solution of 350 g/L iron (II) chloride and
300 g/L magnesium (II) chloride adjusted at a pH of 1.1 while the anolyte that
circulated in loop in the anodic compartment consisted of an aqueous solution
of
20 wt.% hydrochloric acid with 17 wt.% magnesium chloride (MgC12) and 10,000
ppm of ferric iron (Fe3+) as corrosion inhibitor, the balance being deionised
water.
The electrolysis was performed galvanostatically at 80 C during two hours. The
polarization curves, that is, the cell voltage vs. the current density were
recorded
for each anion exchange membrane. The membranes tested were a Excellion I-
100 (SnowPure LLC), Neosepta AMH, ACM, and AHA (Tokuyama Co. Ltd. -
Eurodia), Selemion (Asahi Glass) and Ultrex AMI-7001 (Membrane
International). The polarization curves are presented in Figure 12.
EXAMPLE 10
[0144] Selection of the composition of anolyte for conducting
electrolysis in examples 2a, 2b, 5 and 6 - The selection of the anolyte was
conducted with an electrolyser and set-up identical to that used in Example 9
but
with a synthetic catholyte circulating in loop in the cathodic compartment,
which
was made of an aqueous solution of 350 g/L iron (II) chloride and 300 g/L
magnesium (II) chloride adjusted at a pH of 1.1 and an anolyte circulating in
loop in
the anodic compartment, the composition of which varied as follows: (i) 20
wt.%
MgCiz + 2wt.% HCI; (ii) 20 wt.% MgCI2 + 5 wt. to HCI; (iii) 17 wt.% MgCI2 + 20
wt.%
HCI; (iv) 20 wt.% HCI, all with 10,000 ppm wt. Fe(III) as a corrosion
inhibitor. The
electrolysis was performed galvanostatically at 80 C during two hours. The
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polarization curves, that is, the cell voltage vs. the current density were
recorded
for each anolyte composition. The polarization curves are presented in
Figure13.
[0145] Although the present invention has been described hereinabove by
way of specific embodiments thereof, it can be modified, without departing
from the
spirit and nature of the subject invention as defined in the appended claims.
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46
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