-------
» Northeast Industrial Waste Exchange (New York);
* Piedmont Waste Exchange (North Carolina);
» Southern Waste Information Exchange (Florida);
• Techrad (Oklahoma);
v Tennessee Waste Exchange (Tennessee);
» Virginia Waste Exchange (Virginia);
* Western Waste Exchange (Arizona); and
* World Association for Safe Transfer and Exchange (Connecticut).
The following is a list of the private material exchanges currently in
v - 7
business:
• Zero Waste Systems, Inc. (California);
* ICM Chemical Corporation (Florida);
* Environmental Clearinghouse Organization - ECHO (Illinois);
« American Chemical Exchange - ACE (Illinois);
• Peck Environmental Laboratory, Inc. (Maine);
• New England Materials Exchange (New Hampshire);
• Alkem, Inc. (New Jersey);
» Enkarn Research Corporation (New York);
• Ohio Resource Exchange - ORE (Ohio); and
» Union Carbide Corporation (in-hotise operation only, West Virginia).
i
5.3 EXAMPLES OF WASTE MINIMIZATION PRACTICES
There is a growing'incentive for companies to undertake waste
minimization programs as a consequence of increasing waste disposal costs and
liability. Besides protecting human health and the environment by drastically
lowering the amount of waste generated, waste minimization programs can, in
many cases, provide substantial economic benefits. The following is a summary
of waste minimization practices employed by various industry categories. Much
of this information was obtained tnrough the results of a 1986 study
commissioned by EPA to provide the U.S. Congress with information on the
status of current waste minimization efforts in the country. The
industrial waste generator categories discussed below parallel those
identified in Section 3.3 as being high volume metal/cyanide waste generators,
5-21
-------
5,3.1 Acrylonitrile Production
The most significant means of minimizing waste generation can be
accomplished by improving product yields thereby reducing the formation of
heavy metal impurities. This could be accomplished through improvements in
catalyst development and gas-catalyst contact in the ammoxidation reactor or
through staged addition of NH.» Segregation of acrylonitrile and
acetonitrile purification bottoms from the quench-absorption aqueous effluent
would permit incineration of the concentrated streams and reduce the toxieity
of the wastewater.
5.3,2 Metal Finishing
Waste minimization efforts in the metal finishing industry consist
primarily of methods to minimize consumption of rinse water, extend bath life,
recover baths and rinses, or to use gome form of raw material substitution.
Rinse water accounts for roughly 90 percent of raw waste generation in the
4
industry. Methods to achieve waste reduction which are unique to metal
surface finishing are discussed below whereas waste minimization efforts which
are similar to those employed in electroplating operations (e.g., reduction of
drag-out, use of counter—flow rinses, etc.) are discussed in the following
section.
The use of air-dried, no-rinse chromate conversion coatings for steel,
galvanized steel, and aluminum in the coil coating industry has been rep.orted
by the EPA. The literature also documents successful implementation of
chromic acid recovery through both ion exchange and evaporation, and nickel
recovery from rinses via electrodialysis, each involving subsequent recycling
to the bath and reuse of rinse water.
Raw material substitution can effectively reduce quantities of
contaminated rinses. Cyaniding can be replaced by gas phase carbonitriding
which eliminates the need for the rinse step. Chromic acid rinses
following zinc-based phosphating have been replaced by nonchrome rinses,
although some loss in effectiveness has been observed.
Generation of spent baths can be reduced by various methods aimed at
extending bath life or removing contaminants. Filtering bath solutions is
r\ f*
widely practiced to remove insoluble metallic salts. These otherwise
5-22
-------
precipitate onto heating/cooling process equipment, thereby reducing energy
efficiency, or precipitate onto metal parts resulting in impaired product
21
quality. Soluble salts can also lower bath activity and have been removed
from electroless nickel baths by crystallization with subsequent filtration.
The U.S. Bureau of Mines has experimented with chromic acid etctiant recovery
through use of an electrolytic diaphragm cell. Trivalent chromium is oxidized
22
and reused along with the simultaneous recovery of copper.
Contamination of baths can also be reduced by taking precautionary
measures, such as thorough rinsing, to reduce drag-in. Also, rack maintenance
(e.g., application of fluorocarbon coatings) is effective in preventing
contaminant build-up resulting from dissolution of rack metals.
Primary bath treatment methods resulting in recovery include electrolytic
recovery, ion exchange, crystallization, and evaporation. Evaporation has been
used successfully to recover plating solutions, chromic acid,
23
nitric/hydrofluoric acid pickling liquors, and metal cyanide baths. An
example of recovery of metal finishing wastes is recovery of electroless
nickel plating sodium phosphate salts by using ion exchange resins activated
en-
24
22
with hypophosphorous acid. Liquid membranes have been used by Bend
Research Inc. to recover contaminated dichrotnate rinses and baths.'
Raw material substitution has been applied to eliminate or reduce the
amount of hazardous waste generated by metal finishing processes. As stated
previously, cyaniding baths can be replaced by gas phase carbonitriding which
utilizes ammonia gas instead of cyanide to provide nascent nitrogen. However,
this is less economical for solutions which are used to treat many small
batches requiring different cycle times and high heating rates.
Polysiloxanes, substitutes for cyanide-based stress relievers in electroless
copper plating, are currently marketed by General Electric. Ferric •
chloride or ammonium persulfate solutions can be substituted -for
chromic—sulfuric etchants and strippers if it is compatible with the basis
metal. Peroxide-based secondary pickle solutions have successfully replaced
chromic acid pickling liquor at a wire manufacturer resulting in improved
product quality and economic savings due to recovery of the resulting pure
copper oxide sludge. J As another example, at least 5 companies currently
offer trivalent chrome systems for conversion coating applications which
currently use hexavalent chrome. • Other substitutions include electroless
copper for electroless nickel, plating of zinc instead of nickel, and varying
substitutes for cadmium and silver depending on the application.
5-23
-------
Improved operating practices can also contribute to bath life. More
frequent monitoring of bath activity and temperature can result in timely
correction of deviations thereby improving both product quality and bath
life.11
Several processing alternatives may provide potential substitutes for
waste generating metal plating operations. However, these methods are either
in the developmental stage or otherwise have not vet been widely applied in
the U.S. These include vacuum evaporation methods for coating nickel,
aluminum, and other metals; ion plating of chromium and cadmium; and chemical
vapor deposition. Similarly, ion bean processing may provide an alternative
to case-hardening treatments.
5.3.3, Electroplating
Plating bath life can be extended by taking methods to reduce plating
bath contamination. Examples of practices that will extend bath life include:
use of purer anodes; improved rinsing, rack design, and extended drip time to
reduce drag-in; use of deionized water to compensate for evaporative loss; and
use of treatment techniques to selectively remove contaminants. Examples of
treatment methods include filtering to remove suspended solids, use of carbon
adsorption or chemical oxidation to remove organic breakdown products, and
21
freezing to effect carbonate precipitation from cyanide solutions.
Another plating bath waste minimization option,is substitution of
hazardous plating bath materials with nonhazardous compounds. For example,
cyanide solutions have been effectively replaced by cyanide-free zinc
solutions and pyrophosphate copper plating solutions. However, more stringent
precleaning of the metal substrate is required to ensure high quality
plating. Cadmium-based plating baths can be replaced with zinc graphite
plating, titanium dioxide vapor deposition, and aluminum ion vapor
deposition.
Cadmium can also be replaced by zinc except in applications in alkaline
• 2
environments or when the plate must be exceptionally thin. Hexavalent
chrome can be.replaced by the less toxic trivalent chrome in certain
applications, thus eliminating tbe need for a separate chrome reduction
treatment process. In addition, significant sludge volume reductions
(70 percent) have been reported due to the elimination of excess sulfate ions
5-24
-------
that are introduced during reduction. Much of chromium plating is also
used strictly for decorative purposes and is thus being replaced by other
coating operations (e.g., painting of automobile bumpers).
Waste rinse water generation can be minimized through drag-out reduction,
by optimizing design and configuration of the rinse system, or through
recovery of contaminated water. These can be accomplished through a variety
of ways including;
• Use of counter-current, multiple rinses;
* use of drip tanks, drain boards ,or stagnant rinses with recovery
apparatus immediately following baths;
* use of rinse water for plating bath make-up;
* lowering bath surface tension (e.g., use of nonionic wetting agents)
and viscosity (high temperature, changes in chemical composition);
* lowering concentrations of toxic chemicals;
* reshaping work pieces and rack layouts to improve drainage;
* increasing drip time; and
» use of methods to increase rinse efficiency such as agitation of
immersion rinses or use of spray or fog rinses.
The EPA has reported that acceptable chromium plating can be achieved
with CrO_ concentrations as low as 25 to 50 g/1 versus traditional
27
concentration levels of 250 g/1. Since drag-out ia directly proportional
to concentration (more so if viscosity effects are included), rinse
contamination can therefore be reduced by up to a factor of ten. One author
estimated that the use of wetting agents to decrease surface tension can
27
reduce drag-out by as much as 50 percent.
Automatic process controls enable drip time to be maximized by changing
it to correspond to variations in process throughput or changes in work
pieces. Use of counter-flow rinse tanks arranged in series are capable of
27
achieving theoretical reductions in water requirements UD to 90 percent.
Spray or fog rinsing is widely used and most effectively applied on
rack-mounted, simple shaped parts with high area of exposed surface. Finally,
rinse waters are frequently reused in the same process through recovery
5-25
-------
Ce.g., electrolytic recovery from stagnant rinses) or reused in another
process operation when its contaminants will not adversely affect the
subsequent processing step (e.g., in nickel plating, the same rinse can be
used following alkali cleaning, acid dip and nickel plating tanks). Cost
evaluations of rinsing options have been presented in detail in the
27
literature.
Sludges generated from onsite wasteuater treatment can be reduced by
modifying treatment operations. Waste segregation, use of more effective
precipitation agents, and sludge dewatering are the major categories of waste
reduction options. Segregation of wastes containing highly completed
solutions permits specialized batch treatment and thus optimizes reagent
requirements and subsequent sludge generation. Segregation of streams
containing different metals can result in waste products that are more
amenable to recovery or reuse. For example, nickel hydroxide sludges have
reportedly been reused as plating bath make-up, as have chromium bath scrubber
wastes.
Primary recovery techniques employed by the industry include evaporation,
electrolytic recovery, reverse osmosis, ion exchange, and electrodialysis,
In addition, several plating bath suppliers (e.g., MacDermid, Harshaw, CP
Chemical) reprocess spent baths for their customers.
Evaporation is simple and reliable but also energy intensive and
nonselective; i.e., it concentrates impurities as well as metal components.
Thus, to be economically attractive, evaporation is often utilized in
conjunction with other reduction techniques such as counter-current rinsing,
to concentrate solutions, and deionization of rinse water to reduce build-up
of calcium and magnesium salts in the recovered concentrate. Multiple effect
and vapor recompression evaporators can recover 90 to 99 percent-of heavy
78
metals and are currently used most commonly on chromium baths.
Relative to evaporation, membrane technologies are, in general, more
selective and lower in operating costs but are also more complex processes,
Reverse osmosis has been most effectively applied in the recover? of nickel-
29
rinses and has also been used for cadmium recovery. Its use is restricted
to dilute, prefiltered solutions with moderate pH levels to ensure sufficient
membrane life. Ion exchange is more versatile in its application,
currently applied in nickel, chromium, cyanide, silver and other
5-26
-------
metal-containing rinse solutions. However, it requires a high level of
process control and maintenance. Electrolytic recovery is highly effective
for recovering wastes from concentrated rinses and has found more recent
application in direct recovery from process and treatment baths (e.g., cyanide
destruct tank). Electrodialysis hss also found application for rinse
recovery to remove silver, cadmium and other metals. As with reverse
osmosis, membrane stability and fouling potential restrict its application.
A national survey of elecCroplaters and metal finishing facilities
conducted in 1983 identifies evaporation, ion exchange, and reverse osmosis as
the three most widely applied recovery technologies. Use of these methods
for recovery of specific solutions is summarized in Table 5.3.1. Technologies
which have potential for plating solution recovery, based on pilot testing,
industrial application, or theoretical considerations are summarized in
Table 5.3.2. ' The EPA has estimated that the majority (48 percent) of
the electroplating industry heavy metal discharges are accounted for by nickel
and chromium. Since these are also the most expensive of the metals which are
discharged in large volumes, they account for an even higher fraction of the
32
the value of lost minerals in the industry (83 percent). Recovery costs
for these metals is expected to be significantly offset by the recovery value
32
of the metals.
5.3.4 Printed Circuit Boards
Waste reduction methods are similar to those previously discussed for
electroplaters and metal finishing facilities. Those that are specific to
printed circuit board manufacturers which result in the reduction of
California list wastes are discussed below. These include direct substitution
or recovery of these wastes as well as reduction of other metal containing
wastes which contribute to wastewater sludge generation (F006).
Chromic acid used for desmearing has been successfully replaced by
concentrated sulfuric acid and, more recently, potassium permanganate. The
advantages of the latter are: 1) it does not introduce chromium into the wast*
effluent; 2) it is not hygroscopic like sulfuric acid, therefore bath lives
are extended; and 3) recent developments by Morton Thiokol have resulted in
production of a proprietary additive which will reoxidize the permanganate
5-27
-------
TABLE 5.3.1. APPLICATION OF LEADING RECOVERY TECHNIQUES FOR
ELECTROPLATING AND METAL FINISHING
Application
Chromium plating
Nickel plating
Copper plating
Zinc plating
Cadmium plating
Silver/gold plating
Brass/bronze plating
Other cyanide plating
Mixed plating wastes
Chromic acid etching
Other
Units
Evaporation
158
63
19
7
68
13
10
6
-
6
16
in operation*
Ion Reverse
exchange osmosis
50
38 106
3
3
-
20
_
_
11 6
-
2 1
Source; Reference 31.
*Accorditig to a survey of the U.S. Electroplating and Metal Finishing
industries cited in the reference.
5-28
-------
TABLE 5.3.2. POTENTIAL METAL FINISHING BAT1 RECYCLING PROCESSES
Metal finishing
baths commonly used
Plating - Hard and Decorative
Nickel
Nickel Iron
Copper Cyanide
Copper Acid
Copper Pyrophosphate
Tin, Acid
Tin, Alkaline
Tin Fluoborate
Zinc Cyanide
Zinc j Acid
Tin/Lead, Fluoborate
Cadmium Cyanide
Gold Cyanide, Alkaline
Gold Cyanide, Acid
Silver Cyanide
Electro less Baths: Copper
Nickel
Pickling: Sulfate Copper
^2^2/^2^04 Copper
HN03 Copper
Cleaning: Alkaline Cleaners
Acid Cleaners
s
o
o
fyf
O
4J O
x ••-<
_l U
O <13
& U
« a
o a.
QJ V
F-l ^ >
ta w
X
X X
X
X X
X
X
X
X
X X
X
X X
X
X X
X
X
X X
X
X
X
X
X
w c
«-» o
96 -F)
F-l BJ
n; y
T) FH
O -Ft
u u-i
u n
0 W
4) 4-1
W 3
X
X
X
X
X
X
X
X
X
X
X
a
00
c
eg
,E
y
X
u
d
o
!""•
X
X
X
X
X
X
X
m
•H
W
o
E
o
o
to
aj
>
5rf
X
X
X
X
X
X
X
X
X
Source: Reference 5 and 31.
5-29
-------
thereby increasing bath life and reducing sludge generation in the etch tank.
However, this solution is relatively expensive and can spontaneously combust
if it is exposed to air and allowed to dry.
Plating baths are commonly replenished and treated to enable reuse. For
example, firms commonly remove organic breakdown products (e.g., from
stabilizers and brighteners) from copper, nickel and solder plating baths by
oxidation (e.g., potassium permanganate) followed by carbon adsorption and
21
filtration (e.g., diatomaceous earth).
A large number of printed circuit board manufacturers have switched from
panel plating to pattern plating. Since the latter only involves
electroplating board holes and circuitry, its use reduces the amount of
noncircuit copper which must be subsequently etched away. This, in turn,
reduces the amount of etching waste generated and discharged to onsite
treatment processes. Other processing techniques which can reduce or
eliminate the generation of etching wastes include using dry plasma etching
techniques (e.g., using reactive "gaseous radicals, using nonreactive ion
bombardment), using additive or semi-additive instead of subtractive board
manufacturing, using less toxic etcbants (e.g., ammonium persulfate,
peroxide-sulfuric acid which are widely used in place of chromic acid), or
using in-line recovery methods to extend etehant life (e.g., liquid membrane
copper recovery). Peroxide-sulfuric use in etchants, only recently
adopted by industry, has the advantage of not introducing additional chelators
into the plant's discharge stream. It is also easily regenerated through
crystallization which results in the precipitation of copper sulfate
crystals. These can be easily removed from the etch tank and have potential
reuse applications.
5.3.5 Inorganic Pigments Manufacture
Cadmium and other metal dusts collected in air pollution control
equipment have reportedly been recycled for use in low grade paint. (Versar
i960) Substitutes for red lead primer and chrome yellow (used in traffic
paint) have been identified but generally do not result in comparable
performance, cost, or color characteristics.
5-30
-------
Waste reduction at pigment production facilities has also been achieved
33
through modifications in conventional wascewater treatment systems.
Conventional treatment consists of ehroraate reduction followed by filtration
and landfilling of the collected solids. A modified process, employed by at
least two facilities, consists of the following: (l)use of improved filtration
systems to minimize wastewater participate content; (2)addition of soluble
barium salts to precipitate barium chromate which can then be used to produce
a light yellow pigment; and (3)pH adjustment to alkaline conditions to
precipitate lead and zinc (e.g., as hydroxides or carbonates) which can then
33
be recycled to the process as feedstock salts.
5.3.6 Petroleum_ Refining
Since the generation of hazardous waste from petroleum refining is a
direct result of the attempt to remove existing impurities from the crude
feed, waste minimization in che industry is primarily accomplished by sludge
consolidation. This includes maximization of slop oil recovery and separation
of water and oil from other, nonrecyclable waste products. For facilities
with cokers, much of the API, DAF, and slop oil sludges can be converted to
coke, according to industry representatives. In addition, the current
trend away from production of leaded gasoline will reduce wastes generated
from tetra-ethyl lead production and leaded gasoline storage (i.e., K052)
Leaded tank bottoms can also be reduced by agitation of tne tanks which
effectively transfers solids downstream, eventually ending up in .either
asphalt or coke byproducts. Methods have also been developed to recover this
sludge by dissolving it in a heated, low viscosity distillate with the
resulting liquid sent to slop oil recovery systems. Other processes
currently in use by refineries' include the, Victor extraction process, which
uses steam and air to separate residual oil trapped in the sludge; physical
sludge consolidation processes such as vacuum filtration; thermal, chemical,
or ultrasonic emulsion breaking; solvent extraction (e.g., B.E.S.T. process
from Resource Conservation Co.); and electroacoustic dewatering, Leaded
tank bottoms have reportedly been treated by calcination to recover lead
oxide. Tanic bottoms are reacted at. high temperatures to drive off water and
other volatiles, incinerating residual organics, and oxidizing the lead. '
5-31
-------
Many of the above processes are applicable for recovery of oil from API
separator sludge and DAF float. Other methods include installation of
floating roofs which was found to reduce the oxidation of oil and the
resulting formation of heavy waste material in API separators. Conversion
from induced air to pressurized air in DAP units has resulted in the
generation of less than one half the float volume for the same degree of
solids removal.
Methods to achieve waste minimization for other refinery waste streams of
concern include l)hydrotreating catalytic cracking feed to remove metal
contaminants, thereby extending catalyst life; 2)substitution of chromium
corrosion inhibitors in cooling water with organic chelating agents,
nonoxidizing biocides, and other proprietary compounds; and 3)recovery of
Raney nickel catalysts through roasting, leaching of- aluminate, and
35
preparation of nickel carbonate.
5.3.7 Hood Preserving
For chromium/arsenic preservatives received in drums or bags, closed
systems are available which can minimize residual levels of metal contaminants
in the containers. Alternatively, plastic liners or reusable drums can be
used. Sludge from the work tank can be minimized by careful operating
practices that ensure minimal amounts of dirt, silt, and loose wood fiber
entering the retort before treatment. Most facilities have installed drip
pads and spill basins to collect excess preservative which drips from the wood
pieces after treatment. Other measures to reduce the amount of water
contaminated and thus requiring further treatment includes covering processing
areas, increasing drip time, and diverting run-on. Use of nonchromate cooling
water treatment chemicals would reduce the amount of this compound in the
plant's combined treatment sludge.
5.3.8 Chloralkali Industry
The membrane cell process for the production of sodium hydroxide has
begun to replace the more costly mercury cell process. Since its introduction
in 1980, six plants have opened in the U.S. and, in response, mercury cell
plants have been closing. The membrane cell process eliminates the
5-32
-------
generation of mercury containing hazardous waste. In face, DuPont claims that
its most recently built plant will completely eliminate the production of
hazardous waste. Over the long term, conversion or closing of
non-competitive mercury cell facilities is likely since the membrane process
not only results in less pollution control costs but also requires
4
approximately 30 percent less energy per unit of production.
For plants which continue to use the mercury cell process, waste
minimization options are available. Retorting has been used in the
chluralkali industry to remove mercury from mercury-bearing sludges and solid
wastes. The waste is heated in an oxidizing environment forming mercury gas
which is collected by condensation. Alternatively, wastes are procreated
through hydrometallurgical processes. One facility leaches contaminated muds
with sulfuric acid to concentrate mercury and convert the bulk of the solids
to nonhazardous gypsum. Although the latter is capable of recovering over
99 percent of the mercury contained in the sludge, its capital cost is several
million dollars and thus-appropriate only to large volume facilities.
Solvent extraction has also been suggested for stripping mercury from effluent
wastewaters (see Section 7.2).
5.3.9 Other Industries
Raw material substitution and waste recovery technologies have been the
predominant means cf waste minimization in industries which use
29
silver-containing photographic films. Several companies, including Napp
29
Systems, are marketing silver free films for lithography. To a large
extent, used and spoiled film is already sent to professional recyclers for
silver recovery. Wastewaters containing silver are economically recovered
using technologies such as metallic replacement, chemical precipitation,
electrolytic recovery, reverse osmosis, and ion exchange.
Printing inks may contain heavy metals such as chromium. Contaminated
solutions can be recycled onsite or shipped to ink manufacturers who reuse
29
these materials in the formation of black newspaper ink.
Several examples of catalyst recovery by the inorganic chemicals industry
nave been documented. Solvent extraction has been used to recover vanadium
pentoxide from spent sulfuric acid catalysts using a high molecular weight
amine. Tfte amine solvent is subsequently evaporated leaving a reasonably pure
5-33
-------
ammonium vanadate which is available for reuse. Another example is
fluidization and precipitation of spent nickel catalysts used by inorganic
chemical manufacturers, A nickel salt is formed by dissolving the catalyst in
a mineral acid. This is reacted with soda ash to precipitate nickel
carbonate, which is then collected and reacted, with sulfuric acid to form 8
nickel sulfate solution. Sodium sulfide is added to precipitate iron salts
and the resulting solution is purified through filtration and evaporation. A
similar process is employed by manufacturers of plating chemicals for recovery
of nickel plating solutions.
Cadmium is used as a stabilizer for polyvinyl chloride. In this
application, it can be replaced by organotin compounds which are more
2
efficient but also more expensive.
5.4 WASTE MINIMIZATION SUMMARY
Regulatory trends appear to be moving towards the promotion of waste
minimization. The EPA has recently proposed requirements that generators
certify institution o£ hazardous waste reduction programs. Generators
would be required to reduce the volume or toxicity of hazardous wastes to a
degree determined by the generator to be economically practicable. Three
states currently have established source reduction/pollution prevention
programs: North Carolina, Minnesota, and Massachusetts. In addition,
Tennessee has established a "pilot program", and Kentucky, California,-
Maryland, and Washington have programs currently in development. These
programs vary but, in general, include information exchange, technical
assistance, and economic incentives to companies to encourage development of
their programs.
Table 5.4.1 presents a summary of several documented cases of waste .
9 3?
reduction involving metal/cyanide hazardous wastes. * Additional case
studies can be found in the appropriate sections of this document pertaining
to specific recovery technologies. Although some of the data in fable 5.4.1
are incomplete, this compilation clearly demonstrated the potential economic
benefits which can be achieved through implementation of waste minimization
technologies. In particular, since disposal costs have increased sharply in
recent years, payback periods indicated in the table can be interpreted as
being conservative estimates. _ ,,
-------
Reproduced from
best available copy.
Company and location
Climax Molybdenum
Colorado
Charlotte, N.C.
Uoupaca Foundry ,
Uautpnca, W lac on a in
Stlnidyne , Inc . ,
Sanford. N.C.
Di v in ion, Ine . ,
ElkhMrt. Indiana
Pioneer Metal
Finiahing. Inc..
FrankllnviLle, N.J.
Declie and Co. ,
Mo line, Illlnoia
Eneraon Flectrlc
Co., Special
Product! Divli ion.
Murphy, N.C.
CTE Sylvania.
Chicago, Illlnoia
Data General
Corporat Ion,
Clayton, N.C.
Ian Heath Co.
Birmingham, England
Hiluaukee. Ul
TABLE 5-4.
SIC
Code Product
1061 Raw mnlybJenun
copper, c inc.
iron, manganeae
3312 Steel
3321 Crey and compacted
3432 Plumbing product*
3471 ripe fitting
(abrlcet ion
3471 Electroplating
job ahop
3330 equlpnunt
35 Metal Clnlahlng
36 Stationery aunu-
lacture
3661 Electronic telephone
•witching equipment
3679 board*
Silver plating ol
gUtvara
on p limbing acceaiorlea
1. COMPILATION OF INDUSTRIAL
Haite, mlnlnlcation netliod deacriptlon
,n.t.,,.t,on .f Inte-ep.., can.,. «. P... ™...
preceding.
land filling.
cyanide concentrttlona Iroai plating oparationa.
wictt electrolytic recovery ol! copper.
•elected waataa.
incentive program for coat reduction or product ideaa.
Inatallation of cloaed-loop treatment ajaten and
electrolytic copper recovery eyatea..
and pyrophoephate copper rinaea . -
Cold recovery through electrodialya it and ion exch
WASTE MINIMIZATION
Percent Quantity (•)
93. (Cu)
99. (Hr,)
93. (Zn)
90. (Mb)
96.4 (Cn)
1980
46.0 1982
182.000 19>9
e«i/7.
50.0 40.000
(aludge) |ll/rr
(.lodge)
130.000 1980
100. o (a ib/d?
•olvent (paint
wa.te oil lolldt)
90 Ib/J,
(pi. tin,
•eld. oil.
C.U.Clc)
120 Ib/m
(.olvent)
15 |«l/«k .1976
(CuOH
.ludge)
100.0 400 lon/,r 1981
(proceaa (landfill
wa.te- wa.te)
14 kg HI. and
5 kg Cu per >eek
Capital Annual COft
(ll.OOO) (11,000) period
No change 129.6 Inaedilte
20.99} 1.5 Til
60 120 0.5 rr
210 52.460 ) Tr.
1.900 I5S.750 2.} rr.
674 1,800 1.1-5 jr.
6
30 lao 1.5 .onth
43 -
24.4 9 •onth.
(continued)
-------
TABLE 5.4.1 (continued)
Conpimy and location
Digital Equipment
Coipa , Tempe , AZ
Undine HaniiTac tur-
inf), Trenton, HO
Carol Ina Power and
Light Co., Hc«
Hill, H.C.
DuVe Paver Co.
Inc.. Hfltlheva,
N.C.
E.I. duPoi.t de
Cixon Chemical
America*. Uauaton,
01 In Corpora t Ion,
Standford, CN
California Clec-
troplat Ing, Loe
Angelea, Calif.
Vaaland Hetal Service*,
Inc., Cincinnati. OH
Central Motor*,
Pont lac. HI
Gillette n*ior
Boaton. HA
SIC
Code
3679
3714
4911
4911
739J
28
26
2221
7298
3079
26
28
34
3471
Product
Hetal radlatora
Electric power
E| eet r le power
Hum portrait
photography
Flbera
Polymer product*
Agricultural and
Rlonerllcal products
Coal
Petroleum product
01 ef In a , •roan tic a,
polyole I Ina , da—
•loner*, aolventa.
apeclaltlea, oil/fuel
additive*
Chemlcala
•trip and mill
product*
Plating
Multipurpose plant
Plating of burapera
Barrel Plating of HI
Waato Reduction Capital Annual coat
Haute • In in lent .on method Jeacrlptlon Percent Quantity (*> (*l,fM>0) (J 1,000) period
Ion eichange and eleccrolytic equlpoienc for recovery 100.0 . 27 22 |& atontha
of copper.
reviled equipment operation eatablftahmcnt o( mn ongoing (trltltna ton/yr
all*«r
2.919
Praeeaa change In ADU nanuf acttire. Marketing o( 50.0 1980
Proevca Modification to reduce load to treatment plant. 20.0
A tun (nun hjdroilde rennval Iron ..lodge for reelaautlon. 60.0 1982
buck Into tanha.
Inatallatlon of an Electrochemical Reactor for Cd recovery. 99.4 I960 6.) 17.5
(annual)
Inatallatlon of Reverea Oanoala on Ml dragout tenka. 8) 1979 B. 1
(continued)
-------
TABLE 5.A.I (continued)
Company and location
Baltimore, HD
General Plating,
Detroit , Michigan
Ford Motor Co. .
Sabine, Michigan
ford Ho tor Co..
Sabine. Michigan
Advance Plating
Co.. Cleveland
Ohio
Reliable Plating
Wiaconain
Colorcraft,
Rock ford, Illinois
Deluxe Hoc ion
Picture l.nbora-
. toriet, Hollywood,
| California
t^j PCA Inteinational
Hatthewa, N.C.
Phelpa Dodge
Corp. , Hidalgo,
New Heiico
Kennecott ,
Carfield, Utah
Ubia, Inc.,
Crankton. R.I.
Phillip* Plating Co.
Phillipa. Ul
Crawtordaville, IN
Allied Finishing.
*i:ilter; b:be(or*.
SIC
Code Product
3471 Plating of part*
3471 Metal plating
3471 Metal plating
3471 Automotive parta
plating
3471 Automotive part*
plating
3471 Plating of napkine,
paper towel* and
toilet tiuue
diapeneera
7395 rhoto finlahing
7019 film proceaalng
7393 Fhoto finiahing
33 Copper
33 Coppor
3471 Electroplating
job ahop
3471 Cr Plating
•teel product*
•etal • tapping a
9 and ]7.
Installation of filing film evaporator unit,
InotaLlation of evaporator recovery unit*.
Attachment of In nova Chrove Mapper ion tranafer ayateai
Attachnenc at lnno«a Chrowe Napp*r ion tranafer ayateai
to automatic hoiat lint.
Inatallation of production prototype developer
Recycle bleache*.
Recycle prt-bith, final bath, paper color developer.
facilitate iulfur recovery and energy aaving.
lUe of Noranda eont Inutma proceaa for copper iMlting
to facilitate aulfur and energy living.
Replacenent of counter-current flow rinaing by the
for recovery of plating aolution.
Inatallation of elo*«d-loop rlaing filai evaporator Co
concentrate Cr pitting bath d tag-out for recycle.
Percent Quantity (*) (ft 1,000) (ft 1. 000) period
31 Cd 1.22 pp. 1979 17 1.3 yr*
in effluent
3)0 Ib/dy
chrouic acid 100
60. 0-90.0
(Cr)
99.0
(water)
BO. 0-90.0 39.4 hg/wa. 1980 13 - Aawrtitation
(Cr) (H2CrO4) by Cr and H;0
99.0 92 1.1/dy pavinga alon*
(water)
80. 0 30
«2.0
(viter)
90.0
(Ag)
100.0
(Ag)
100.0
(water)
94.0
(vaate-
watcr)
1.8 kg Cr/hr .1979 -
180 kg/day i960 - 100 baaed on
Cr conaunption
• lone
841 Cr 1979 - - 3-2.3 yra.
conaunption,
13-201 •ludg*
generation
-------
A survey of 610 hazardous waste generators in Massachusetts was conducted
in 1985 to identify current and planned source reduction efforts. Of these
facilities, 238 were identified which practiced source reduction activities
for metal containing wastes, 10? for cyanides, and 71 for petroleum refining
wastes. Current source reduction activities and percent reduction in volume
achieved are summarized in Table 5.4.2. Predominant! methods employed include
waste segregation, process modification, improved housekeeping, precipitation,
improved rinsing, and chemical detoxification for cyanides.
This section was not intended to represent a complete survey of waste
reduction practices available to generators of metal/cyanide wastes. The
limited scope of this survey only permitted a broad overview of available
methods to be presented, supplemented with specific examples for high volume
waste sources, A comprehensive literature survey of waste reduction practices
•3Q
is being undertaken by the EPA Office of Solid Waste. The survey data
will be compiled in the form of a computerized data base and is intended to
provide technical assistance for both states and private companies. This
should be available by 1988 or early 1989 and is expected to represent a
significant improvement over current compilations of waste minimization
39
data. Other useful sources of information which are currently available
include several surveys which provide lists of articles, by industry, on waste
10 40 41
reduction practices, ' '
5-38
-------
TABLE 5.4.2.
SOURCE REDUCTION ACTIVITIES PRACTICED BY RCRA
WASTE GENERATORS IN MASSACHUSETTS
Source reduction
technique
Waste aggregation
Procea* modification
Better housekeeping
Waste recycling
Raw material institution
Waste rime
Neutralization
Filtration
Distillation
Praduet- reformulation
Precipitation
Improved rxttsing
Chemical detoxification
Sedimentitioa
Clarification
Evaporat ion
Carbon adsorption
Xon exchange
Flotation
Slectrodialysia
Other
Total source reduction
Aqueous
Facilities
using
method*
a)
9.T
11.3
8.8
5.9
1.3
5.0
5.9
6.3
0.4
2.1
9.7
9.7
2.1
4.6
5.9
2.1
0.4
3.8
1.3
1.7
2.1
238=
netals
Waste
reduction^
(Z)
46
39
42
62
22
43
65
61
25
64
47
31
60
53
93
-71
10
47
90
56
22
51
Cyan:
FacilitieB
using
method*
(J)
12.1
13.1
9.3
1.9
5.6 '
1.9
3.7
3.7
-
0.9
7.5
15.0
12.1
2.8
3.7
1,9
-
1.9
-
1.9
0.9
107'
ide.
Waste
reduction0
(Z)
5D
50
30
92
51
55
54
50
-
10
56
33
61
22
74
82
-
57
-
19
15
48
Petroleum wastes
Facilities
using Waste
method8 reduction*1
(I) (Z)
28.6 19
14.3 5
14.3 5
14.3 100
-
14.3 . 100
_
_
14.3 100
-
-
-
-
_
-
-
_
-
-
_
~" —
7e 50
BFercentsge of facilities practicing seas form of source reduction.
"Percentage of waste generation prior to implementation of source reduction technique.
'Total cumber of facilities practicing so-jrce reduction.
Source: Adapted free Reference No. 6,
5-39
-------
SECTION 5.0
REFERENCES
L. Minnesota Waste Management Board, Hazardous Waste Management Report.
1983.
2. Congress of the United States. Office of Technology Assessment. Serious
Reduction of Hazardous Haste: For Pollution Prevention and Industrial
Efficiency. U.S. Government Printing Office, Washington, B.C. 1986.
3. Carner, P. Telseon with M. Kravett, Alliance Technologies Corporation.
W.R. Grace and Company. Lexington, HA. June 1986.
4. Versar, Inc. Technical Assessment of Treatment Alternatives for Hastes
Containing Metals and/or Cyanides. Springfield, VA. Performed for U.S.
EPA Office of Solid Waste under Contract No. 68-03-3149. October, 1984.
5, Breton, M. et al. Alliance Technologies Corporation. Technical Resource
Document: Treatment Alternatives for Solvent Containing Wastes. Prepared
for U.S. EPA HWERL, Cincinnati, OH under Contract Ho. 68-03-3243.
August, 1986.
6. Roeck, D.R., et al. GCA Technology Division, Inc. Hazardous Waste
Generation and Source Reduction in Massachusetts. Bedford, MA. Contract
No. 84-198, MA Dept. of Env. Mgt., Bureau of Solid Waste Disposal,
June, 1985 (Draft).
7. Wilk, L. et al. Alliance Technologies Corporation. Technical Resource
Document: Treatment Alternatives for Corrosive Wastes. Prepared for U.S.
EPA HWERL, Cincinnati, OH under Contract No. 68-03-3243. Sept., 1986.
8. Versar, Inc. National Profiles for Recycling: A -Preliminary Assessment.
(Draft) Prepared for U.S. EPA Waste Treatment. Branch under Contract
No. 68-01-7053 July, 1985.
5-40
-------
9. Versar, Inc. and Jacobs Engineering Corporation. Waste Minimization
Issues and Options. Volume 3. U.S.EPA Office of Solid Waste.
EPA 53Q/SW-8&-043. October, 1986.
10. Overcash, Michael R. Techniques for Industrial Pollution Prevention. A
Compendium far Hazardous and Non-Hazardous Waste Minimization. Lewis
Publishers. 1983.
11. Versar, Inc. and Jacobs Engineering Corporation. Waste Minimization
Issues and Options. Volume 2. U.S.EPA Office of Solid Waste.
EPA 530/SW-86-042. October, 1986.
12. Methods Described to Minimize and Manage Electroplating Sludges. The
Hazardous Waste. Consultant. July/August, 1985.
13. Steward, F. A. and W. J. McLay. Waste Minimization: Part 1:
Introduction, Alternate Recovery Technologies. Metal Finishing, August,
1985, pg. 23.
14. Steward, F, A. and W. J. McLay. Waste Minimization; Part 2: Concentrate
Return Methods. Metal Finishing, September, 1985, pg. 55.
15. Steward, F. A. and W» J. McLay. Waste Minimization; Part 3: Non-Return
Recovery Methods. Metal Finishing, October, 1985, pg 63.
16. Steward, F. A. and W. J. McLay. Waste Minimization: Part 4: Recovery and
Regeneration of Process Baths. Metal Finishing, November, 1985, pg 69.
17. Steward, F. A. and W. J. McLay. Waste Minimization: Part 5: Recycle of
Treated Wastewater. Metal Finishing, December, 1985, pg 47.
IS. "ersar, Ir.c. ar.d Jacobs Engineering Corporation, *~aste Minimization
Issues and Options. Volume 1. U.S.EPA Office of Solid Waste.
EPA 530/SW-86-041. October, 1986.
5-41
-------
19. Tucker, S.P., and G.A. Carson, NIOSH. Deactivation of Hazardous
Chemical Wastes, Cincinnati, OH. Environmental Sci. Tcchnol.,
9(3):215-220. 1985. ,
20. U.S. EPA. Development Document for Effluent Limitations, Guidelines and
Standards for the Metal Finishing Point Source Category. Effluent
Guidelines Division, Washington, D.C. EPA 440/l-82-Q91b. August, 1982.
21. Nurmo, Thomas et al. Alliance Technologies Inc. Waste Minimization in
the Printed Circuit Board Industry - Case Studies. Prepared for U.S. EPA
HWERL under Contract No. 68-03-3243. 1986.
22. Basta, Nicholas et al. Total Metals Recycle is Metal Finisher's Goal.
Chemical Engineering, August 8, 1983. pg 16.
23. Cushnie, G. C. Centec Corporation, Reston, VA. Navy Electroplating
Pollution Control Technology Assessment Manual. Prepared for the Naval
Civil Engineering Laboratory, Port Hueneme, CA. NCEL-CR-84.019.
February, 1984.
24. Martin, M., Bend Research Corporation. Bend, OR. Telephone conversation
with Lisa Milk, Alliance Technologies Corp., Sept. 24, 1986.
25. Steward, F. A., ERC/LABCY Company. Process Changes to Reduce the
Production of Industrial Sludges. Presented at1 the 1981 Annual Mtg. of
the AIChE, New Orleans, LA. Nov. 8-12, 1981.
26. Kirk Othmer Encyclopedia of Chemical Technology. John Wiley and Sons,
New York, NY. Third Edition. 1978.
27. U.S. EPA. Control and Treatment Technology for the Metal Finishing
Industry: Ir.-Plsnt Changes (Sussaary Report). U.S. EFA. IERL Cincinnati,
OH. EPA 625/8-82-008. January, 1982.
5-42
-------
28. Clark, R. Massachusetts Bureau of Solid Waste Disposal. Massachusetts
Hazardous Waste Source Reduction Conference Proceedings, MA Dept. of
Environmental Management. October, 1983.
29. Campbell, M.E. and W.M. Glenn, Pollution Probe Foundation. Profit From
Pollution Prevention: A Guide to Industrial Waste Reduction and
Recycling. Toronto, Canada. 1982.
30. Kohl, J. and B. Triplett. Managing and Minimizing Hazardous Waste Metal
Sludges: North Carolina Case Studies, Services, and Regulations. North
Carolina University, Raleigh Industrial Extension Service. 1984.
31. Noll, K. E., et al. Illinois Institute of Technology. Recovery, Reuse,
and Recycle of Industrial Waste. U.S. EPA Office of Research and
Development. EPA 600/2-83-1. November, 1983.
32. Grosse, Douglas W. Treatment Technologies for Hazardous Wastes: Part 4 -
Metal Bearing Hazardous Waste Streams. JAPCA Vol 36, No. 5. May, 1986.
33. Risstnann, E.L., et al. Waste Minimization Audits: How they can Lead to
Reductions ,in Hazardous Waste Generation. DuPont Waste Minimization
Symposium. Wilmington, Del. October 16, 1986.
34. Stoddard, S. K. Alternatives to the Land Disposal of Hazardous Wastes:
An Assessment for California. Toxic Waste Assessment Group, Governor's
Office of Appropriate Technology, State of California. 1981.
35. Tsuen-Ni, Lung, et al. Chemical Reclaiming of Nickel Sulfate from
Nickel-Bearing Wastes. Conservation and Recycling. Vol. 6, No. 1/2. 1983.
36. U.S. Federal Register. 51 FR/10177. March 24, 1986.
37. Laughiin, R.G.W. et al. Ontario Research Foundation. Metal Finishing
Industry Technical Manual: Waste Abatement, Reuse, Recycle, and Reduction
Opportunities in Industry. Prepared for Environment Canada. Jan., 1984.
5-43
-------
38. RC8A Performance Management Standards for Waste Minimization Said Not
Needed Now. Environment Reporter. Current Developments. Nov. 7, 1986.
39. Elaine Eby, U.S. EPA OSW, Telephone conversation with T. Nunno, Alliance
Technologies Corporation, March 9, 1987.
60. Hunt, Gary and Roger Schecter. Pollution Prevention Pays: Bibliography by
Industrial Category. North Carolina Department of Natural Resources and
Community Development. January, 1986.
41. Mathews, J.E. Industrial Reuse and Recycle of WastewaterSj Literature
Review. U.S. EPA RSKERL, Ada, OK. EPA 600/2-80-183. December, 1980.
5-44
-------
SECTION 6.0
MEMBRANE SEPARATION TECHNOLOGIES FOI METAL REMOVAL
6.1 PROCESS DESCRIPTION
The membrane processes considered here include conmercially proven
technologies such as ultrafiltration, reverse osmosis, and electrodialysis.
Some discussion is also provided for other membrane technologies such as
Donnan dialysis and coupled transport whose applicability for the treatment of
hazardous waste streams has not yet been commercially demonstrated. Reverse
osmosis and electrodialysis are used to recover plating compounds from
rinsewater and to permit possible reuse of rinse waters as plating bath
make-up. Ultrafiltration alone is of little value for these applications, but
is used in combination with chemical treatment to physically contain metal
sludges. It is also used as a. pretreatment for other processes, such as
reverse osmosis, which are subject to fouling and plugging due to the presence
of particulates or high dissolved solids levels of certain salts in the feed.
Although the waste treatment/recycling applications are not extensive, these
processes have found growing acceptance in applications such as desalination
of seawater and brackish waters, and as unit operations in the food and
pharmaceutical industries.
According to Cheryan , the world-wide market for membranes, less than
S10 million in I960, reached S400 to S600 million/year in 1986. More "than 30
manufacturers of membranes are identified by Cheryan in his ultrafiltration
2-5
handbook. This handbook and several other books on membrane technology
are recommended for those concerned with the theory, development, and
applications of membrane technology to individual process and waste streams.
!
The primary function of a membrane is to allow preferential containment
and transport of certain components present within waste streams. Membranes
can be classified in a number of ways in accordance with factors such as their
origin, chemical composition, structure (e.g., pore size and asymmetry of pore
6-1
-------
structure), and mechanism .of membrsne action; e.g., adsorptive vs. diffusive,
ion exchange, osmotic, or nonselective (inert) membrane. Figure 6.1.1,
taken from Cheryan, providei a. classification of various separation processes
based on particle or molecular size and the primary factor affecting the
separation process. As shown in the Figure, membrane processes such as
ultrafiltration, reverse osmosis, and eiectrodiaiysis permit separation-of
dissolved molecules down to the ionic range in size, provided the appropriate
membrane is used.
The distinction between netnbrane processes such as ultrafiltration and
reverse osmosis is somewhat arbitrary and has evolved with usage and
convention. Table 6.1.1 shows some characteristics of several membrane
processes, including osmosis and dialysis, two processes with no apparent
utility for hazardous waste treatment. They have been included for reference
and completeness, along with microfiItrat ion, a process similar to
ultrafiltration that is often used as a pretreatment to remove suspended
solids that may interfere with the operation of molecular separation processes.
Another useful classification system is the normal operating
concentration range of membrane and other technologies. Figure 6.1.2 provides
this information for a number of processes, including reverse osmosis and
eiectrodialysis. These operating concentration ranges are based on a number
of factors which include the increase of osmostic pressure with increasing
concentration to levels that exceed membrane capacities, selectivity of the
membrane process, flux (defined for processes such as reverse osmosis as the
volume flow rate per unit area and pressure), and cost.
The flux obtainable with reverse osmosis, ultrafiltration, and other more
conventional filtration media is shown in Figure 6.1,3. As shown in the
Figure, ultrafiltration flow rates ran^e from roughly 0.1 to
10 gal/ft /day/psi. Ultrafiltration systems are typically operated at
pressures ranging from 10 to 100 psig, resulting in; flow rates that are still
several orders of magnitude below conventional filtration processes, but with
size retention of the order of 10 to 200 angstroms (0,001 to Q.02ym), as
opposed to one or more microns for conventional filters. These size values
are arbitrary and it is customary to refer to a "molecular weight cutoff" when
attempting to classify ultrafiltration membranes. Ultrafiitration is
generally considered suitable for separation of molecules ranging from about
1,000 to 1,000,000 in molecular weight. Thus ultrafiltration-, as designated
6-2
-------
a-
Oi
Priroory Factor
AHecting Separation
Size
Charge
Vapor Temp, Pressure
Solubility
Surface Activity
Densl t y
Angstroms
Microns 1C
| Mlcroflltert | | Cloth and Fiber Fillers
Ultrallltralion | Screens, Etc.
Reverse Osmosis |
Dialysis |
Eleclrodlalysls
Ion Exchange
Distillation/Freeze Concentration
Solvent Extraction
1 Foam and Bubble Frocl ionollon 1
1 Ultrocentrlluges
| Cent rl 1 uges
Liquid Cyclones 1
1 Gravity Sedimentation 1
1 III.
IO I02 I03 I04 I05 I06 I07
)~4 IO"3 IO"E 10"' 1 10 IO2 IO3
Ionic Macromolecular Micron | Fine Coorse
Range " " Range " * H«ticli -H Particle —Particle
Range | Range Range
Figure 6.1.1. Useful range of separation processes, showing the range of particle
or molecular size covered by each process and the primary factor
governing each separation process.
Source: Reference 1.
-------
TABLE 6.1.1. MEMBRANE SEPARATION PROCESSES
Process
Principal
driving force
Function of membrane
Permeate
Retentate
ULtrafiltration Pressure
Reverse osmosis Pressure
Electrodialysis
Donnan dialysis
Discriminates on the
basis of molecular size,
shape, and flexibility
Selective transport of
water from concentrated
solution
Water and Large molecules
small molecules
Electromotive Selective ion transport
Force
Concentration Ion transport with
charge equalization
Coupled transport Concentration Ion transport through
complexing agent in
membrane
Water
Water and
ionic solutes
Metal ions
Specific
metal ions
Solvent
Nonionic solutes
Hydrogen ions
from transfer
fluid
Other solute
ions
Dialysis
Osmosis
Concentration Selective solute
transport
Water and Large molecules
small molecules
Chemical
potential
Microfiltration Pressure
Selective transport of
water into more con-.
centrated solution
Removal of particulates
Water
Solutes
Water and dis- Suspended
solved solutes particles
Source: References 1 and 6.
-------
CHEMICAL PRECIPITATION
71
ISO
1
100
REVERSE OSMOSIS
400
3,300
ELECTROOIALYStS 1
300
ION EXCHANGE
i
eo
£9.000
,000
DISTILLATION
1
100 eoo 400 eoo 1,000 2,000 4,000 6,000 10,000 20,000 40,000 eo.ooo 100,000
Milligrams per liter (ports per million)ol Totol Dissolved Solids
Figure 6.1.2. Normal operating concentration range of separation technologies.
Source: References 7 and 8.
-------
o-
I
10'
Humon Hair
Red Blood Cells—L°_
Tolcum Powder
Small Bacteria
Influenza
Virus
- ...
°
-I
o
a
Starch Molecule
Egg Albumin
=\-
"
o
(C
•5 10
U
5
"o
Glucose Molecule
Chloride Ion
Reverse
Osmosis
Membranes^
MicropOfOUB
Fillers
Ullroflltrallon
Membranes
1
Conventional
Particle
Filters
FLUX, gol/sq It-Doy-Apsl
I03
Figure 6.1.3. Pore size vs. flow rate for separation media.
Source: Reference 7, page 90.
-------
by this retention characteristic, is not sufficient by itself to remove
dissolved ionic species such as metal ions (and cyanides). Selective ion
retention by ultrafiltration membranes, such as that which occurs in reverse
osmosis, may take place, but the effect is slight and not significant in terms
of effective separation. However, both ultrafiltration and microfiltration
are capable of effectively collecting colloidal metal suspensions following
precipitation of dissolved metal ions; and they also find application as a
pretreatment for reverse osmosis and other processes to protect membranes from
clogging.
Reverse osmosis, sometimes called hyperfiltration, may be used to
concentrate dilute solutions of many inorganic species, including dissolved
metal ions and many organic solvents* Reverse osmosis systems are available
from many manufacturers for the treatment of metal-bearing waste streams.
Ideally, they permit only the transfer of water, selectively retaining all
other dissolved species within the waste stream. Operating pressures used for
reverse osmosis are high; of the order of 300 to 1,500 psig, in order to
overcome the osmotic pressure of the solute and to provide adequate flux. As
shown in Figure 6.1.2, normal operating concentration ranges vary from very
low values to as high as 60,000 mg/L. As the concentration of the solute
increases during reverse osmosis, additional pressure is required to maintain
the water permeation • rate. At some point in the separation process, further
transfer of water through pressure increases will become impractical because
of membrane and equipment limitations. Other processes (e.g., evaporation)
g
will be needed to achieve higher concentration levels, if necessary.
Electrodialysis processes use an electrical potential gradient and
special synthetic membranes, usually ion exchange type resins, to produce an
enriched stream and a depleted stream. Cation and anion exchange membranes
are arranged alternatively to form compartments in a stack maintained between
two electrodes. Upon application of an electrical field, the ions entering
the compartments within the stack migrate in opposite directions. Depending
upon the selectivity of the membrane, the ion will either pass through the
first membrane it encounters or be held within its original compartment.
Thus, salt solutions are concentrated or diluted in alternate compartments.
Other membrane processes that have been identified in the literature as
potentially suitable for the recovery of metal ions from aqueous solutions
include Donnan dialysis and coupled transport. Donnan dialysis operates on
6-7
-------
the principle thac two solutions separated by a membrane will remain
electrically neutral. Thus, metal ione in a wastewater compartment will
interchange with the hydrogen ions in an acidic solution contained in another
compartment that is separated by a cation exchange membrane from the
wastewater. In coupled transport, a process similar to liquid ion exchange, a
porous membrane, containing a liquid eomplexing agent within its pores, is
used to affect separation.' The metal ions in the wastewater compartment
combine with the conplexing agent and migrate through the membrane to the
second compartment. Here the complex is broken, releasing the metal ions to
solution, with the regenerated completing agent in turn becoming available for
further reaction/interaction with the metal ions in the wastewater.
Despite some success in the laboratory, both Donnan dialysis and coupled
transport have not been commercially applied, largely because existing
membranes have short-life expectancy. Accordingly, these processes are
not discussed at the same level of detail provided Cor the more advanced,
commercialized membrane technologies. Before proceeding with discussions of
tnese technologies, the following subsection will discuss briefly the types of
membranes and commercial designs available for membrane separations.
6.2 MEMBRANE STRUCTURE AND SYSTEM DESIGN
Membrane structures can be classified according to their ultrastructure
as either microporous or asymmetric. The latter are also referred to as
"skinned" membranes. Microporous membranes are designed to retain all
particles above a certain size. However, particles'that are approximately the
same size as the pores may enter into the pores and plug them. Microporcus
structures with pore sizes in the ultrafiltration range (10 to 200 angstroms)
generally have not been very successful. The few designs that are
commercially available have low flux and are subject to rapid plugging.
The development of the asymmetrical membrane by Loeb and Sourirajan in
1960 marked the beginning of modern membrane technology. These membranes are
characterized by a thin "skin" on one surface, usually 0.1 to 0.5 urn in
thickness, while the main body of the membrane supporting the skin is of the
order of 40 to 200 pnrin thickness and highly porous. The combination of a
thin skin, supported by a highly porous substrate, results in high flux with
good selectivity.
6-8
-------
Asymmetric membranes rarely get plugged in the fashion that microporous
structures do, although they are subject to flux lowering phenomena such as
fouling and concentration polarization. These factors are controlled by
pretreatment, system design, and operating conditions. Cleaning cycles
are also used, but there is a danger that the powerful cleaning agents
required can damage or attack the membrane. For example, the cellulose
acetate membrane used for many reverse osmosis applications has limited pH,
temperature, and chlorine tolerance. Thus, cleaning to correct fouling can be
a problem. Second generation membranes are available which minimize
difficulties associated with cleaning. New membranes, such as those being
developed from ceramic materials, may virtually eliminate problems such as
irreversible fouling.
The following discussions identify the types of membranes and system
designs available for specific membrane technologies. Since several
variations are generally available, the user should contact the manufacturers
of such equipment to identify the most appropriate system. Lists of
1 2 12
manufacturers can be found in several references * * and in McGraw Hill's
Chemical Engineering Equipment Buyers' Guide.
6.2.1 Ultrafiltration/Microfiltration Systems
Ultrafiltration membranes are generally not defined by their pore sizes,
which range from 10 to 200 angstroms and higher, but by the size or equivalent
molecular weight of particles excluded. Although the size cutoff is
arbitrary, one definition by Lonsdaie is that ultrafiltration membranes
retain species in the 300 to 300,000 molecular weight range. Because
ultrafilcration deals with the separation of larger molecules, it is.not
suitable for the separation of dissolved metal ions. However, it does find
use as a pretreatment method or as a means of removing chemically precipitated
metallic species. Microfliters with a pore size of greater than 0.1 ym are
also used to effectively collect precipitates.
Ultrafiltration/microfiltration membranes are made, from a wider selection
of polymers than are reverse osmosis membranes. Cellulose acetate and
poiyamide were the earliest of the commercial'membranes. In addition to
these, several other polymeric materials are available. These are comprised
6-9
-------
of thin skin composite membranes formed on. the surface of a porous support
polymer, usually a polysulfone. However, for specific applications, the
composite structure can be tailored from other materials to enhance chemical
and biological resistance and improve other properties such as selectivity.
• Users of ultrafiltration membrane technology have their choice of tour
basic equipment designs; 1) tubular with inner diameters greater than
10 microns; 2) hollow fibers with inner diameters less than about 1,3 taicrons;
3) plate type units; and 4) spiral-wound modules.-* • The tubular module is
the simplest design. However, because of its small surface area per module,
it is only used in specialized applications. The membrane is either inserted
into a porous tube or is cast in place. The feed is pumped through the tube
and the permeate passes radially through the membrane and porous tube out
through an exst line. The concentrate or reject stream exists from the
downstream"side of the tube. Although the tubular modules do have low surface
area, they are easy to clean and are less susceptible to plugging by suspended
solids than are other membrane types.
Hollow fiber membranes (acrylic copolymer) used ia ultrafiltration employ
a membrane skin on the inside of the hollow fiber. Each hollow fiber has a
fairly uniform bore with available sizes ranging from 8 to 49 mil (0.19 to
1.25 mm) in diameter with a cross-sectional thickness of about 200 urn.
Bundles of fibers are normally sealed in a shell and tube arrangement,
although.
The permeate passes through the membrane and fiber wall and is collected
on the outside of the fibers. The concentrate passes out the opposite end of
the fiber bore. Hollows fibers have a fairly low pressure rating; thus, flow
rates (flux) will also be low, A major advantage is the ease of cleaning
achieved through backflushing, due to the self-supporting nature of the fibers.
The plate and frame-type units represent an early design and consist of
membrane covered support plates stacked in a frame'arrangment. Permeate
exists via the support plate and concentrates leave the module opposite the
feed end. Internal flow within the module may be arranged L'n a combination of
parallel and series flow patterns by using section plates. The stacking
arrangement is usually horizontal for ultrafiltration._
The spiral-wound design is compact, relatively inexpensive) and provides
larger surface areas per unit volume of equipment. The spiral-wound modules
consist of membrane and spacer materials that are trapped around a perforated
6-10
-------
center tube which collects the permeate. Figure 6.2.1 shows a schematic of a
spiral-wound structure along with a cross-section which illustrates the flow
path of the permeate. Although economic treatment of larger volumes is
possible with this design, it is more apt to plug than other designs. Also,
it cannot be cleaned mechanically.
The characteristics of the feed will play a major role in determining
which of the membrane materials and designs should be selected. Feeds
containing larger suspended particles are best processed in larger diameter
tubular units.. Other factors, such as ease of cleaning, pressure losses,
degree of concentration, and other considerations, all contribute to the-
overall utility and cost of a system. Pilot plant studies, a service offerred
by many manufacturers, should be undertaken before proceeding with final
.system selection.
6.2.2 tReverse Osmosis Systems
Three types of membranes are conrnierically available for reverse osmosis:
cellulous acetate, aromatic polyatnides, and thin film composites. Cellulose
acetate membranes have high flux and high salt/metal rejection properties and
are relatively easy to manufacture. Among the disadvantages of these
membranes are; (1) a fairly narrow temperature range (maximum recommended
temperature of 3G°C)j (2) a rather narrow pH range (preferably pH 3-6;
(3) poor resistance to chlorine; (4) a tendency to "creep," bringing about a
gradual loss of membrane properties (notably flux); and (5) susceptibility to
raicrobial attack.
The aromatic polyamides (aramids), commercialized by DuPont in 1970, have
an asymmetric structure similar to cellulose acetate. They are not susceptible
to biological attack, resist hydrolysis, and can be operated over a wider
range of pH (pH 3 to 11) and slightly higher temperature (40°C) than cellulose
acetate membranes. However, they are readily degraded by low levels (0.2 ppm)
of free chlorine, a major drawback for some applications.
Thin film composites are formed by depositing a film of a polymeric
material on a porous support structure, usually a polysulfone. An advantage
of these thin film composite membranes is their ability to withstand more
severe environments. However, not all polymers can be fabricated into
6-11
-------
PERMEftTE
PERMEATE
COLLECTION
TUBE
OUTER COVER
FEIO
SPACER
PERMEATE FLOW
Arrow! indieott
Ptrm«(rti Flo*
PRODUCT
BRINE
FEED CHANNEL
SPACER
TRICOT PRODUCT
WATER
COLLECTION
CHANNEL
PRODUCT
ANT<-T£L£SCOP>NS •
DEVICE I
PRODUCT
TUBE
MEMBRANE SURFACE
MEMBRANE SUPPORT
, BACKING
AMESIVE
Figure 6,2.1. Spiral wora cartridge schematic and cross section
showing flow or permeate (feed flow perpendicular),
6-12
-------
structures that ate suitable for reverse osmosis applications. Polysuifones,
for example, cannot withstand high (e.g., about 100 psig) pressures, nor can
they be fabricated with pore sizes with less than a 500 to 1,000 molecular
weight cutoff. Thus, the polysulfones cannot be used directly as asymmetric
membranes for reverse osmosis. However, they find use in ultrafiltration
applications and as backings for certain reverse osmosis barriers such as
polyethylene itnine/toluene diisoeyanates. Membrane development remains a major
focus of membrane technology.
The equipment used to conduct reverse osmosis separations is similar to
that used for ultrafiltration* Tubular, spiral-wound, and hollow fiber
systems are hollow fibers, but the hollow fiber reverse osmosis system does
differ from that used for ultrafiltratioti. This membrane.is made from an
aromatic polyariide, with an inside diameter o£ about 42 urn and an outside
diameter of about 85 JOT« The fiber has-an asymmetric structure. However,
unlike the ultrafiltration hollow fibers, the skin is located on the outside
of the fiber, necessitating the employment of a different system configuration
to separate the permeate (which flows through the bore of the hollow fibers) and
the concentrated feedt Up to 4.5 million of the fibers can be assembled into
a bundle for use in reverse osmosis equipment. This system provides the
highest membrane area per unit volume for any reverse osmosis system.^
6.2.3 Electrodialysis Systems
Electrodialysis is based on the migration of ions through seta of
alternate cation and anion exchange selective membranes that permit the
passage of positive and negative ions, respectively. The selective _
membranes should possess the following characateristics:
o low electrical resistance;
o good selective qualities;
o good mechanical properties;
o good structural stability; and
o high chemical stability.
6-13
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Since it is difficult- to optimize these properties, only a few companies
produce electrodialysis membranes commercially' Two general types of
membranes are available, heterogeneous and homogeneous. Heterogeneous
membranes are manufactured by nixing a commercial ion exchange resin (50 to
70 percent) with a binder polymer such as polyvinyl chloride. The ion
exchange resins are usually crosslinked copolyners of styrene and
divinyibenzene. Cation or anion exchange groups are introduced into the
copolymer by sulfonation and chloromethylation/amination with a triaraine.
Plastic mesh or cloth is used as a support for the ion exchange/binder mixture
that constitutes the membrane.
Homogenous membranes consist of a continuous homogeneous film onto which
an active group is introduced. The membranes can be reinforced or
nonreinfcreed. Some of the properties of some commercially available
membranes are shown in table 6.2.1, as provided in E> Korngold's chapter on
electrodialysis in Reference 2. A more extensive description of these and
other specialty membranes is provided in other references. "
In an electrodialysis stack, cation and anion exchange membranes are
alternated between two electrodes. AB shown in Figure 6.2.2, during
electrodialysis, one cell will contain a concentrated solution and the other
will contain a dilute solution. In industrial units, several hundred cell
•pairs can be assembled between two electrodes. The system can be used for
desalination, separation of nonelectrolytes from electrolytes, and electrolyte
concentration. Power consumption is directly proportional to ion
concentration and operating costs are more favorable for low feed
concentration. Also, because ion rejection varies from about 45 to 55 percent
per pass (as opposed to up to 99 percent for reverse osmosis), a number of
passes oust be used if a low concentration of dissolved solids is required in
the dilute stream.
The application of reverse osmosis for seawater conversion is much more
advanced than electrodialysis technology. Because of the high energy
consumption required at seawater concentration levels (approximately
32,000 ppm), electrodialysis is not economically attractive, although programs
to develop high temperature operation and new membranes may improve the
economics considerably. Cost data from the early 1980s indicate that
electrodialysis becomes competitive for desalting at levels of roughly
2,000 ppm total dissolved solids.
6-14
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TABLE 6.2.1. PROPERTIES OF COMMERCIALLY PRODUCED MEMBRANES
I
(—'
Ui
Manufacturer
lonac Chemical Co.
New Jersey
American M.icliinc and
lTouiulry
Connecticut
Ionics Inc.
Massachusetts
Asalii Glass Co. Ltd.
Tokyo, Japan
Tokuyama Soda Ltd.
Tokyo, Japan
Asalij Chemical Industry
Co. Ltd.
Tokyo, Japan
Ilcn-Ginion University ol'
the Negcv, Research &
Development Authority
lieershcva, Israel
Name of
membranes
lonac
A.M.I-.
Nepton
Sclemion
Ncosepla
A.C.I, or
Acipex
Ncginst
Membrane
MC-3142
MC-3470
MA-3148
MA-3475
IM 12
C-60
A-60
CK6I AZL 183
AR III DZL 183
CMV
AMV
CL25T
AV4T
UK 1
DA 1
NIIGINST-IID
NBGINST-IID
NEGINST-HC
NEGINST-IIC
Thickness
(mm)
0.15
0.35.
0.17
0.40
0.13
0.30
0.30
0.60
0.60
0.15
0.14
0.16
0.15
0.23
0.21
0.35
0.35
0.2
0.2
Capacity
(meq/gm)
1.06
1.05
0.93
1.13
—
1.5
1.6
2.7
1.8
1.4
1.8-2.0
1.5-2.0
2.6
1.5
0.8
0.8
1.6
1.7
Electrical resistance
(f) cm2 in 0.1 N NaCI)
9.1
10.5
10. 1
23
4
6
5
9
14
61
4.0
3.5
4.0
6.5
4.5
12
10
6
8
Reinforcement
Yes
Yes
Yes
Yes
Yes
No
No
Yes
Yes
Yes
Yes
Yes
Yes
Yes
Yes
Yes
Yes
No
No
Source: Reference 2.
-------
ELECTRODE
RINSE WATER
Cathode
©
FEED WATER
Dilute Concentrate
Stream Stream
ELECTRODE
RINSE WATER
I
Anode
©
Permeable Membrane
(T)Anlon Permeable Membrane
Figure 6.2.2. Diagrammatic representation of electrodialysis.
Source: Reference 6.
-------
-------
6.2.4 Other Membrane Systems
As noted previously, Donnan dialysis and coupled transport are two
membrane technologies that appear to have some application for the treatment
of metal-containing aqueous wastes. However, problems of membrane stability
have limited their development. DuFont presently markets "Safion", a
perfluorosulfonic acid membrane that is being evaluated as a Donnan dialysis
membrane for removal of nickel in electroplating wash water. Anion exchange
membranes for the removal of copper, cadmium, and zinc cyanide complexes are
also being evaluated. Referenced reports studies using auarternized
polyvinyl pyridine and polyvinyl benzylcbloride films grafted on a
polyethylene base. Ion transport rates were reportedly proportional to ion
exchange capacity.
The coupled transport process relies on a liquid, water immiscible,
organic completing agent held within the pores of a micrpporous membrane. The
metal ion, introduced with the feed solution to one side of a cell reacts with
the liquid. It is then transported through the membrane to the product
solution where it is released and the organic transport medium is
regenerated. Both Donnan dialysis and the coupled transport processes operate
without the need for electric current or .high hydraulic pressures as required
for electrodialysis and reverse osmosis, respectively. The only energy
required is that needed to pump the feed and stripping solution through the
cells.
6.3 ULTRAFILTRATION/MICROFILTRATION FOR TREATMENT OF METAL WASTES
As a result of the high molecular weight cutoff (approximately 1,000) of
ultrafiitration membranes, they cannot be applied directly to recover metals
present as dissolved solids in, for example, electroplating rinsevaters.
Ultrafiitration has been used commercially to recover or treat electrophoretic
paints, oil in water emulsions, proteins from the dairy industry, and
rinsevaters from alkaline metal cleaning baths. However, there are no known
applications for the recovery of metals from aqueous waste streams, with the
exception of its use ae a pretreatment method or as a means of recovering
precipitated materials. Consequently, the following discussions will address
membrane systems used to separate precipitated metals-from waste streams.
6-17
-------
-------
The following discussion relies heavily on material provided by Memtek
Corporation, a supplier o£ this technology.'5-17 Strictly speaking, the
Merate'* technology is microfiltration, rather than ultrafiltration, since a
membrane with pore sizes of the order of 0.1 urn is used. Although the nature
of the metabrane(s) is proprietary, it reportedly is inert and can withstand
any solution pH.
6.3.1 Process Description
Advanced membrane processing of wastewater, as described by Memtek,
utilizes a system where insoluble precipitated contaminants are separated from
solution through chemical pretreatment followed by the use of cross-flow
tubular membranes to contain the precipitate. These membranes have a nominal
pore size of 0.1 microns which allows for complete rejection of all
particulate or euspended solids larger than this size. The mechanism requires
the solution to be pumped under low pressure and turbulent flow conditions
down the center of a membrane module. The pressure exerted on the solution
forces clean solution through the membrane. The solids rejected at the
membrane surface are carried by the flowing liquid back to the beginning of
the system. This design allows solutions of up to 10 -percent solids to be
filtered while producing particulate free effluent and concentrating the
solution to a higher percent solids (see Figure 6.3.1).
The system differs from conventional membrane systems by incorporating
chemical pretreatment to render metal ions insoluble. Since all precipitated
solids and turbidity are retained by the membranej,the effluent quality is
related to the residual soluble ions* When applied to heavy metal wastes with
appropriate pretreatment chemistry, including co-precipitation effects, the
toxic metal content of the effluent can be extremely low.
The basic components of the system are a chemical reaction section,
concentration tank, process pump, membranes, and an integrated cleaning
system. These systems utilize low pressure pumping (40 psi) and high
turbulent flow (15 fps) through the membrane to effect good filtration rates.
Typical design flow rates for these membrane systems used in industrial
wastewater applications are 200 to 400 gfd (gal/fL2 of membrane surface area
per day). Filtration capacity is provided by installing the required number
of modules in parallel to produce the design effluent.
6-18
-------
VO
CONCINmATION TAN!
ClIANINO
Figure 6.3.1. A typical simplified flow diagram of wastewater
treatment for computer related industries.
Source: References 15 and 17.
-------
The pretreatment required to attain effluent specifications is specific
to the particular application. For example, a different pretreatment program
is required for ehelated compounds, hexavalent chrome, cyanide, and hydroxide
reactions. The goal of all chemical pretreatment is simply to convert
dissolved ions to precipitated compounds so that they can be effectively
removed to produce the required effluent concentrations. Using variants of
standard pretreatment chemistry, some systems have been designed to produce an
effluent in the parts per billion range. However, the presence of oil and
grease can cause premature fouling of the membrane] If present, additional
pretreatment will be required to remove these constituents.
Each membrane section is provided with an integrated cleaning system to
restore the membrane performance when fouling occurs. This procedure uses
only chemical cleaning processes to restore flow rate to the design level.
The chemical cleaning procedure is designed to complete the cleaning in a
short time period, typically less than 2 hours, applied once each week.
During treatment, the solids content of the feed increases in the
membrane modules to form s concentrated slurry with.generally 2 to 5 percent
solids. To increase this concentration, a portion of the slurry is removed
from the system and routed to a settling tank or filter press where further
sludge consolidation occurs; e.g., 5 to 20 percent or 30 to 40 percent solids,
respectively. . :
6.3.2 Process Performance
Table 6.3.1 summarizes operational performance for a range of typical
applications for ultrafiltration. The table shows the ability of the membrane
to process high and low concentration wastes to extremely low effluent levels.
6.3.3 Costs
Costs for the various sizes of standard advanced membrane filtration
units are presented in Table 6.3.2. The cost estimates include treatabiiity
study, engineering,, all pretreatment equipment, control panel, piping,
membranes, pumps, installation and start-up. These costs are based on typical
pretreatment equipment, but do not include costs for pretreatment equipment
such as cyanide destruction. Operation and maintenance costs include
6-20 :
-------
TABLE 6.3.1. TYPICAL SYSTEM PERFORMANCE
Contaminants
Aluminum
Arsenic
Cadmium
Chromium
Copper
Cyanide
Fluoride
Gallium
Germanium
Gold
Iron
Lead
Manganese
Mercury
Nickel
Rad ium*
Rhodium
Silver
Tin
Uranium
Zinc
BOD
COD
Suspended solids
Feed (mg/L)
10-1000
1- 43
25- 115
3- 275
1-1525
5- 300
18-5000
4- 20
20- 110
1- 12
2-1500
2- 25
1- 10
3- 30
4- 300
1- 10
20- 500
10- 200
20- 75
1- 15
2- 400
50-5000
20-3500
—
Effluent (mg/L)
0.5
0.05
0.05
0.1
O.-l
0.1
1.0
0.5
0.5
0.15
0.02
0.05
0.02
0.02
0.02
0.6
0.1
0.1
0.1
0.001
0.1
**
**
Non-detectable
Concentration given in picocuries/liter.
** 95 percent removal.
Source: Reference 15.
6-21
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TABLE 6.3.2. ESTIMATED CAPITAL AND OPERATING COSTS OF STANDARD
ADVANCED MEMBRANE FILTRATION SYSTEM
Flow rate (gpm) Dimensions
20
45
75
100
150
500
6'W
6'W
12'W
12'W
18'W
40 'W
x i&'L x
x 24 "L x
x 24 'L x
x 45'L x
x 55 'L x
x 60'L x
Capital cost ($)
ll'H
ll'H
ll'H
ll'H
ll'H
ll'H
90,000
130,000
170,000
225,000
330,000
900,000
Typical8
O&M cost ($/yr)
8,500
15,400
19,100
25,500
37,400
113,300
^Depends greatly on influent waste characteristics.
Source: Reference 15, 1987 costs.
6-22
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electricity, routine periodic repairs, and routine cleaning. They do not
include costs for pretreatment chemicals, labor, post-treatments, or membrane
replacement. As shown in Table 6.3.3, these costs are highly site—specific
and can be appreciable.
6.3.4 Overall Status of Process
Availability™"
Over 100 full-scale industrial systems, ranging in size from 10 to
400 gpm, have been installed by Memtek at.facilities that include printed
circuit board manufacturing, electroplating, battery manufacturing, and
photographic processing. Each application will require treatability studies
to optimize and integrate the chemical pretreatment and membrane systems.
Application—
This system can be applied to metal-containing aqueous waste streams
provided solubility limits associated with the chemical precipitation step are
within required limits. In most instances this does not appear to be a
problem. Membrane fouling may also impose limitations, although Memtek
reports that cleaning will restore flux rates.
Environmental Impact—
Assuming that concentration levels found in the permeate are below
regulatory limits, the principal environmental impact will result from the
sludge generated by the process. Dewatering, followed by solidification,
encapsulation, or some other treatment method, will bejrequired.
Advantages and Limitations™
Chemical precipitation followed by microfiltration appears to be an
effective means of reducing the heavy metal contaminant levels of aqueous
waste streams. Treatment processes can effectively reduce contaminants levels
down to 100 ppb or lower. The principal limitation results from the hazardous
sludge generated which must be treated before it can be land disposed. The
cost of treatment and disposal will depend upon the contaminant and its
concentration and could be appreciable.
6-23
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'i
to
TABLE 6.3.3. ESTIMATED OPERATING COST FOR MEMBRANE FILTRATION SYSTEMS
TREATING METAL-CONTAINING WASTEWATERS3
Waste description
General rinse water"
Chelated wastewater
Wastewater of chelates,
non-chelates, and spent
concentrates
Same as above
Same as above
System
size
(gpm)
—
40
100
60
35
Metal
concentratio
in feed
(ppm)
17
480-665
27-70
56-105
46-85
Costs
Chemicals
0.05
9.10
- 0.80
4.00
3.40
($/l,000 gallons of feed)
Electricity
0.60
0.80
0.70
0.70
0.75
Solids
disposal
0.20
5.20
0.83
0.40
0.60
Total
0.85
15.10
2.33
5.10
4.75
aWastes from printed circuit board producers containing primarily copper
with some lead and nickel.
"Also contains~2 ppm chrome.
Source: Reference 15, 1987 costs.
-------
6.4 REVERSE OSMOSIS
6.4.1 Process Description
Reverse osmosis (RO) is a treatment technique used to remove dissolved
organic and inorganic materials, and to control amounts of soluble metals,
TDS, and TOG in wastewater streams. The technology has been applied in the
metal finishing industry to recover plating chemicals from rinsewater, such
that both plating chemicals and rinsewaters can be reused,
RO involves passing wastewater through a semipermeable membrane at &
pressure greater than the osmotic pressure caused by the dissolved materials
in the solvent. Thus, the osmotic flow, defined as the flow from a
concentrated solution to a dilute solution, is reversed due to the increase in
pressure applied to the system. The process is schematically presented in
Figure 6.4.1.
To obtain reasonable water fluxes (approximately 10 gal/ft day), the
feed solution must be pressurized well above the equilibrium osmotic
pressure. The expression for osmostic pressure can be written as:
CRT
where C is the volume concentration, R has the same value as the universal gas
constant, and T is the absolute temperature. In practice, reverse osmosis
systems are operated from about 4 to 20 times the equilibrium osmotic
2
pressure, with pressures of 1,000 psi or greater not uncommon.. As shown in
Figure 6.4.2, osmotic pressure increases with solute mass fraction and
decreases with molecular weight. As the concentration of the solute increases
during reverse osmosis, additional pressure must be applied to maintain flux.
The detrimental effect of increasing concentration is further complicated by
concentration polarization, a term which refers to accumulation of solute at
the surface of the membrane resulting in a further increase in osmotic
pressure. Proper design will minimize the polarization effects, e.g., through
the use of turbulence within the feed stream.
6-25
-------
Pressure
Semipermeable
™" membrane
Semipermeabls
memftrane '
a. Osmotic flow
b. Osmotic equilibrium
c. Reverse osmosis
Figure 6.4.1. Principles of normal and reverse osmosis.
Source: Reference 9.
Ideal aqueous solution
25'C
0.2 -
0,1
0.01 0.02 0.05 0.1 0.2 0.5 1.0
Solute mass fraction
Figure 6.4.2, Osmotic pressure as a function of mass fraction
and molecular weight,
Source: Reference 11.
6-26
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The operation of a reverse osmosis system is affected primarily by the
.feed characteristics, operating pressure, and membrane type. These factors
will affect the flux and percent rejection which/ in turn, define system size
requirements and effluent quality, respectively.
Flux determines the system size for a given waste flow rate; i.e., higher
flux permits the use of smaller systems. Flux is the volume flow of permeate
per unit membrane area. It is proportional to the effective pressure driving
18
force, according to the following relationship:
J = K UP - ATT)
where J is the flux, K. is the membrane constant, APIs the applied pressure
across the membrane, and £,TT is the osmotic pressure across the membrane.
Since osmotic pressure is approximately proportional to molar feed
concentration, flux increases with increasing operating pressure and decreases
with increasing feed concentration. Thus, chemicals which form high-molecular
weight complexes will have higher flux for & given weight percent in
solution. More concentrated solutions can be achieved by utilizing 3 large
effective driving pressure. Increases in temperature of the waste feed will
also increase the flux by lowering viscosity. However, although increased
operating temperatures will improve the performance of the system in the
short-term, the lifetime of the membrane will be shortened.
18
Percent rejection is defined as follows:
(feed concentration) — permeate concentration , „_„,
% Rejection = rr—: :—r- \ ' x 1004
J (.feed concentration)
Higher percent rejections will result in better quality (higher purity) of the
permeate and concentrated streams. Percent rejection is primarily affected by
the membrane type, although rejection will decrease with increasing feed
18
concentration.
The application of reverse osmosis to the treatment of metal-containing
wastes is often limited by the pH range in which the membrane can operate.
Table 6.4.1 shows the characteristics cf setae ccczsercielly available
6-27
-------
TABLE 6.4.1. COMMERCIALLY AVAILABLE MEMBRANE MATERIALS
I
M
00
Type
Hollow fiber
Spiral wound
cellulose
acetate
RC-100
Description
Hollow fine fiber
asymetric membranes
or aromatic polyamide
Flat sheet composite
membrane of cellulose
acetate with mesh
spacers, rolled into
cartridge
Flat sheet composite
membrane of polyether/
amide on polysulfone,
rolled into cartridge
Allowable
PH
Source range
E.I. DuPont 4-11
Wilmington, DE
Osmonics, Inc. 2.5-7
Hopkins, MN
Fluid System
Div. of UOP, Inc.
Dow Chemical USA
Midland, MI
Fluid Systems 1-12
Div. of UOP, Inc.
San Diego, CA
Typical
operating
pressure
kg/cm2(psig)
29.2-58.4
29.2-58.4
(400-800)
29.2-58.4
(400-800)
Flux rate
L/m2/day
(gal/ftz/day)
@ 77°F & 400 psi
73(1.B)<9>
1140(28)
530(13)
Typical flux
per module
L/day(gal/day)
3785(1000)
5680(1500)
3785(1000)
Module
replacement
coats (1)
750
350
1000
Source: Reference 19-
-------
membranes. The cellulose acetates have a very small pH range, and thus cannot
be used for recovery of cyanide plating baths where pH is well'above 7. The
polyetber/amide on polysolfone material appears to result in the membrane that
is least affected by pH.
Many seraipermeable membranes can be fabricated either in the form of a
sheet or tube, which is then assembled into modules. Figure 6.4.3 shows the
three basic module designs, which include:
Tubular—A porous tubular support with the membrane case in place or
inserted into the tube. Feed is pumped through the tube,
concentrate is removed downstream, and the permeate passes through
the membrane/porous support composite.
Spiral Hound—Large porous sheet(s) wound around a central permeate
collector tube. Feed is passed over one side of the sheet and the
permeate is withdrawn from the other,
Hollow Fiber—Thousands of fine hollow fiber membranes (40 to 80 ym
diameter) arranged in a bundle around a central porous tube. Feed
enters the tube, passes over the outside of the fibers, and is
removed as concentrate, water permeates to the inside of the fibers
and is collected at one end of the unit.
Reverse osmosis systems typically consist of a number of modules
connected in series or parallel, or a combination of both arrangements. In a
series arrangement, the reject stream' from one module is fed directly to
another module, such that greater product concentration is achieved.
Alternatively, the reject stream may be recycled to the feed stream of the
same unit. Series treatment inay be limited in some cases by the ability of
the membrane to withstand concentrated contaminants. The system capacity can
be increased through the use of a parallel arrangement of modules; however,
product quality will not be enhanced. Schematic flow diagrams of two series
systems are shown in Figures 6,4.4 and 6,4,5.
To ensure a minimum permissible reject flow rate per nodule, and thus
provide adequate turbulence, each successive stage contains a smaller number
of modules than the preceding stage. The system shown in Figure 6.4.4 is
designed for 87.5 percent water recovery. While the degree of rejection is
dependent on the particular ion under consideration, Figure 6.4.5 shows how
higher pyrity water can be obtained by feeding the produce to a second stage.
6-29
-------
A.TUBULAR MEMBRANE.
b. S?iRAL-WOUNO MODULE
SOU TO
ASSEMBLE ,-•;.-*
FLOW
ATs OUT
fEsaciTt 5106 B»CK)NC »
MATSBIAL WITH Mf.MBR*NE ON "%
E*CK SIDE *ND CLUE5 ancuMD
E»5sS AND TO CSNTift tusi
C. HOILOW-FI3SS MOOULt
=NO
POBOUS "5=3 ]/
TOR -0" SINC END PLATS
Figure 6,4.3. Reverse osmosis membrane module configurations.
Source: Reference 20.
6-30
-------
J.000 pern
0.50 mad
8' Bora
—< STAQE 2
PRODUCT
0.875 iT>9d
I 72 cpm
0,25 msd
133 psm
0,125
470 ppm
O.SOmgd
0 25 mgts
7,630 ppni
nJ
REJECT
0.12Sm^J
14,?80pon>
Figure 6.4.4. Schematic of a three-stage RO plant.
Source; Reference 2.
R.Q. PLANT FEED
2.00 ft\5d
WASTE STREAM |
Mlmga I
2.000 W>m TDS
STAGE 1
1 «0 mgd
126 ppm
STACt i
J5 ps
RECYCLE
0.28 msd
527 ppm
0.60 c.»d
5.SSS pom
Figure 6.4,5. RO system to increase-product purity.
Source; Reference 2.
6-31
-------
Although they are able to operate at higher pressures, tubular modules
are not applicable for most industrial applications because of targe floor
space requirements and high capital costs. Comparatively, hollow fiber and
spiral wound modules have lower, and similar, capital costs- Hollow fiber
modules require less floor space, but spiral wound nodules are not as
19
susceptible to plugging by suspended solids. For best performance, the
feed to any of these systems should be treated to remove gross amounts of
solids and to prevent fouling by precipitation or biological growth.
Pretreatoent Requirements*™
Colloidal and organic matter can clog the membrane surface, thus reducing
the available surface area for permeate flow. Also, low-solubility salts will
precipitate on the membrane during the concentration process, similarly
reducing membrane efficiency. Pretreatment techniques such as pH adjustment,
activated carbon adsorption, chemical precipitation or filtration
(approximately 5ym) may be required to ensure extended service life.
Operating costs for membrane systems are a direct function of the
concentration of the impurity to be removed, due in part to increased
maintenance and membrane replacement costs.
Multi-charged cations and anions are effectively removed from the
wastewater by reverse osmosis. However, most lov molecular weight, dissolved
organics are, at best, only partially removed. Their presence could present
operational difficulties and may require expensive pretreatment for their
removal. The use of reverse osmosis for recovery/reuse of process wastes is
also currently somewhat limited because many membranes are attacked by
solutions with a high oxidation potential (e.g., chromic acid) or excessive pH
levels. However, future development of membranes which are able to withstand
21
harsher environments is expected.
Post-Treatment—
Reverse osmosis applied to plating bath wastes is usually supplemented
with an evaporation system in order to adequately concentrate constituents for
21
reuse. The amount of feed concentration permitted in a unit is limited by
6-32
-------
Che membrane characteristics. Reverse osmosis units can concentrate moat
plating and other systems operated at ambient temperatures where atmospheric
21
divalent metaLs (nickel, copper, cadmium, zinc, etc.) from rinsewaters to a 10
21
to 20 percent solution. Further concentration must be achieved through
the use of a small evaporator. Evaporators are especially necessary for
plating and other systems operal
evaporative losses are minimal.'
6.4.2 Process Performance
Reverse osmosis applications in recycling many electroplating bath wastes
are somewhat limited due to membrane degradation in the extreme pH regions.
However, research has been conducted to recover both acidic and basic plating
rinsewaters, which has led to the development of more chemically resistant
membranes.
The Waiden Division of Abcor, Inc. (Wilmington, MA) conducted a number of
studies of reverse osmosis systems for recovery of plating rinsewaters.
Initially, studies were conducted to test the applicability of membrane types
to various plating rinses. Test samples were prepared by diluting actual
18
plating bath solutions with de—ionized water. Bath properties (total
dissolved solids, pH) and test solution properties (concentration, pH) for the
wastes tested is presented in Table 6.4.2. The percent TDS rejection values
shown in the table were averages for tests at the low end of the concentration
levels. Percent reduction values for the metals of interest were generally
slightly higher. Overall, membrane performance was affected by feed
constituent concentrations, operating pressure, operating temperature, flow
rate, and pH. Flux and rejection data were not affected by changes in pH, but
extreme pH values were found to decrease membrane life.
Additional tests were performed to evaluate the effects of feed
concentrations on flux and percent-rejection. A summary of the operating
parameters and results for some of these tests is presented in Table 6.4.3 for
the chromic acid rinse. As shown, both flux and rejection decrease somewhat
with increasing feed concentration. These results are typical for all of the
wastes studied. While rejection and flux results were satisfactory for the
chromic acid test runs, hydrolysis and degradation of the membranes occurred.
6-33
-------
TABLE 6.4.2. SUMMARY OF REVERSE OSMOSIS EXPERIMENTS
Properties of
test solutions Average % TDS
rejection by
Properties of bath Concentration module type
range
Plating baths Source of bath % TDS pH (% TDS) pH range ABC
Chromic Acid Whyco Chromium Co. 27.5 — 0.3-4.5 A.5-6.1 98 94 97
(Neutralized) (37.1)
Chromic Acid Whyco Chromium Co. 27.5 0.53 0.4-9.0 1.2-1.9 90 94 97
(Unneutralized)
Copper Pyrophosphate Honeywell (M&T) 31.9 8.8 0.2-11.4 6.8-8.5 96 97 97
Nickel Sulfnmate Honeywell (llarstan) 31.0 4.2 0.5-12.0 4.9-6.1 91 93 91
Nickel Fluoborate llampden 25.7 3.5 0.9-5.8 3.4-6.1 64 — 91
o> Colors & Chemicals
** Zinc Chloride General Electric 19.8 4.5 0.2-4.2 5.3-6.1 9Z 90 89
(Conversion
Chemical)
Cadmium Cyanide American 26.3 13.1 0.3-3.1 11.5-12.5 95
Electroplating Co.
Zinc Cyanide American 11.4 13.9 0.5-2.4 12.3-13.3 95
Electroplating Co.
Copper Cyanide American 37.0 13.3 0.6-3.7 11.8-12.5 98
Electroplating Co.
Rochelle Copper Whyco Chromium Co. 12.7 11.2 0.13-3.2 9.8-10.6 99
Cyanide
A - DuPont B-9 permeator, polyamide hollow-fiber membrane.
B - T.J. Engineering 97H32 spiral-wound module; cellulose acetate membrane.
C - Abcor TM 5-14 module, tubular configuration; cellulose acetate membrane*
Source: Adapted from Reference 18.
-------
TABLE 6.4.3. ANALYTICAL RESULTS FOR REVERSE OSMOSIS TREATMENT
OF SPENT CHROMIC ACID PLATING RINSE
f t a. **, *» *"* 4*A«/4
waste reed
module3 %-TBS % of bath
Hollow fiber
Spiral 0.40 1.5
Tubular
Hollow fiber
Spiral 1.83 6.7
Tubular
Hollow fiber
Spiral 4.11 15
Tubular
Hollow fiber
Spiral 9.43 34
Tubular
Operating conditions
Pressure Tetnp, pH of
(psig) (°G) feed
400
600 29 1.9
800
400
600 29 1.2
800
400
600 29 1.2
800
400
600 28 0.9
800
% - Rejection
Fluxb
2.59
15.3
10.0
1.97
13.2
8.58
1.20
10.6
7.31
leak
leak
6.60
Basis :
IDS
84
97
99
95
94
97
90
92
95
leak
leak
94
Basis:
Cr+6
97
96
98
87
86
91
91
92
7
leak
leak
97
aThree commercially-available membrane modules were tested:
DuPont B-9 .hollow-fiber module (polyamide membrane);
T.J, Engineering 97H32 spiral-wound module (cellulose acetate
membrane); and
Abcor, Inc., TM5-14 tubular module (cellulose acetate membrane).
bGallon/minute/single DuPont B-9 permeator size 0440-035.
Gallon/day/ft^ for spiral-wound and tubular modules.
Source: Reference 18.
6-35
-------
Other available data, widely scattered throughout the literature,
indicate that reverse osmosis is a proven technology for treating
electroplating waatewater. Systems are being used commercially to recover
brass, hexavalent chromium, copper, nickel, and zinc from metal finishing
solutions. While the ultimate goal is zero discharge, evaporators may be
needed to concentrate the solution to the required bath strength. Rinses such
as Macts nickel, bright nickel, and nickel sulfamate, all can be treated in a
22
zero discharge system; however, duplex nickel cannot* Some typical
membrane rejection values for cations and anions are shown in Table 6.4.4.
They are similar to values reported by other sources. * ' '
6.4.3 Cost of Treatment
Capital costs for reverse osmosis systems vary with operating parameters,
membrane type, modular design, and waste feed characteristics. According to
Reference 10, the capital and annual operating costs for a typical reverse
osmosis system used in the electroplating industry were $20,000 and $5,0,00,
respectively. Due to savings associated with recovery of plating chemicals,
wastewater treatment, and sludge disposal, the payback period was 4.3 years.
The above values represent 1979 dollars, but present day costs do not appear
to have increased appreciably.
Capital costs are primarily a function of the membrane surface area
needed to provide the necessary flux. The packaged reverse osmosis units
available from manufacturers contain a fixed number of membrane nodules along
with auxiliaries such as a feed pump, prefilters, and other equipment needed
for pretreatment. The packaged membrane modules can be readily replaced or
expanded if the need arises. Installation costs are minimal since the units
are normally skid mounted and require only utility connections.
The capital cost of a system is approximately the same for spiral-wound
or hollow-fiber membrane units; tubular units are more expensive', but may be
required if fouling is a problem for a particular waste. Costs are also very
similar for most of the commercially available reverse osmosis membranes;
i.e., cellulose acetate, polyamide, and polyether/amide. Figure 6.4.6 shows
the relationship between equipment costs and membrane surface areas fcr a
6-36
-------
TABLE 6.4.4. MEMBRANE REJECTIONS
Name
Cations,
Sodium
Calcium
Magnesium1
Potassium
Iron
Manganese
Aluminum
AfflffijQniuiB
Coppec •
Nickel
Strontium
Hardness
Cadsiura
Silver
Anions
Chloride
Bicarbonate
Sulfat*
Nitrate
Fluoride
Silicate
Phosphate
Bromide
Borats
Chromate
Cyanide
Sulfite
Thiosulfate
Ferrocyanide
Percent
rejection
94-96
96-98
96-98
94-96
98-99
98-99
99*
88-95
96-99
97-99
96-99
96-98
95-98
94-96
94-95
95-96
99*
93-96
94-96
95-97
99*
94-96 .
35-70**
90-98
90-95**
98-99
99*
99*
Maximum
concentration
percent
3-4
*
*
3-4
*
*
5-10
3-4
fl-10
10-12
~
*
8-10
*
3-4
5-8
8-12
3-4
3-4
—
10-14
3-4
—
8-12
4-12
8-12
10-14
8-14
*Kust watch for precipitation, other ion
controls naximuc concentration.
**Dependent oo pH.
Source: Reference 23.
6-3?
-------
30
cr
oo
I-
10
O
U
uj
c\l
a
IL
<
U
25
20
15
Minimum Size
Unit
NOTE: Unit Includes Prcfillcr, Illyh Pressure F«:cc) Pump,
Mcinbrnno Modules, activated Carbon filler. Auxiliaries,
Preasscmblcd, rcc]iilrinrj ulllM/ connections only
Basis - Cellulose Acetate Membr.-inc
10
_L
J_
100 200 300 '100 500 GOO
MEMBRANE SURFACE AREA
700
000
900
Figure 6.4.6. Reverse osmosis system capital cost vs. membrane surface area.
Source: Reference 25.
-------
spiral-wound, cellulose acetate membrane system. The costs are in 1983
dollars, but are not appreciably different today if the Chemical Plant Index
is used as an indicator (a value of 316.9 in 1983 versus 319.7 in April 1987),
The low operating cost of a reverse osmosis system is one of the most
attractive features of the technology. The only utility needed for the system
is electricity and the feed pumps generally draw less than one kilowatt of
power. However, additional costs may be incurred for membrane replacement and
feed pretreatment. Also, additional eiepense may be required for an evaporator
when reverse osmosis alone is not capable of providing a output that is
concentrated enough for direct reuse as an electroplating solution. In the
case of a zinc cyanide system, a cost for an evaporator of $40,000 was
estimated to supplement a £25,000 expenditure for a reverse osmosis system. A
810,000 annual savings was insufficient to offset the $12,000 operating costs,
for the combined reverse osmosis/evaporator system.
However, where reverse osmosis can be used to produce a satisfactory
plating solution, chemical recovery benefits can be appreciable. Figure 6.4.7
shows the savings possible from the recovery of nickel salts from a Watts
nickel plating line. A detailed analysis of the economics of a reverse
osmosis installation for drag-out recovery is presented in Table 6.4.5. The
values are given in 1983 dollars; as noted, values are not appreciably
different from present-day values if the Chemical Plant Index is used as a
cost indicator.
In summary, the cost-effectiveness of reverse osmosis is dependent upon
the following factors: production rate, type and concentration of rinsewater
constituents, water supply, wastewater disposal costs, and useful lifetime of
membranes. As more chemically resistant membranes are developed, reverse
osmosis systems will have more cost-effective applications for metal
containing waters. Also, with the implementation of the land disposal ban and
the resulting rise in sludge disposal costs, reverse osmosis will become a
more cost-effective alternative to conventional neutralization practices.
6-39
-------
120
DRAG-OUT (Cal/h)
BES:S: SOOO hours per year Operation on V/sus Nickel
Plating line Rinse Tank
Assume 90t Overall Drag-Out Recovery : Savings Based
on Chemical Recovery Plus Reduced Pollution Control
Costs,
Bath Composition— NiSQ
' NiCI
HJjoz/gal 6 $1.30/'b
6 oi/gal Q $1.75/ib
5 qz/cal 8 50.eS/lb
figure 6.4.7. Annual savings from nickel plating drag-out recovery,
Source: Reference 25,
6-40
-------
TAJLE 6.4.5. ECONOMICS OF REVERSE OSMOSIS SYSTEM FOR NICKEL
SALT RECOVERY, OPERATING 4,000 h/yr
'Item Amount
Installed cost, 530-ft2 unit ($):
Equipment:
RO system including 25 urn filter, pump less 17,000
ten membrane units
Activated carbon filter 2,000
Auxiliaries, piping, and miscellaneous 3,OOP
Subtotal: 22,000
Installation, labor and material 3,OOP
Total installed cost 25,000
Annual operating cost (l/yr);
Labor and naintenance at JlO/hr 1,600
General plant overhead 1,000
Raw nuterials;
Module replacement, 2-year life 1,800
(10 x $350/module) x 0.5 yr
Carbon f°r carbon filter
Prefilter element (25 un)
Electricity costs (S0.45/kwh)
Total operating cost: 6,700
Annual fixed costs ($/yr);
Depreciation, 10 percent of investment 2,500
Taxes and insurance, 2 percent of investment 500
Total fixed costs: ' 3,000
Total cost of operation: 9,700
Annual savings ($/yr):
Plating chemicals;
4 Ib/hr nickel-salt at ll/lb 16.000
1.5 02/hr brightener at $0.10/az 600
Hater and sewer charges: saving 270 gal/hr at
60.80/1,000 gal ^.900
Total gross annual savings: 17,500
Net savings = annual savings - (operating cost +
fixed cost) (t/yr) 7,800
Net savings after taxes, 452 tax rate
7,800 x. 0,55 + 2,500* (t/yr) 6,800
Average ROI " net savings after taxes/total installed
investment x 100 (X) . 27
Cash flow from investment • net savings after taxes +
depreciation (S/yr) 9,300
Payback period • total installed investment/cash
flow (yr) 2.7
aiOZ investment tax credit - £2,500 (or 0.10 x 25,000).
Source; Reference 25.
6-41
-------
6.4,4 Overall Status of Reverse Osmosis
Availability—
Reverse osmosis technology is available from a fairly large number of
neobrane and system equipment manufacturers. Many of these firms are listed
.12
in references previously cited ' and in the Chemical Engineering Equipment
Buyers' Guide which is published annually by MeGraw Hill.
Application-
Reverse osmosis appears to have found widespread acceptance in the
electroplating industry for the recovery of metals in rinsewater. A list of
current reverse osmosis installations in the electroplating industry, adapted
from Reference 24, is shown in Table 6.4.6. This list is expected to increase
as new membranes are developed to meet the demands of the harsh electroplating
environment.
Environmental Impacts--
The reverse osmosis technology, as applied in the electroplating
industry, is a recovery technology capable of achieving zero discharge in the
certain applications. There should be no detrimental environmental impacts
associated with this technology provided the reject stream can be recycled.
Advantages and Limitations—
When properly applied, reverse osmosis systems should achieve economic
benefits associated with chemical recovery and the elimination of the expense
of hazardous waste disposal. The disadvantages of reverse osmosis are
associated largely with the limited lifetimes of membranes in some
applications, resulting in cost penalties for membrane replacement and
pretreatment. Most suppliers will favor conducting treatability studies Co
ensure successful application of their systems to a specific waste stream.
6-42
-------
TABLE 6,4.6, CURRENT RO INSTALLATIONS IN THE ELECTROPLATING INDUSTRY
Type of bath
Type of membrane
and configuration
No. of installations/
zero discharge
Bright nickel
Nickel sulfsmate
Watts nickel
Copper sulfate
Zinc sulfate
Brass cyanide
Copper cyanide
Hexavalent chromium
Cellulose acetate
Spiral wound
Polyamide
Cellulose triacetate
Thin-film composite
Hollow-fiber
Spiral wound
Thin-film composite
Spiral wound
Polyamide
Cellulose triacetate
Hollow-fiber
Polyamide
Hollow-fiber
Thin—film composite
Spiral wound
150/yes
12/no
1/901 recovery
5/902 recovery
2/90% recovery
Under investigation
Source: Reference 24,
6-43
-------
6.5 ELECTSODIALYSIS
6.5.1 Process Description
Electrodialysis is one of the more recent technologies applied to the
recovery of plating chemicals from rinse solution. Electrodialysis uses an
electric field as the driving force to remove charged ionic species from a
feed stream. Anion and cation exchange membranes allow anions and cations,
respectively, to pass from the feed stream to a concentrated ionic solution.
As noted previously, several types of cation and anion exchange membranes
are available commercially. Some properties of commercially produced
membranes were shown previously in Table 6.2,1. As with reverse osmosis
systems, electrodialysis systems are available as packaged units equipped with
electrical components, pumps, motors, pretreatment features, recycle,
temperature control, cleaning, and other features. These can be arranged in
parallel or series as required by the application and its process streams.
Properly designed and operated, electrodialysis unite have proven to be
25
effective and reliable.
By packaging several cell pairs of membranes (typically 50 to 300 cell
20
pairs ) between electrodes and manifolding the streams, a concentrated
stream and a depleted stream, from which 45 to 55 percent of the ions have
2
been removed, are generated. Further ion reduction of the depleted stream
can be accomplished in additional stages. However, electrodialysis cannot
process highly deionized water because of the poor electrical conductivity of
such waters. A flow sheet of a three stage electrodialysis system used for
desalting is shown in Figure 6.5.1. As noted in Reference 2, the feed and
recycle pumps operate at pressures of about 50 psig, a pressure sufficient to
supply as many as six stages without intermediate pumps. Excessive feed
pressures must be avoided to prevent leakage. However, the flow rate in any
stage must be sufficient to create adequate turbulence to keep concentration
polarization below scaling limits.
Electrodialysis removes dissolved matter from water, leaving nonionized
material (such as many organics, suspended matter, silica, etc.) present in
the ion depleted water. This can cause problems if, for example, a build-up
6-44
-------
__ I.DOQmjd 4,500 Dpi*
7-0 ) 0429
V_y "•*
CONCENTPtATE
; f ~\ i ,000 mjd
' Q J9.S44 ppm
STAGE
1
2.02S ssm
STAGE
2
311 OP<™
STAGE
3
1.000 mad SlOfiWL
_ TO WASTE _
0.4Z9 mjd
t4,C34 ppffl
CONCiNTSATE RBCVCH
0,571 mja 14,034 ppm
FEED 1.429 mgd 4,500opm
Figure 6.5.1. Schematic of a three-stage ED plant.
Source: Reference 2.
6-45
-------
of organic compounds in the purified stream IB undesirable. Also, all ionic
species are noneelectively recovered* Pretreatment (lor example, to reduce
concentrations of .hardness components and other dissolved impurities or
organics) and maintenance requirements must consider these possibilities and
their implications,
A potential problem with all applications is the possibility of reaching
excessive current densities because of the high concentration of ions at the
membrane interfaces. Possible consequences of this are the precipitation of
metals euch as calcium and magnesium and the electrolysis of water to hydrogen
and hydroxide ions. Undesirable effects leading to membrane fouling and local
overheating of membranes can result.
Pretreatment or system design features can avoid problems resulting from
electrolysis. For example, introducing turbulence or reducing the total ionic
content of the concentrate stream have been successful in reducing fouling and
electrolysis. To avoid fouling tendencies, almost all manufacturers recommend
periodic reversal of the applied voltage while simultaneously re-routing the
feed and concentrate.
There are no fundamental limits, other than solubility, on the maximum
concentration level obtained in. the concentrate. However, power consumption
is directly proportional to the ion content of the feed. This contrasts with
reverse osmosis, in which separation costs are less strongly influenced by
concentration. Consequently, e-lectrodialysie operating costs are favorable
for low feed ion concentration and become less so as concentration increases.
In addition, electrodialysis, according to Reference 21, is generally used to
produce a concentrated solution, such that evaporation units are not
required. Where a valuable concentrate is being provided, salts may be
concentrated to 20 percent or more, significantly beyond that feasible for
12
reverse osmosis systems.
6.5.2 Process Performance
As noted in Reference 25, there are now more than 100 applications of
electrodialysis to process rinsewacers from electroplating processes. At
least three vendors are currently sar.ufac taring systems for treatment of
wastes from gold, chromium, silver, and zinc cyanide plating operations and
from nickel plating operations. Other plating baths treated successfully by
electrodialysis include tin and tin lead fluoborate, and trivalent chromium
6-46
-------
baths. Application to hexavalent chromium plating is questionable because of
the potential for degrading presently available membranes. Another
application that has been successfully demonstrated involves the recovery of
chromic acid and sulfuric acid from spent brass etchants. This particular
electrodialysis system, used for acid recovery, was developed at the Bureau of
9 A "? 7
Mines * and is now available from Scientific Control, Inc. in Chicago,
Illinois. The system is applicable to wastes containing copper as the primary
20
contaminant. A more detailed description of this process and other
eiectrodialysis processes can be found in References 10, 20, and 25, and in
other primary references cited.
6.5.3 Cost of Treatment
Typical costs for electrodialysis systems to treat plating rinsewaters
range from £30,000 to £45,000, depending on the application. Capital
costs for the "Chrome Napper" system available from Innova Technology, Inc. in
Clearwater, Florida, range from $9,900 to $30,000, including installation and
power supply. This has been successfully used for the recovery of chromic
acid from electroplating wastewaters. Systems are sized according to bath
temperature, dragout concentrations, number of rinse tanks, concentration of
the bath, and the volume of spent solution to be treated per unit time.
Scientific Control, Inc. sells electrolytic electrodialysis units to
recover chromic/sulfuric acid brass etchants. Unit sizes are based on the
amount of copper the system is capable of removing per unit of time.
Available unit sizes range from 0.05 to 0.5 Ib copper removal per hour.
Capital equipment costs (1986 dollars) for these units range from $24,000 to
$80,000. These costs do not include1 installation which would include a hoist,
plumbing, and a ventilation/exhaust system. Additional costs for the exhaust
system could range from $5,000 to $15,000, depending on tfhe size required.
Operating and maintenance costs are relatively low. Membranes will need to be
replaced approximately every 9 months, depending on usage, at a replacement
cost of approximately 10 to 15 percent of the original equipment costs.
Additional maintenance costs will include approximately ilO/month for
replacement of filter cartridges la pre-filter system is incorporated into the
unit). The estimated payback period for the system is approximately 2 years,
29
based on savings in treatment and disposal costs.
6-47
-------
6.5.4 Overall Status of Process
Availability—
Over 100 electrodialysis systems are now employed commercially for the
25
recovery of metals from electroplating rinsewaters. At least three
manufacturers (Scientific Control, Inc.,•Chicago, IL; Innova Technology,
Clearwater, FL; and Ionics Inc., Watertown, MA) are currently manufacturing
electrodialysis equipment for this application. Other membrane-oriented
1—5 12 20
equipment suppliers are listed in the references " * , as well as in the
Chemical Engineering Equipment Buyers' Guide.
Application—
The principal area of application of electrodialysis appears to be the
recovery of metals from electroplating bath rinsewaters. Electrodialysie and
reverse osmosis are competitive processes for these applications.
Electrodialysis would appear to have the advantage when concentration levels
are low (operating costs are low), or wh'en recovery values justify the expense
of achieving concentration levels higher than those possible with reverse
osmosis.
Environmental Impacts—
Because electrodialysis is a recovery process, the environmental impacts
are limited to those resulting from pretreatment and post-treatment.
Fretreatment operations generating wastes include filtration to remove solids,
oil, and grease, chemical precipitation to remove scaling components, and ion
exchange to remove organics that could lead to biological fouling or
electroplating difficulties. Post-treatment requirements are minimal, but
could involve treatment or purification of process streams suffering from a
gradual accumulation of contaminants in a near zero discharge system.
Advantages and Limitations—
A significant advantage of electrodialysis over reverse osmosis is its
ability to concentrate solutions up to their solubility limit, thus avoiding
6-48
-------
the need for auxiliary equipment such as evaporators. Further advantages
include the following, as noted in Reference 2:
• The units operate continuously (ion exchange without regeneration);
» The only utility required for operation is a DC power source;
• The units are compact; and
• Operating cost is low; electrical power consumption averages
$0.25/hour.
Disadvantages of the process vary, depending on the application. All
ionic species are nonselectively recovered, including ionic bath impurities.
Conversely, organic brighteners, wetting agents, and other nonionized
compounds will accumulate in the dilute stream, limiting its reuse potential.
A potential problem with any application is the possibility of exceeding
the maximum voltage set by the solution conductivity at the membrane boundary
layer. The consequence of this condition is electrolysis of water to hydrogen
and hydroxide ions and the possible resulting precipitation of metal
hydroxides which will foul the membranes.
6.6 OTHER MEMBRANE PROCESSES .
Previous reference'has been made to other membrane technologies, notably
Donnan dialysis and coupled transport. Although these processes offer
potential advantages over other technologies, including the commercially
available membrane technologies, they have yet to achieve commercial status.
Recent discussions with representatives of firms involved in the study and
development of these processes have indicated that no additional work is in
progress. ' However, the Bend Research Corporation, a developer of a
coupled transport process, is actively seeking licensing arrangements for a
process they feel is viable and demonstrated.
6-49
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REFERENCES
1. M. Cheryan. Ultrafiltration Handbook. Technomic Publishing Company,
Inc., Lancaster, PA. 1986.
2. G. Belfort. Synthetic Membrane Processes. Academic Press Inc., Orlando,
FL, .1984.
3. R.E. Resting. Synthetic Polymeric Membranes. 2nd Edition, John Wiley &
Sons, Inc., New York, NY. 1985.
4. D.R. Lloyd. Material Science of Synthetic Membranes. Americai Chemical
Society, Washington, B.C. 1985.
5. Sourirajan, S., and T. Matsuura. Reverse Osmosie/Ultrafiltration Process
Principles. National Research Council of Canada, Ottawa, Canada. 1985.
6. J.D. Birkett. Dialysis/Electrodialysis. In: Unit Operations for
Treatment of Hazardous Industrial Wastes. Noyes Data Corporation, Park
Ridge, NJ. 1978.
7. J.B. Berkowitz. Unit Operations for Treatment of Hazardous Industrial
Wastes. Noyes Data Corporation, Park Ridge, NJ. 1978.
8. E.I. DuPont DeNemours & Company. Perroaset Permeators, Waste Treatment
by Reverse Osmosis. October 1984.
9. L.E. Applegate. Membrane Separation Processes. Chemical Engineering.
June 11, 1984.
10. I.E. Biggins. Industrial Processes to Reduce Generation o£ Hazardous
Waste at BOD Facilities, Phase- II Report. July 1985.
11. C.H. Gooding. Reverse Osmosis and Ultrafiltration Solute Separation
Problems. Chemical Engineering. January 7, 1985.
12, The Hazardous Waste Consultant-. May-June 1985.
13. H*K. Lonsdale. The Growth of Membrane Technology! Journal of Membrane
Science, 10:81. 1982.
14. H.F. Hamil. Removal of Toxic Metals in Electroplating Wash Water by a
Donnan Dialysis Process. U.S. EPA-600/2-82-098. December 1982.
15. C.J. Fournier. Heratek Corporation, Billerica, MA. Unpublished Material
and Bulletins supplied to Alliance. April 1987.
15. Kemtek Corporation, Bulletin No, PB-01. 1986.
17. T.V. Tram. Advanced Membrane Filtration Process Treats Industrial
Wastewater Efficiently. Chemical Engineering Progress, March 1985.
6-50
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18. Donnelly, R.G., Goldsmith, R.L., McNulty, K.J., and M. Tan. Reverse
Osmosis Treatment of Electroplating Wastes. Plating. May 1974,
19. Crampton, P., and R. Wilmoth. Reverse Osmosis in the Metal Finishing
Industry. Metal Finishing. March 1982.
20. Wilk, L. et al. Technical Reeource Document - Treatment Technologies for
Corrosive-Containing Wastes. Prepared for U.S. EPA, HWERL, Cincinnati,
OH. October 1986.
21. Higgins, T.E., CH2MHILL. Industrial Processes to Reduce Generation of
Hazardous Waste at DQD Facilities - Phase 2 Report, Evaluation of 18 Case
Studies. Prepared for the DOD Environmental Leadership Project and the
U.S. Army Corps of Engineers. July 1985.
22. Finkenbiner, K., and P. Cartwright. Reverse Osmosis Technology for Toxic
Heavy Metals Treatment. Presented at HAZPRO '85, Baltimore, MB. May 12,
1985.
23. J.D. Birkett. Reverse Osmosis. In: Unit Operations for Treatment of
Hazardous Industrial Wastes, Noyea Data Corporation, Park Ridge, NJ.
1978.
24. P.S. Cartwright. An Opdate on Reverse Osmosis for Metal Finishing.
Plating and Surface Finishing. April 1984.
25. G.C. Cushnie, Centec Corporation, Reston, VA. Navy Electroplating
Pollution Control Technology Assessment Manual. Final Report prepared
for the Naval Facilities Engineering Command under Contract No.
F08635-81-C-Q258. NCEL-CR-84-019. February 1984,
26. Soboroff, D.M., Troyer, J.D., and-A.A. Cochran. Regeneration and
Recycling of Waste Chromic Acid-Sulfuric Acid Etchants, Bureau of Mines
Report of Investigations No. 8377. 1979.
27. McDonald, H.O., and L.C. George. Recovery of Chromium From
Surface-Finishing Wastes. Bureau of Mines Report of Investigations
No. 8760. 1983.
28. D. Pouli, Innova Technology, Clearwater, FL. Telephone conversation with
Lisa Wilk, Alliance Technologies Corporation. August 26, 1986.
29. S. Gary, Scientific Control, Inc., Chicago, IL. Telephone conversation
with Lisa Wilk, Alliance Technologies Corporation. August 29, 1986.
30. J.B. Hsu, Southwest Research Institute, San Antonio, TX. Telephone
conversation with Alliance Technologies Corporation. April 1987.
31. D. Friesen. Bend Research Corporation, Bend, OR. Telephone conversation
with Alliance Technologies Corporation. April 1987.
6-51
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SECTION 7.0
LIQUID-LIQUID EXTRACTION
7.1 BACKGROUND
Liquid-liquid extraction involves the separation of s component from a
waste solution by transfer to a second liquid. The extractant is immiscible
in the waste, but exhibits a preferential affinity for the constituent.
Although not a widely applied treatment technology, liquid extraction has
potential for removal of many toxic constituents from wastewaters. Liquid
extraction is particularly attractive in cases where the solutes are present
at high enough concentration levels to provide recovery value or when other
treatment methods are lees effective.
In the mining industry, the solvent extraction of metal salts from
aqueous solutions has acquired commercial importance, particularly for the
recovery of copper, nickel, cobalt, uranium, vanadium, and other metals from
aqueous effluents. However, the application of solvent extraction to treat
waste effluents remains undeveloped. As a unit operation, extraction lags in
terms of the amount of research which has been conducted and the availability
of quantitative design methods. Performance is highly waste and site-specific
due to competing reactions, desired selectivity in multi-component wastes, and
operating limitations of pH, temperature, and processing time. The impact of
these variables can only be accurately determined through laboratory scale
testing, using various extractants or combinations of extractant, to determine
distribution isotherms and reaction kinetics. Thus, much of the technological
development of extraction has been carried out by manufacturers of specialized
equipment and has not- been disseminated through open literature.
The simplest extraction system is comprised of three components: 1) the
solute, or material to be extracted; 2) the carrier, or the noniolute portion
of the feed mixture to be separated; and (3) the solvent, which is immiscible
7-1
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with the carrier phase. Discussions of extraction require distinctions to be
made between the light and heavy phases, the dispersed and continuous phases,
and the raffinate and extract phases. The terminal streams from an extractor
are the extract and the raffinate. This is shown in Figure 7.1.1 for the case
of countercurrent extraction.
As a recovery process, the use of solvent extraction involves the
following steps:
1. Extraction—Constituents are transferred from aqueous phase to
organic phase using an organic solvent as an extractant.
2. Back-Extraction/Stripping—The constituent to be recovered is
transferred from the organic phase to a concentrated aqueous phase.
Solvents used for the extraction of metals include three basic types:
cation or acidic extractants; anion exchangers; and solvating agents, as shown
in Table 7.1.1. Metal cations react with the cation or acidic extractant,
typically an organic acid, to form neutral complexes that are preferentially
dissolved by the organic phase. The following equation describes a cation
exchange:
h0* * nRH ^ MRn + nH*
As shown, hydrogen ions are exchanged for the metal cation in proportion to
its valence. Thus, Fe is preferentially extracted by acid extractants in
the presence of divalent ions such as Cu or Ni . The degree of
extraction of the metal will also increase with the pH of the aqueous phase
since, at low pH, the extractant cannot release its hydrogen ion in exchange
for the cation.
As shown in Table 7.L.I, anion exchangers in solvent extraction are
generally protonated forms of primary, secondary, and tertiary high-molecular
weight amines and quaternary compounds. The extraction of metal complexes
proceeds primarily by either ion exchange or an addition reaction as
follows:
R4N+X- + MYn+l ^
or
7-2
-------
Extract
5 + CKA)
* •
Peed
Solvent
^S^&:
£2r*&X
S^g-5-.-.;.
1%^!"
•H&S&&
•it**.'.—,'
"&?.?'&?,
X'V^Vi-V-'
f ;4 *. Carrter>7';-57?r,"l^
S • J. r Sol vent;;- ^g?*^.
i.C"»Solutt (distributed;
Ig^jA^y^^.^.. .J. . . ., . «^u_^4aii^
Raffinite
XK+C+fil
»
^v)^
Solvent
5 .
Raffinate
4 (+e+o
Extract
Fetti
Light solvent
Heavy solvent
Figure 7.1.1. Simple system illustrating extraction technology.
Source: Reference 1.
7-3
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TABLE 7.1.1. CLASSIFICATION OF EXTRACTION REAGENTS
Extractant type Extraction mechanism
Example extractants
Cation or acid
extractants
Anion exchangers
Solvating agents
Extraction by compound
formation
Extraction by ion-pair
formation
Organic acids, such as csrboxylic,
sulfonic, phosphoric, phosphonic,
phospbinic acids; and acidic
chelating agents.
PolyphonyImetalloid type,
polyalkyIsulfonium type, polyalkyl-
ammonium type, and salts of high—
molecular-weight aliphatic amines.
Extraction by solvation Carbon-, sulfur-, or phosphorus-
bonded oxygen-bearing extractants;
alkylsulfides; etc.
Source: Reference 2.
7-4
-------
In any reaction, both extraction mechanisms occur. However, the controlling
mechanism is determined by the free concentration of X in the aqueous phase
and the dominant species present in the aqueous phase (i.e., MX ,
^(n-2)^ ^ ^ m ^ MX~ , etc.). Thus, a knowledge of individual
system chemistry is of prime importance in determining reaction kinetics and
optimal system design.
Extraction by solvation requires the transfer of a formally neutral
species from the aqueous to the organic phaee. This occurs through solvation
of the metal ion of a neutral salt species or, in the case of formation of a
2
complex acid species, eolvation of the proton:
or
HMVn + xS ^=2
In acidic and solvation extraction of metal complexes, the extraction agent
will replace primary and/or secondary waters of hydration, thus rendering, the
complex soluble in the organic phase.
The degree of extraction of a metal by a solvating extractant depends on
a number of factors, including: 1) the nature and concentration of the
anionic coordinating ligand X which, in turn, influences the type of metal
complex formed; 2) the degree of hydration of these aqueous metal complexes;
and 3) the relative strength of the water-metal and extractant-metal bonds.
These factors determine the nature of the competition between water and tbe
extractant for the solvation sites. Ideally, the metal ion will be completely
2
stripped of its hydration layer,
Tables 7.1.2 through 7.1.4 summarize the structures and properties of
frequently used cation and acidic extraetants, anion exchangers, and solvating
exchangers.
7.2 PROCESS DESCRIPTION
Solvent extraction can be used for the recovery of concentrated solutions
or for treatment of wastewater streams prior to discharge. The former is used
for the treatment of spent process solutions sucb as used pickling and plating
baths, and process bleed streams. The purpose of recovery in this case is to
7-5
-------
TABLE 7.1.2. STRUCTURE AND PROPERTIES OF ACIDIC EXTRACTANTS
Name
Formula
As Received Extractant
Molecular Weight Active Flash
of Active Extractar.t, Specific Point,
Extractant wt, % . Gravity C
Veisatic 10
I — C — COOH
175
99.6
0.91 129
Rj +• R2 t R3 = Ca
Di-2-ethylhexyl- (C4H9CH(C2Hs}CHjO},POOH
phosphoric acid
Octylphenyl- ROPO(OH). t (RO)jPOOH
phosphoric acid
__
K -
SYNEX105!
R-
S03H
^U"
^_A^
R = C^ n i §
322
458
100
50
0.98
0.92
Source: Reference 2.
7-6
-------
TABLE 7.1.3. STRUCTURE AND PROPERTIES OF ANION EXCHANGERS
As Received Extrsctant
Name
Formula
Molecular Weight Active
of Active ' Extractant,
Extractint • wt. %
Specific
Gravity
Flash
Point,
°C
Primary Amines (RNHj )
Primene JMT
LA-2
Adogen 283
Adogen 364
Alamine 336
Adogsn 368
Hostarex A327
Adogen 381
Alamine 308
Hostarex A324
Adogen 382
Adogen 464
Aliquat 336
R-(CH3)3C(CH5C88
0.84
' 0.83
0.83
0.81
0.81
0.82
0.81
0.82
—
0,81
0.82 '
0.84
0,88
-
180
-
—
168
—
203
—
-
166
-
—
132
Source: Reference 2,
7-7
-------
TABLE 7.1.4. STRUCTURE AND PROPERTIES OF SOLVATING EXTRACTANTS
Name
Ethsrs (RjORi) or (R2OC
Diisopropyl ether
DibutylceUosolve
Alcohols (ROH)
n-Butanol
n-Pentanol
Kstones (RjCORj)
Methyl isobutyl ketone
Molecular
Formula Weight
Carbon-Oxygen-Bonded Donors
RI =(CH3)jCH • 102
R2=C,H9 .174
R = C4H9 74
R = CSH,, 88
RI = CH3> R2 = (CH3}SCHCH2 100
Specific
Gravity
0,726
0.837
0.81
0.82
0.804
Flash
Point,
°C
-25
_
32
33
14
Viscosity
(2S°C),
cP
0.38
1.34
,2.46
3.31
0,55
R, = R2 = Rs = C4H,O
Phosphoric acid esters
Tri-n-butylphosphate
Phosphonic acid esters
HostarexPQ2J2
HostarexPO224
Phosphine oxides
Trioctylphosphine oxide R, = R7 = R3 = C8H,,
266
R, = Ra = C^HsO, R3 = C4H9 2SO
R, = R, = C8Kj,O, R3 = C6Hn 418
386
Sulfides(RSR)
Dihexyl suifide
Sulfur-Containing Exrracianis
= C6H,3 202
0.97
0.94
0.91
193
(solid)
3,56
5
16
Source: Reference 2.
-------
purify or recycle the solution by removing or reducing the concentration of
undesirable impurities. Process costs oust be balanced against the value of
recovered solution and, in some cases, the recovery value of the impurity.
For example, build-up of iron and "copper in pickling and etching operations,
respectively, can be partly or entirely extracted. The regenerated solution
can be reused after makeup and the recovered metal sold to secondary metal
smelters.
Solvent extraction can also be used to treat low concentration liquid
effluents. Typical applications often involve large volume flows, such as
rinse waters used to remove dragout from pickling and plating baths. Mine
waters, wet scrubber solutions, and drainage waters from dumps are also
examples of dilute effluents which might utilize extraction.
System designs and configurations are highly waste and site specific.
Configurations can include multiple stage extraction, use of more than one
solvent for sequential extraction to obtain higher purity separations, and
pretreatment (e.g., precipitation, filtration) and post-treatment
(e.g., carbon adsorption to remove organics from the raffinate) options.
These considerations are discussed below.
In commercial applications, the extractant is typically used at a 10 to
40 percent active level in a non-toxic, inexpensive solvent such as
kerosene, Depending on the application, this solution can be used in
volumes equal to that of the waet« stream. It is contacted countercurrently
with the waste feed, usually at a slightly elevated temperature, to improve
4 5
exchange kinetics and improve phase separation. ' Initial solvent
selection will be based on its extraction efficiency and selectivity as
determined through laboratory testing. However, solvent purchase price,
anticipated loss (e.g., resulting from incomplete phase separation), and
breakdown will determine the overall economic viability of the extractant.
Breakdown can occur as a result of build-up of organic additives which are
present in the wastewater. These will be carried over into the organic phase
and will eventually hinder phase separation. Alternatively, breakdown can
occur due to long-term incompatibility with regenerants.
Solvent selection and processing conditions will determine overall system
efficiency. Reaction kinetics can be enhanced by increasing the concentration
of the extractant, raising the waste pH, increasing"the relative volume of
7-9
-------
extractant to feed, and improving mixing. However, since each of these
impairs phase separation, an optimal balance must be determined. Phase
separation can also be enhanced by using solvent modifiers (e.g., decyl
alcohol, tributyl phosphate), optimal solvents (e.g., psraffinic solvents give
better separation than kerosene or high aromatics), and reagent combinations.
As an example of the latter, secondary reagents used with suifonic acids can
improve separation rates by 2 or 3 times over sulfonic acids alone.
Choice of regeneratit is dependant on the speed and efficiency of the
4 '
separation. A typical regenerant will contain 10 to 15 percent acid at an
organic/aqueous volume ratio of 15 to 1. Volume ratios of the original
waste feed and the final metal-laden regenerant depend on the overall ease and
required completeness of separation. Increases in metal concentrations of 20
to 30 times the feed level are not uncommon.
The most prevalent type of equipment used in commercial applications is a
mixer-settler which consists of a high intensity mixer and & large baffled
settling chamber. Host applications use several units in series to provide
high volumetric throughput. Equipment and maintenance costs are low, but
expenses can escalate rapidly if significant solvent make-up or extensive
post-treatment are required. Post-treatment equipment may include
electrolytic or evaporative recovery units for the concentrated metal stream
and carbon adsorption systems for organic removal from the raffinate. Other,
%",v
more advanced, equipment includes reciprocating plate column contactors,
centrifugal contactors, and electrostatic coalescers. Beaker teats have shown
up to 90 percent reductions in phase separation time with the latter.
A large number of processes using solvent extraction have been proposed
for the treatment of liquid waste effluents. Some of these have included
pilot or laboratory studies, as summarized in Section 7.4. However, very few
proposed applications have actually been carried through to commercial
operation. Those which have been identified are summarized below.
7-10
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7.2.1 lecovery of Zinc from Pickling Liquor
Pickling in the galvanizing industry is commonly performed with
hydrochloric acid. The spent solution contains about 1 to 2 percent free HC1
and an iron content of 100 to 130 g/L. The solution will also contain 20 to
120 g/L of zinc, as well as smaller concentrations of California List metals
2
such as chromium and nickel. The high zinc content prohibits conventional
treatment of the liquor via thermal decomposition to iron oxide and
hydrochloric acid.
The Metsep process was designed to separate zinc from the iron chloride
2. 8
solution by continuous resin ion exchange. * Anionic zinc-chloro complexes
are absorbed on a strong-base, ion-exchange resin. The resin is eluted with
water to yield zinc chloride which is converted to a sulfate medium by solvent
extraction with Di-2-ethylhexylphosphoric acid (D2EHPA) and stripped with
eulfuric acid. The product is a zinc sulfate solution suitable for
electrowinning (an electrolytic deposition process used for the recovery of
metals from solution) and subsequent re-sale (e.g., sale to an electrolytic
zinc refinery). The solvent extraction raffinate is used for hydrochloric
acid production.
The MeS Process, an alternative to the Metsep process, has been developed
for the recovery of zinc from pickle liquors. It uses a solvent extraction
circuit for the initial separation of zinc from iron in the pickle liquor.
Zinc is preferentially extracted, as a zinc chloride complex, with
Tri-butylphosphate (TBP). Iron extraction is minimal, since it is primarily
in the ferrous state. The preferential extraction of zinc over iron is
somewhat less with TBP than with optimal amine extractants, but this is
9
balanced by operational advantages such as higher loading.
Zinc is stripped from the organic solution with water or dilute sulfuric
acid. The zinc chloride strip solution is mixed with sulfuric acid mother
liquor in a boiler, thus evaporating hydrochloric acid and crystallizing zinc
sulfate. The zinc sulfate is separated by centrifugation. By adjusting the
conditions in the boiler, chloride-free zinc sulfate suitable for
electrowinning can be produced. The HC1 and iron oxide by-products are also
recovered. HC1 is returned to the process and iron oxide is available as &
saleable product although exact purities were not specified.
7-11
-------
The MeS procesa was developed by MX-Processor AB in Sweden and has been
9
piloted in Holland at a galvanizing plant with encouraging results. The
raffinate produced contained less than 100 ppot zinc. Extractant residues were
removed by activated carbon adsorption.
7.2.2 Recovery of Zine~Cyanide Plating Batha
Zinc electroplating is carried out front alkaline zinc cyanide solutions
which generate contaminated rinsewater reauiring treatment for cyanide
destruction. Zinc cyanide can be efficiently extracted from alkaline
solutions by quaternary amines. The Union Carbide Corporation * has
developed a process based on the simultaneous extraction of both zinc and ,
cyanide. The decontaminated raffinate was recycled as fresh rinse water. The
atnine extractant was regenerated by stripping with sodium hydroxide and
recovered zinc and cyanide were recycled as plating bath make-up.
A typical composition of the contaminated rinse water was 40 ppm cyanide
and 23 ppm zinc. Solvent extraction reduced these values to 0.4 ppm cyanide
and 0.07 ppm zinc. Active carbon treatment further reduced these levels and,
at the same time, reduced entrained and dissolved amine to 0.1 ppm. A ratio
of feed to strip solution of 162:1 produced a strip solution containing 3 to
4 g/L zinc. The same procedure has also been demonstrated for cadmium cyanide
plating rinse waters, however, it is not known whether the process has been
operated commercially.
7.2.3 Recovery of Copper Plating and Etchant Baths
In general, the higher purchase price of copper relative to zinc and the
simplicity of electrolytically recovering it from extractants makes its
recovery comparatively more favorable. The hydrometallurgical copper industry
has widely applied selective extractants that have a high affinity for copper
in weakly acidic and ammoniacal solutions, while simultaneously rejecting
ferric iron. Another application includes the recovery of copper containing
etchants as described below.
Etching of copper, with the use of an ammoniacal solution, is a common
procedure in the manufacture of printed circuit boards for the electronics
7-12-
-------
industry. Spent anmioniacal etching solution contains free ammonia, one OT
more anraoniacal salts, copper, and oxidants. Maximum etching efficiency is
obtained when the ammoniac a 1 solution contains 110 to 130 g/L of copper. It
gradually diminishes as the copper concentration approaches 150 to
12 13
170 g/L* * Thus, to keep etching efficiency constant and optimal, the
etching solution must be continuously regenerated or replaced with fresh
solution.
14
A process patented by the Criterion Corporation- (see Figure 7.2,1)
completely removes copper from spent etchants to produce a fresh product that,
after makeup, can be re-sold to printed circuit board producers. The process
uses an (LIX64N) solvent to selectively extract copper and chloride ions;
process of this type has been operated in the United Kingdom by Proteus
Reclamation Ltd., recovering 300 kg/day of copper with Acorga P5100 used as
extractant.
A similar process which has been used for onsite etchant recovery is the
Mercer Process (see Figure 7.2.2). This process withdraws etchant
directly from the etching line and recirculstes it through a solvent
extraction circuit, maintaining the copper concentration in the etchant within
the optimal range of 110 to 130 g/L. Treatment of the rinsewater obtained
from rinsing the circuit boards after etching is also integrated into the
process.
In the first extraction stage the etchant is mixed with an organic
solution containing LIX54 it) kerosene. The initial copper concentration in
the etchant, approximately 130 g/L, is reduced to 90 g/L. The regenerated
etchant is returned to the etching line after careful removal of entrained
organic solvent. The copper in the rinsewater is extracted in a second
extraction stage. At the same time, any entrained etchant from the first
extraction stage is washed out. Copper is stripped from the organic solvent
with barren copper electrolyte in the stripping stage. The solvent is reused
and copper metal is produced by electrowinning on titanium cathodes.
The Mercer process is in successful operation in two prototype
installations and commercial units are new -marketed by P.R. Processutveckling
AB» Sweden. Etchant makeup has reportedly been- reduced by 95 percent and no
negative influence on product quality has been encountered.
7-13
-------
Regenerated
etchant
Spent etchant
I
__j
CuSQ*
Cu cathodes
Figure 7.2.1. Process for recovery of copper from spent
ammoniacal chloride etchant.
Source; Reference 14.
7-14
-------
Etching
Washing
Extraction 1
Extraction 2
Solvent extraction
r
Strip
~ H70
HjSOi
Electro-
winning
i
I
L_J
CuS04
Cu cathodes
Figure 7.2.2. The MECER-process for on-line regeneration
and copper recovery from ammonlacal etchant.
Source: Reference 10.
7-15
-------
7.2.4 Recovery of Nickel Plating Baths
Recovery of nickel from plating baths and rinsewaters is another feasible
application of solvent extraction. One process proposed by Plett and
Pearson is based on extraction of nickel with D2EHPA. The solvent is used
in its sodium form to avoid pH changes during extraction, as follows:
2NaD2EHP(org) «• Ni2* ^ * Ni(D2EHP)2orgj + 2Na*
A typical feed solution may contain 1 to 2 g/L nickel. Laboratory tests
showed that nickel could be effectively removed in two extraction steps with
2
the resulting raffinate containing 4 mg/L. By loading the solvent with
nickel, the transfer of sodium to the strip solution was minimized. Nickel
was stripped from the solvent with dilute sulfuric acid and recovered from the
strip solution by electrowinning,
7.2.5 Recovery of Chromium Plating Baths
During the plating of chromium, a buildup of impurities such as Fe(III),
Cr(III), Ni, Cu, and Zn gradually takes place, making the bath unusable for
plating. Rinses containing similar contaminants in more dilute concentration
are also generated. Two alternative solvent extraction procedures are
available for recovery of these wastes: extraction of Cr(VI), or extraction
of impurities.
Chromate and dichromate ion extraction from acid solutions by use of TBP
or an amine extractant have been well documented. Cuer et al. have
investigated the applicability of the extractants Alamine 336, LA-2, and TBP.
They report excellent stability for TBP and have designed a process for the
recovery of 99.5 percent of chromium (VI) from combined industrial effluents.
The particular cases described in the reference refer to recovery of waste
liquors originating from the production of chromium anhydride (CrO-,) and
recovery from processes using this product (e.g., chromium plating and metal
treatment). The extracted chromic acid is recovered as sodium chromate by
stripping with a sodium hydroxide solution. Concentrations of 200 g/L of
7-16
-------
Cr(VI) in the atrip solution are reached. A similar process is reported in a
Japanese patent, which includes the use of a chromic acid wash of the organic
solvent to remove extracted impurities such as iron and chloride.
2
The approach taken by MX Processor Afl, Sweden to recover spent
chromium plating baths is to extract the impurities, thus regenerating the
solution for reuse. A flow sheet of this process is outlined in
Figure 7.2.3. The extraction ie carried out by use of a mixture of HDNNS and
TBP. Since the acidity of the plating bath limits the extraction efficiency,
dilution of the bath with water significantly improves the operation. Water
balance is partly maintained by the natural evaporation during plating. Small
amounts of chromic acid may be extracted, but can be selectively scrubbed with
water which can be used for dilution of the plating bath before extraction.
Five molar H.SO, or HC1 is used for stripping, giving a metal
concentration in the strip solution of more than 60 g/L«
7.2.6 Removal of Mercury from Chior-Alkali Effluents
The mercury associated with brine effluents from chlor-alkali plants must
be reduced from approximately 10 ppm to the parts-per-billion level prior to
discharge. A solvent extraction method has been proposed by Gronier, which is
based on the rapid extraction of mercury from chloride medium by
18
high-molecular-weight tertiary and quaternary amines. With a mercury
contamination of 10 ppm in the effluent, a concentration factor of 2,500 is
claimed to be obtainable, leaving a strip solution containing 25 g/L mercury.
The amount of mercury that is extractable in a particular case depends, to
some extent, on the pH of the brine. With tertiary amines, better than
99 percent is achieved at pH 3, but there is a decrease at higher pH. The
most significant problem with this system is the loss of organic material to
the brine effluent.
A closely related process, based on amine solvent extraction, has been
19
patented by Chapman and Caban. It is not known whether any commercial
applications have been implemented.
7-17
-------
H,0
| Cr(lll)
Me sulfaie
Figure 7.2.3. Process for removal of impurities and
regeneration of chromium plating bath.
Source: Reference 2.
7-18
-------
7.2.7 Miscellaneous Applications
Research on the recovery of noble metals, primarily from plating
solutions, has been reported. Gold and silver are extracted from cyanide
solutions with quaternary amines, ' Gold can be stripped from the
organic solvent with alkaline potassium cyanide solutions.
22
Rothmann et el. have proposed the use of an aroine extractant for
recovery of chromium and vanadium in effluents from the processing of
.eteelmaking slags. However, difficulties were reported with the precipitation
of silica, which interfered with phase separation.
7,3 PRSTREATMENT AND POST-TREATMENT REQUIREMENTS
Pretreatment requirements for liquid-liquid extraction are not nearly as
stringent as those required for other physical separation techniques;
e.g., carbon adsorption or membrane systems. Since the equipment used in
extraction is not very .susceptible to fouling or plugging, suspended solids
removal is less critical. However, optimal operation will frequently require
removal of organics which will otherwise dissolve in the extractant and
interfere with subsequent operations; e.g., phase separation. Pretreatment
operations can also be applied to prepare the waste for optimal separation
efficiency. Pretreatnente include pH adjustment, dilution, flow equalization,
and temperature increases to optimize reaction and phase separation time.
Post-treatment will be required for the raffinate to remove residual
metals/cyanides and solvent which has either dissolved in or been entrained in
the coalescer effluent. Treatment options typically consist of adsorption or
possibly biological treatment for organic destruction. Concentrated metal
solutions from the regeneration step will be removed from the acidic media by
electrowinning if recovery is economically viable. Alternatively, if
concentrations are too dilute, they may be increased using evaporation or a
membrane technology which is capable of operating in an acidic environment.
Finally, if economic recovery is not viable, removal through precipitation may
be most feasible.
7-19
-------
7.4 PERFORMANCE DATA FOE LIQUID-LIQUID EXTRACTION
Performance data for solvent extraction is limited. Much of the
development of processes and equipment has been carried out by manufacturers
of customized equipment and is, therefore, considered to be proprietary. In
other cases, pilot studies often do not supply enough information to indicate
whether the corresponding commercial-scale process might be feasible. Data
which were identified in this study are presented below.
7.4,1 Copper and Nickel Extraction from Metal Finishing Sludge
A pilot-scale study, conducted by the Department of Metallurgical
Engineering of the Indian Institute of Technology, explored the use of LIX64N
(10 percent by volume) to extract copper and nickel from metal finishing
23
wastewater and sludge. Samples were generated by redissolving metals from
sludge with sulfuric acid to produce a solution containing 4.45 g/l copper,
0.16 g/L nickel and 2.5 g/L zinc. In addition, a synthetic sample containing
4.3 g/L copper, 7.9 g/L nickel, and 2.5 g/L zinc was made by dissolving metal
sulfates in water. The extractant used was a 10 percent solution of LIX6AN in
kerosene and the stripper consisted of 4 N sulfuric acid. Extractions were
carried out in a 500 uL separating funnel by manual shaking.
Table 7.4.1 illustrates the effect of pH on the extraction of copper with
LIX64N from the synthetic solution. Multi-stage extractions were subsequently
conducted at a pH of 2.0 as shown in Table 7.4.2. Tables 7.4.3 and 7,4.4 show
similar diagrams for nickel. Finally, Table 7.4.5 is a. summary of results for
experiments conducted on the sludge leach solution. The data suggest that
this process may be economical since the solvent is fully recovered.
7.4.2 Metal Extraction from Metal Finishing Wastewater
A study conducted for the EPA by Curtis W. McDonald, Texas Southern
University, explored the use of high-molecular-weight amines for the removal
of toxic metals from metal,finishing wastewater such as cadmium, chromium,
24
copper, nickel, and zinc. The researchers used a 25 percent Alamine 336
solution diluted in xylene, with an extractant/water ratio of 1 to 100. The
high ratio is desirable in concentrating the metal and in avoiding emulsions.
7-20
-------
TABLE 7.4.1. EFFECT OF pH ON SINGLE-STAGE EXTRACTION OF
COPPER USING 10 PERCENT VOL/VOL LIX64N
PH
1.3
1.8
2.0
2,2
2.5
Copper concentration
in strip solution
Cg/L)
0.89
1.78
2.28
2.22
2.04
Distribution
coefficient
(D)
0.26
0.70
1.12
1.06
0.89
% Extraction
20.07
41.05
53.20
51.80
47.00
Composition of leach solution: Cu 4.3 g/L, Ki 7.9 g/L,
and Zn 2.5 g/L; vol. of ao. phase 100 mL; vol. of
organic phase 100 mL.
Source: Reference 23.
TABLE 7.4.2. MULTISTAGE CO-CURRENT EXTRACTIONS OF
COPPER BY LIX64N
No. of stages
employed
3
4
5
% Extraction % Extraction
calculated actual
89.5 73.6
95.0 80.6
97.6 .93.3
Feed composition: Cu 4.3 g/L, Ni 7.9 g/L, and
Zn 2.5 g/L; vol. of aq. phase 100 tnL; vol. of
org. phase 100 niL; pH 2,0.
Source: Reference 23.
7-21
-------
TABLE 7.4.3. EFFECT OF pH. ON SINGLE-STAGE EXTRACTION OF NICKEL
Concentration of
nickel in strip
pH solution
8.0 4.4
8.2 5.3
8.5 5.9
9.0 4.3
9.5 4.0
Distribution
coefficient
(D) 1 Extraction
1.26 55.7
2.02 67.0
2.96 74.8
1.18 54.3
l.Oi 50.4
Composition of feed solution: Cu 0.3 g/L, Ni 7.9 g/L,
and Zn 2.5 g/L; vol. of aq. phase 100 mL; vol. of
organic phase 100 mL.
Source: Reference 23.
TABLE 7.4.4. MULTISTAGE .CO-CURRENT EXTRACTION OF NICKEL
No. of stages
employed
2
3
% Extraction
calculated
95.0
98.4
1 Extraction
actual
94.41
97.04
Feed composition: 0.3 g/L copper, Si 7.9 g/L, and
Zn 2.5 g/L; pH 8.5, vol. of aq. phase 100 mL; vol.
of org. phase 100 mL.
Source: Reference 23.
7-22
-------
TABLE 7.4.5. MULTISTAGE CO-CURRENT EXTRACTIONS OF Cu AMD Hi
CONTAINED IN THE LEACH SOLUTION OBTAINED PROM
TEE LEACHING OF HYDROXIDE SLUDGE
Metal
Copper
Nickel
pH
2.0
• 8.5
No. of stages
employed
4
2
£ Extraction
actual
93.97
96.25
Original feed composition: 0.16 g/L nickel; 4,4 g/L
copper, arid 2.4 g/L zinc; vol. of aq. phase 100 mL;
vol. of org. phase 100 mL.
Source: Reference 23.
7-23
-------
Chromium extraction was found to be affected by the quantity of
hexavalent versus trivalent chromium, since the latter is not as easily
extracted. However, baaed on the effect of chloride concentration (HC1) on
extraction, as shown in Table 7.4.6, selective extraction appears to have
potential application. Results of a simultaneous extraction of the same three
metals is shown in Table 7.4.7. A dose of 33 mL of concentrated HC1 was mixed
with 1-liter of the wastewater prior to extraction. For both of the
experiments, no appreciable amount of copper or nickel was extracted.
Stripping of the loaded extract was performed with 4.0 M NaOH with more than
99 percent of the metals being stripped. As previously indicated in
Table 7.4.3, metals extraction is pH dependent and the addition of HC1 or NaOH
reagents affect extraction efficiency. The solvent was reused 15 times with
no loss of efficiency. 'Reagent loss was estimated by an increase in TOC
content of 50 ppm in the aqueous phase/extraction.
7.4.3. Lab _Scale_ Study Using Sequential Extractions
A lab scale study was performed by Clevenger and Novak on a simulated
4
regenerate waste from an electroplating ion exchange unit. Four chelating
compounds dissolved in chloroform were studied for recovery of Fe, Zn, Cu, Hi,
and Cr. Results are shown in Figures 7.4.1 through 7.4.4 for single-stage
extractions. It was shown that high metal removal efficiencies could be
achieved with pronounced selectivity for copper at low pH for two of the
chelates. Using these selectivity data, the researchers experimented with
various schemes to identify optimal sequential extractions.
Although nearly complete removal and high selectivity could be achieved
with sequential extractions, the investigators realized that recovery of both
the metal and the chelator would be necessary for the process to be
economically viable. Metals were efficiently extracted with 2.4 M HCi and
0.75 M HNO- solutions. However, the chelators could not successfully be
reused due to significant loss o£ extraction capabilities following acid
recovery. Since the chelators alone are generally more expensive than the
recovery value of the metals, this process would not be economically viable on
a commercial scale.
7-24
-------
TABLE 7.4.6. SELECTIVE EXTRACTION OF CHROMIUM, CADMIUM, AND ZINC
Chromium Cadmium Zinc
Mean % extd. Std. dev. Mean % extd. Std. dev. Mean % extd. Std. dev.
First extraction 88.6 4.0 0.0 0.0 0.0 0.0
0.002M chloride
Second extraction 0.65 1.09 94.1 1.65 8.1 2.1
0.03M chloride
Third extraction 0.55 0.87 4.8 1.3 80.9 3.5
0.4M chloride
Total metal 89.8 98.9 89.0
extracted
Composition of wastewaters: Cr - 10.0 to 56.8 ppm
Cd - 4.1 to 5.9 ppm
Zn - 5.2 to 9.2 ppm
Cu - 0.3 to 0.5 ppm
Ni - 0.4 to 0.5 ppm
Source: Reference 24.
-------
TABLE 7.4.7. SIMULTANEOUS EXTRACTION OF CHROMIUM, CADMIUM, AND ZINC
Chromium Cadmium Zinc
Mean % extd. Scd. dev. Mean % exCd. Std. dev. Mean % extd. Std. dev.
First extraction 90.6 1.2
Second extraction 0.0 6.0
Total metal 90.6
extracted
98.0 0.11 83.3 0.11
1.4 0.86 15.4 1.8
99.4 98.7
i;
cr Composition of wastewaters: Cr - 8.0 to 9.0 ppm
Cd - 3.7 to 4.0 ppm
Zn - 4.8 to 5.2 ppm
Cu - 0.3 to 0.5 ppm
Ni - 0.4 to 0.5 ppm
Source: Reference 24.
-------
100
90
:? eo
§30
c
S 20
123156789 IO
PH
Figure 7.4.1. Metal extraction efficiency
as a function of pH using
0.1-M thenoyltrifluoracetone
and chloroform.
Source: Reference 24.
100;
90
80
5 TO
g so
b,
t 50
| -O
§ 30
tt
^ ?o
LU
10
1 2 345676 9 10
pH
Figun; 7.4.2. Metal extraction efficiency
as a function of pH using
0.1-M acetylacetone and
chloroform.
Source.: Reference 24.
3
£
j-,
u
-
u
u.
u
1
^
cr
tu
100 '•
90-
80 -
70 -
60 -
50 •
40-
30
?oJ
1O-
Cu _ .. ... . „-• • •
ca - ' A--''-'
N, ,- /
/ »
/ m!
t S
f *^-m~- — '" *^
Cr *-*
f
i
i
I
Zn '
,23456789 10
pH
Figure 7.4.3-
Metal extraction efficiency
as a function of pH using
0.03-M sodium diethyidithio-
carhamate and chloroform.
Source: Reference 24.
1OO
90 -
80 -
70
SO
50
4O
3O
20
1O
9 6
PM
Figure 7.4.4. Metal extraction efficiency
as a function of pH using
0.1-M 8-hydroxyquinoline
and chloroform.
Source: Reference 24.
-------
7.4.4. Metal Recovery From Scrap
An example of nickel, colbalt, and iron recovery from metal scrap
., lathe turnings, mill shavings) wi
process involved five processing steps:
(e.g., lathe turnings, mill shavings) was provided in Che literature. The
1. Pyroraetallurgical treatment to convert Mo and W into their carbides;
2. Electrolytic dissolution of Fe, Co, and Ni followed by partial
stripping to concentrate the CaCl2 electrolyte;
3. Separation of Fe, Co, and Mi by extraction with a high molecular
weight amine;
k. Stripping the Co/Fe organic extract with the weakly acidic
condensate from step (3); and
5. Cathodic deposition of Co and Ni in separate half-cells. The
stripped electrolyte then goes to step (2).
The process produces Ni/Pe and Co/Fe mixtures, which reportedly does not
significantly affect the market value of the nickel or cobalt. The process,
depicted schematically in Figure 7.4.5, does not generate any liquid discharge.
7.5 STATUS AND COSTS OF EXTRACTION
As stated previously, extraction is not a widely applied technology for
the treatment of metal/cyanide wastes. The design and effectiveness of an
extraction system will be highly specific to the waste type, constituent
concentrations, and waste quantity. The difficulty in identifying an
appropriate system and the relatively complicated nature of the process itself
has undoubtably hindered its acceptance, particularly among smaller waste
generators. Instead, its widest application has been in larger firms which
have installed custom-designed systems through extensive support from the
equipment and reagent suppliers.
Extraction will probably only be used in situations where recovery of
valuable constituents or, recovery of baths via removal of contaminants,
cannot be achieved tnrougn more conventional means. Due to the dependency of
design variables on site—specific factors, generalizations on equipment
7-28
-------
TREATED METAL SCRAP
CATHODE
CAC12
• CATHODE
KEL
r
SOL'N
^
CATHQDIC
HALF
CELL
' ^
•s
SOL'N
f
ANODIC
HALF
CELL
tOBAL
r
CATHODIC
HALF
CELL
Ni
ELECTROLYTE
(CAC12)
SOLID
FILTER I
CARBIDE
DILUTE
CO,FE,NI
CHLORIDE
SOLUTION
COBALT/I RON
ELECTROLYTE
EVAPORATOR
CONDENSATE
CONC,
SOL'N,
Cr=250 G/L
AMINE
EXTRACTION
AMINE/METAL 4-
COMPLEX
(CAC12)
COBALT/IRON
STRIPPING
Ni, CA
CHLORIDE SOL'N
Tppmaay i
AMINE
CONDENSATE
CONDENSATE
Figure 7.4,5. Schematic of metal recovery from scrap.
Source; Reference 5.
7-29
-------
selection cannot be made. Similarly, the limited coat data presented in the
literature cannot be generalized, since it is highly dependant on reagent loss
and recovery value of the metals or batba. For example, solvent reagents cost
roughly £l.50/pound, which can represent a significant operating cost if
appreciable losses occur.
7-30
-------
REFERENCES
1, Humphrey, J.L., Rocha, J.A., and J,R. Fair, The Essentials of
Extraction. Chemical Engineering, pp. 76-95. September 17, 1984.
2. Lo, T.C., Malcolm, H.I», -and C. Hanson. Handbook of Solvent Extraction.
John Wiley & Sons. pp. 53-89. 1983.
3. King Industries Product Bulletin SY2-378. August 18, 1978.
4. Clevenger, T.E., and J.T. Novak. Recovery of Metals from Electroplating
Hastes Using Liquid-Liquid Extraction. Journal of the Water Pollution
Control Federation, 55(7):984-989. July 1983.
5. L.V* Gallacher. Liquid Ion Exchange in Metal Recovery and Recycling.
American Electroplaters Society Conference. April 14, 1980.
6. King Industries. Using Syvex Liquid Ion-Exchange Reagents. Product
Bulletin.
7. Haines, A.K., Tunley, T.H., TeRiele, W.A.M,, Cloete, F.L.D., and
T.D. Sampson. J. South Afr. Inst, Min. Met. 74, 149. 1973.
8. Tunley, T.H., Rohler, P., and T.D. Sampson. J. South Afr. Inst. Min.
Met. 77,423. 1976. R. Wood. Process Eng. 6, 6. 1977.
9. H. Reinhardt. Some Hydrotaetallurgical Processes for the Reclamation of
Metal Waste. Paper presented at IWTU Conference, Waterloo. 1978.
10. Moore, F.L., and W.S. Groemier. Plating Surface Finishing, 26. August
1976.
11. P.L. Moore. Separation Science 19(4)489. 1975.
12. W. Hunter. The Use of Solvent Extraction for Purification of Silver
Nitrate Electrolyte. Paper presented at the 2nd. International
Conference of Precious Metals, New York, NY. May 1978,
13. W. Hunter. Electrolytic Refining. U.S. Patent 3,975,244. 1976.
14. Hamby, W.D., and M.D. Slade. Process for Regenerating and for Recovering
Metallic Copper from Chloride-Containing Etching Solutions. U.S. Pattent
4,083,758. 1978.
15. Flett, D.S., and D. Pearson, Chem. Ind. 639. 1975.
16. Cuer, J.P., Stukeos, W., and N. Texier. Proceedings of International
Solvent Extraction Conference (1SEC), Lyons 1974, Vol. 2, Society of
Chemical Industry, London. p. 1185. 1974.
7-31
-------
17. Nishimura, S., and M. Watanabe. Japan Kokai 76, 90,999. 1976.
IS. Reinhardt, H., and H.D. Ottertun. Swedish Patent Application 76-13686-0.
19, W.S. Gronier, Application of Modern Solvent Extraction Techniques to the
Removal of Trace Quantities of Toxic Substances from Industrial
Effluents. Oak Ridge National Laboratories, TN. Report ORNL-TM-4209.
1973.
20, Chapman, T.W., and R. Caban. Extraction of Mercuric Chloride from Dilute
Solution and Recovery. U.S. Patent 3,899,570. 1975.
21. Ivanoviakii, M.D., Meretukov, M.A., Potekhin, V.D., and L.S. Strizhko,
Izvest. Vysah. Ucheb. Zadev. Tsvetnye Metally 17 (2), 36. 1974.
22. Rothmann, H., Bauer, G., Stuhr, A., and H.J. Retelsdorf. Metall.
(Berlin) 30 (8), 737. 1976.
23. Phule, P.P., Dixit, S., and R. Mallikarjunan. The Recovery of Principal
Metal Values from Metal Finishing Hydroxide Sludges. Indian Journal of
Technology. Vol. 23, 462-464. December 1985.
24. C.W. McDonald. Removal of Toxic Metals from Metal Finishing Wastewater
by Solvent Extraction. Industrial Environmental Research Laboratory,
Texas State University. 1977.
25. King Industries. Price Schedule, Effective October 1, 1986.
7-32
-------
SECTION 8.0
ADSORPTION FOR METAL REMOVAL
8.1 CARBON ADSORPTION
Adsorption involves the interphase accumulation or concentration of
substances at an. interface. The process can occur between any two phases,
such as liquid-liquid or liquid-solid interfaces. The material being
concentrated or adsorbed is the adsorbate, and the' adsorbing phase is termed
the adsorbent.
Activated carbon adsorption involves separation of a substance from one
phase, typically an aqueous solution, and the concentration of the substance
at the surface of an activated carbon adsorbate. Activated carbon is most
widely used for the removal of organic1 contaminants and is most, effective when
the organic solutes have a high molecular weight and low water solubility,
2
polarity, and degree of ionization. However, studies in the field of
metallurgy have indicated that carbon adsorption of many metallic compounds
can be successfully achieved and has found commercial application for certain
aqueous waste-streams. ' However, adsorption efficiency varies
considerably between different compounds.
' Activated carbon is available as a powder (PAC) .or in the form of
granules (GAG). GAC is most commonly used because its larger size is most
amenable to handling in conventional contacting and regenerating
equipment. However, despite handling and regeneration problems, PAC is
preferred in some treatment schemes; e.g., when used in combination with
biological treatment. * Both types of carbon have 'effective surface areas
far in excess of their nominal external surface areas. Surface arems are on
the order of 500 to 1,500 'square meters per gram,.resulting primarily from a
network of internal pores 20 to 100 angstroms in diameter. Porosities can be
D
as large as 80 percent. The characteristics of the micropore structure are
8-1
-------
largely dependent on the activation process, which is a controlled sequence of
dehydration, carbonization, and oxidation of rav materials including coal,
wood, peat, shell, bone, and petroleum based residues.
The equilibrium capacity of an activated carbon for a contaminant is a
function of the effective carbon surface area and the surface binding
process. Adsorption equilibria are governed by two types of interactions:
solute-adsorbent, which describes the carbon's affinity for the solute
(contaminant), and solute-solvent, which involves the solubility of the solute
in the liquid media. In general, an inverse relationship between the extent
of adsorption of a substance from a solvent and its solubility in that solvent
9
can be anticipated. Overall, the relative affinity of a solute for either
phase will be determined by the lyophobic (i.e., solvent-disliking)
characteristics of the solute and the affinity of the solute for the adsorbate.
Activated carbon adsorption of inorganic compounds is more complex and
compound specific than adsorption of organic compounds, primarily due to the
charged nature of inorganic species in aqueous media. The important
physical-chemical properties of activated carbon selected for inorganic
electrolyte adsorption are: specific surface area, pore structure, and surface
chemistry of the adsorbent. *
Specific surface area may be defined as that portion of the total surface
area that is available for adsorption. Specific surface area is proportional
to adsorption which, in turn, is dependent on pore size and pore size
distribution. The pore size can range from leas than 20 angstroms and size
distribution is dependent on the source materials and the activation process
employed. Available surface area is nonpolar in nature, but depending
upon the activation process, active sites can be formed yielding a slightly
polar surface.
•In addition to specific area, the adsorption capacity of activated carbon
is primarily influenced by surface chemistry; e.g., the formation of
carbon-oxygen complexes at the carbon surface and the anionic adsorption
capacity. ' Formation of surface functional groups is dependent on the
activation process and carbon source. Two broad categories of activated
carbon can be identified based on activation temperature and atmospheric
12
conditions: L-type, which tends to adsorb bases, and B-type, which tends to
adsorb acids*
8-2
-------
L-type carbons are prepared by exposure to oxygen at 300°C to 400°C or by
solution oxidation. H-type carbons are prepared by outgassing at 800°C to
1000°C followed by cooling in an inert atmosphere and exposure to oxygen at
12
room temperature. Typical surface oxide functional groups formed by these
methods include:
* carboxyl,
* phenolic bydroxyl,
• lactone and quinone,
» carboxylic acid,
* anhydrides, and
* cyclic peroxides.
The surface oxide groups have significant effects on the adsorption
capacity of activated carbon. Electrokinetic studies have shown that H-type
activated carbon exhibits a positive surface potential whereas L-type carbons
exhibit a negative surface potential. This is due partly to the high pH which
developes when H-type carbon is brought in contact with water and the low pH
which occurs when the surface functional groups of L-type carbon are
hydrated. * The surface charge characteristics of both carbon types can
be readily'modified by the introduction of a strong base or acid to the
10 • :
system.
In general, the following system parameters have the greatest influence
on metals removal by activated carbon:
• pH, •
» metal concentration,
• activated carbon dose,
• ionic strength,
* temperature, and
* presence of ligands.
8-3
-------
The dominant solution parameter controlling adsorption of inorganic
13
chemicals is pH. As mentioned above, pH has a controlling influence on
the surface charge characteristics of the adsorbent. The distribution of
metal ions in solution is also a function of pH, with lower pH favoring
14 • '
solvation of metal ions. The overall pH effect on adaoption o£ metal ions
can be attributed to electrostatic attraction, which is a function of the
charge of both the solid adsorbent and the adsorbate.
Studies have shown that the removal efficiency of inorganics from a waste
stream by activated carbon increases with concentration of either the
+ 2
adsorbant or the solute. For example, researchers demonstrated that Cd ,
Cr , Cr , and CN all show improved removal efficiencies at higher
initial concentration throughout the pH range tested. * Similarly! by
increasing the activated carbon/Cd ratio by 100 tines, a threefold increase in
removal efficiency was realizedt
Adsorption of Cd(II) was demonstrated to decrease with increasing ionic
c 14
strength of the solvent. This suggests that the extent o£ adsorption is
sensitive to changes in concentration of supporting electrolyte, indicating
that electrostatic interaction may be a significant conponent of adsorption in
plating and other metal containing solutions.
Since adsorption is an exothermic process, adsorption efficiency might be
expected to improve with decreasing temperature. However, it has been
+2 14
demonstrated that adsorption of Cd increases with temperature. This
is unexpected thernodynatnically and suggests that some extrinsic process which
responds to temperature increase is at work.
Complexation of metal ions by inorganic and organic ligands can
dramatically increase or decrease adsorption compared to a ligand-free
13
system. Studies have demonstrated that mercury and cadmium show a
significant increase in adsorption efficiency using chelating agents such as
APDC, TETA,-NTA, and EDTA. "
System pH also influences adsorption of metal-ligand complexes. The
effect of pU on cadmium removal was determined in a system using EDTA,
tartate, citrate and TR1EN as completing ligands. Adsorption increased
with pH for 3 out of 4 ligands tested. The other, EDTA, showed improvements
with initial increases in pH up to pH 6, but adsorption capacity fell off as.
higher pE was approached.
8-4
-------
The relative feasibility of using different adsorbents is usually
determined in the laboratory by developing adsorption isotherms. The
adsorption isotherm is a functional expression for the variation of adsorption
with concentration of adsorbate in bulk solution at constant temperature.
Typically, the amount of adsorbed material per unit weight of adsorbent
increases non-linearly with increasing concentration. It must be noted that
adsorption isotherms can vary widely for different carbons, and isotherm data
cannot be used interchangeably.
The two most common isotherm expressions used are the Freundlich Equation
and the Langtnuir Equation. The Freundlich equation in an empirical expression
but is often useful as a means for data description." The Langmuir model,
originally developed for adsorption of gases onto solids, is predicated on
three assumptions: (l)adsorption energy is constant and independent ol surface
coverage; (2)adsorption occurs on localized sites with no interaction between
adsorbate molecules; and (3)maximum adsorption occurs when the surface is
covered by a monolayer of adsorbate.
The Freundlich equation can be expressed as follows:
where: x = mass of adsorbate, mg
m = mass of dry adsorbent, g
k - constant, indicative of adsorption capacity
C = equilibrium solution concentration, tng/1
1/n = constant, indicative of adsorption intensity.
Data for the Freundlich equation are usually fitted to the logarithmic form of
the equation:
log (-} - log k + (-) log C
m n
This expression is a straight- line with a slope of L/n and an intercept equal
to the log k when C • 1 (.log 1 » 0).
8-5
-------
The Freundlieh equation generally shows good agreement with both the
Langmuir equation and experimental data, over moderate ranges of
concentration, C.
In its linearized form, the Langnmir equation can be expressed as follows:
£ - i- «. i c
x ab a
m
where: x = mass of adsorbate, rag
m * mass of dry adsorbent, g
C = equilibrium solution concentration, mg/1
a = solid phase concentration corresponding to complete
coverage of available adsorption sites (mass solute
adsorbed/mass carbon; for complete monolayer)
b • constant related to the enthalpy of adsorption
The coefficients of the Langmuir equation can be calculated by performing a
linear regression of Che data or determined graphically, by plotting G/(x/m)
veraua C on arithmetic graph paper (slope = I/a, intercept = 1/ab).
The adsorption data from both models are useful in estimating the
relative effectiveness of adsorbents for a given application. However, care
must be exercised in assessing 'performance when the wastestream contains a
large number of competing adsorbates. Most users will be forced to rely on
laboratory scale adsorption isotherm results and prior industrial experience
to assess performance and appropriate system design for a specific wastestream.
8,1,1 Process Description
Although activated carbon adsorption has been shown to effectively treat
some metal and cyanide containing waste streams, it is generally employed in
the treatment of organic containing wastes. Consequently, the focus on
8-6
-------
applied technology and research has been on the treatment of organics.
However, research has been conducted on unique systems used for the treatment
of metal and cyanide containing wastes, as discussed later in this section.
A schematic of an activated carbon adsorption system for the treatment of
hexavalent chromium is shown in Figure 8.1,1, After exhaustion of the
adsorption capacity, the activated carbon is regenerated with sulfuric acid,
The regeneration acid is pH adjusted to precipitate chromium, and the sludge
is removed via filtration. Carbon ie returned to the process. There is
typically an accompanying loss in adsorption capacity as a result of a small
but significant depletion in effective surface area. This can-result from a
build up of hard to remove adsorbate, attrition, and other mechanisms.
8.1,1.1 Pretreattnent/Post-Treatment Requirements—
Pretreatoent of the feed to carbon adsorption columns is often required
to improve performance and/or prevent operational problems. As discussed in
Reference 9, there are four primary pretreatments which may be required:
• equalization of flow and concentrations of primary waste
constituents;
• filtration;
' * adjustment of pH; and
• adjustment of temperature.
Generally, the flow to the adsorber columns and the concentration of the
primary waste constituent are not constant. Since variations in either can
have a detrimental impact on system performance, it is necessary to make
provisions to equalize flow and minimize concentration surges. Flow
equalization is accomplished by employing a surge ta.nk prior to the column.
Concentration equalization is also accomplished somewhat by employing surge
tanks, however, this may have to be supplemented by mechanical agitation.
Mixing prevents concentration surges which can lead to premature column
leakage and breakthrough or conversely, low concentration swings resulting in
premature regeneration of an underloaded adsorber column.
8-7
-------
OTHER WASTE
STREAM
GEL/Cr +6
STREAM
CARBON
ADSORPTION
BEDS
00
00
CLEAN WATER
FILTRATE
REGENERATION
EQUIPMENT
SLUDGE
HANDLING
FILTER
SUMP
SOLIDS TO
DISPOSAL
pH
ADJUSTMENT
TO MUNICIPAL
SANITARY SEWER
Figure 8.1.1. Regenerative carbon adsorption system.
-------
It is a general requirement that Che feed to column be low in suspended
solids. It is difficult to set an upper limit on the absolute level of
acceptable suspended solids because the physical nature of the solids is as
important as their concentration. For example, finely divided, siity solids
tend to pass through the bed, but coarse material of varying particle size can
rapidly form a mat on top of the bed, thereby constricting flow. In general,
if the column feed is turbid or the suspended solids level is greater tnan
10 mg/L, pretreatment for solids removal will be required.
In addition to suspended solids, many waste contaminants can interfere
with carbon adsorption. For example, if calcium or magnesium are present in
concentrations greater than 500 ng/L, they may precipitate and plug or foul
the column. Oil and grease in excess of as little as 10 mg/L has been
18
reported to interfere with column operation. The presence of many other
compounds can influence adsorption of the contaminant of concern through
competition for available adsorption sites.
Removal of suspended solids and other waste contaminants noted above may
be achieved by pretreatment with multi-media pressure filters. Such filters
complement fixed bed adsorption processes and can be readily integrated into a
total design. Other filtration options include membrane filtration, when a
highly clarified feed is desired, and uitrafiltration, if high molecular
weight contaminants are encountered (over 1,000).
Activated 'carbon adsorption systems for metals are sensitive to changes
in pH, particularly when the contaminants to be removed are either weakly
acidic or weakly basic. Control can be easily achieved by installing pH
measurement and acid/base reagent addition systems in the surge tank to
maintain the desired pH feed to the adsorption columns. Finally, provisions
for feed heating nay be required since adsorption of metals has been shown to
vary with temperature, as discussed above.
Under proper design and operating conditions, the treated water will
generally be suitable for discharge to surface waters. Other aqueous streams
such as backwash, carbon wash and transport waters are recycled or sent to a
settling basin. Acidic regenerants, which are typically used for metals, may
be treated through neutralization and precipitation.
8-9
-------
8.1.1.2 Operating Parameters—
Optimal process design ot both the adsorption and regeneration, or
desorptian, systems is dependent on the waste's physical, chemical, and flow
characteristics. Isotherms, determined in a laboratory, measure the affinity
of activated carbon for the "target" adsorbates in the process liquid. This
provides data for determining the type and amount of carbon which will be
required to treat the full scale process stream. Carbon requirements will be
based on a limiting constituent for which attainment of acceptable effluent
concentration is the most difficult.
Table 8.1.1 gives properties of some commercially -available granulated
19
activated carbons. Properties of a typical powdered activated carbon are
20
shown in Table 9.1.2. Adsorption properties of the two types of carbon
are generally comparable, the principal difference being the particle size.
The fine size of the PAG makes it unsuitable for use in the contacting and
regeneration equipment used for adsorber applications but makes it ideal for
flow through processes (e.g., biological treatment) equipped with filtration
systems for carbon removal.
A typical continuous adsorption system consists of multiple columns
filled with activated carbon and arranged in either parallel or series. Total
carbon depth of the system must accommodate the "adsorption wavefront";
i.e., the carbon depth must be sufficient to purify a solution to required
specifications after equilibrium has been established. Bed depths of
a
8-40 feet are common. Minimum recommended height-to-diameter ratio of a
column is 2:1. Ratios greater than 2:1 will improve removal efficiency but
result in increased pressure drop for the same flaw rate. Optimum flow rate
must be determined in the laboratory for the specific design and carbon used.
Q
For most applications, 0.5 to 5 gpm per square foot of carbon is common.
Optimal adsorber configuration will be based on influent characteristics,
flow rate, type of carbon, effluent criteria, and economics. Figure 8.1.2
illustrates several arrangements typically used for adsorber systems.
There are two basic modes of operation for columns; namely, fixed beds and
moving or pulsed beds. In the fixed bed mode, the entire bed is removed from
service when the carbon is reactivated. In the moving or pulsed bed, only the
exhausted (inlet) portion of the bed is removed as new adsorbent is
simultaneously added to maintain bed volume.
8-10
-------
TABLE 8.1.1. PEQPEETIES OF SEVERAL CQMMERCIALLY AVAILABLE CARBONS
PHYSICAL PROPERTIES
Surface area, »2/g (BET)
Apparent density, g/ctn^
Density, saekweehed and drained, lb/g^
Real density, g/em3
Particle density, g/em
Effective cize, am
Uniformity coefficient
Pore volume, cnr/g
Mean particle diameter, tm
SPECIFICATIONS •
Sieve size (U.S. std. series)8
Larger than Mo, 8 (max. X)
Larger then No. 12 (max. 2)
Smaller than Ho. 30 (max. 2)
Smaller than No. 40 (max., S)
Iodine No.
AnriiSXDR llQ» HILnl.lIIU115
Ash (Z)
Moisture as packed (max. Si)
IC1
Aoerica
Hytirodarco
(lignite)
600 - 650
0.43
22
2.0
1.4 - 1.5
0.8 - 0.9
1.7
0.95
1.6
8
—
5
—
650
b
b
b
Calgon
Filcrasoris
300
(bituninous)
950 - 1050
0.48
26
2.1
1.3 - 1.4
O.B - 0.9
1,9 or less
0.85
1.5 - 1.7
8
—
5
—
900
70
8
2
Westvaeo
WV-L
(bituminous)
1000
0.48
• 26
2.1
1.4
0.85 - 1.05
1.8 or less
0.85
1.5 - 1.7
8
—
5
__
950
70
7.5
2
Witco
51?
(12x30)
(bituminous)
1050
0.48
30
2,1
0.92
0.89
1.44
0.60
1.2
—
5
5
_.
1000
85
0-5
1
aOtl»er sizes of carbon are available on request from the manufacturers.
^No available data from the.manufacturer.
— Mot applicable to this size cmrbon.
TYPICAL PEOP2RIIES OF 8 X 30-KESH CARBONS
Total surface area, E^/g
Iodine number, min
Bulk density, lib/ft3 bsckwashed and iJiained
Particle density wetted in water, g/en^
Pore volume, cm^/g
Effective sits, asm
Uniformity coefficient
Mean particle dia. , am
Pittsburgh abrasion mmber
Moisture as packed, max.
Molasses RE (Relative efficiency)
Ash
Mean-pore radius
Lignite
carbon
600 - &50
5QO
22
1.3 - 1,4
1.0
0.75 - 0.90
1.9 or less
1.5
50 - 60
9Z
100 - 120
12 - 182
33 A
Bituoinous
coal caiboa
950 - 1,050
950
26
1.3 - 1.4
0.85
0.8 - 0.9
1.9 oi less
1.6
70 - SO
2Z
40 - 60
5 - K
14 A
Soyrce: Reference 19.
8-11
-------
TABLE 8.1.2. TYPICAL PROPERTIES OF POWDERED ACTIVATED CARBON
(PETROLEUM BASE)
Surface, Area m2/g(BET) 2,300 - 2,600
Iodine Mo. 2,700 - 3,300
Methylene Blue Adsorption (mg/g) 400 - 600
Phenol Ho. 10-12
local Organic Carbon Index (TQCI) 400 - 800
Pore Distribution (Radius Angstrom) 15 - 60
Average Pore Size (Radius Angstrom) 20 - 30
"5
Cumulative Pore Volume (cm /g) 0.1 - 0.4
Bulk Density (g/cm^) 0.27 - 0.32
Particle Size Passes: 100 mesh (wtZ) 97 - 100
200 mesh (wt%) 93 - 98
325 mesh (wtX) . 85-95
Ash («t%) 1.5
Water Solubles (wtl) 1.0
pH of Carbon 8-9
Source: Reference 20.
8-12
-------
out
r
m
out
UPFLOW IN SERIES
DOWNFLOW IN SERIES
out
in
to..
I
i
I
4.
I
|
w in -
f
",
i
t
t
t
1
out
UP FLOW IN PARALLEL
DOWNFLDW IN PARALLEL
out
MOVING
BED
Figure 8.1.2.
Carbon Bed Configurations,
Source: Reference 21.
8-13
-------
Arrangement of columns in series permits the first column to become
saturated with impurities while a solution of required purity is obtained
through the second, or polishing, column. Upon reaching saturation, the first
column is emptied and refilled with fresh or regenerated carbon. Fluid flow
is redirected to the second column so that the replenished column is now in
the downstream position, resulting in a variation of countercurrent flow
between the waste stream and the adsorbent.
Adsorption beds can be operated in either upflow or downflow mode. A
downflow mode must be used where the adsorber is relied upon to perform the
dual role of adsorption and filtration. Although lower capital costs can be
realized by eliminating the need for pretreatment filters, operating costs
escalate since more efficient and frequent backwashing of the adsorbers is
2
required. Application rates of 2-10 gpm/ft are employed, and backwash
2
rates of 12-20 gpm/ft are required to achieve bed expansions of
o
20-50 percent. The use of a supplemental air scour can be used to increase
efficiency of the backwashing.
While pre-filtration is normally required to prevent blinding of
upflow-expanded beds with solids, smaller particle sizes of adsorber can be
employed to increase adsorption rate and decrease adsorber size. Application
rates can also be increased in the upflow-eicpanded mode, even to the extent
that the adsorbent may be in an expanded condition.
9 .
The design, flow, and configuration arrangements discussed above offer
21
the following advantages and limitations:
Adsorbers in Parallel
Adsorbers, in Series
For high volume applications
Can handle higher than average suspended
solids ( 65-70 ppm) if downflow
Relatively low capital costs
Effluents from several columns blended,
therefore, less suitable where effluent
limitations are low
Large volume systems
Easy to monitor breakthrough at tap
between units
Effluent concentrations relatively low
Can handle higher than average suspended
solids ( 65-70 ppm) if .downflow
Capital costs higher than for parallel
systems
8-14
-------
Moving Bed " - Countercurrent carbon use (most
efficient use of carbon)
- Suspended solids must be low (<10 pptn)
— Best for smaller volume systems
- Capital and operating costs relatively
high
- Can use such beds in parallel or series
Upflow-expanded - Can handle high suspended solids (they
are allowed to pass through)
- High flows in bed (>15
The above systems are not generally used with the much finer powdered
activated carbons. The PAC systems now used involve mixing the PAC with the
waste stream to form a slurry which usually can be separated later by methods
such as filtration or sedimentation. A novel technique where powdered
activated carbon is used to make activated carbon beads, based on a suspension-
polymerization technique, has proven effective for treating some metal
. , 22
containing waste streams.
8.1.2 Experimental Data and Demonstrated Performance
Information gathered from activated carbon manufacturers and industry
indicates that few activated carbon systems are being used specifically for
the treatment of metal and/or cyanide bearing waste streams. Thus, data for
full scale applications are incomplete and essential operating parameters or
pollutant removal characteristics have either not been generated or are
considered to be proprietary information. However, the literature includes a
number of efforts where the feasibility of activated carbon for metal and/or
cyanide removal has been demonstrated. Specifically, some degree of success
has been reported for the adsorption of arsenic, cadmium, chromium, mercury,
and cyanide, as described below.
Arsenic—A number of different adsorbents were tested for their abilities
23
to remove arsenic from a variety of aqueous solutions. However, the
results show that activated carbon was not the best adsorbent tested.
Three types of activated solids were chosen for toe study including,
activated alumina, activated bauxite, and activated- carbon. Experiments were
8-15
-------
carried out with freshwater, seawater diluted ten times, undiluted seawater,
and a 0.67 M sodium chloride solution. In experiments for As adsorption,
all reaction flasks were flushed with nitrogen gas to prevent the oxygenation
of As to As .
The results demonstrate that As is far more adsorbable than As ,
and that As was removed from solution much faster by activated alumina
than by any other adsorbent (see Figure 8.1.3). In general, the rate of
adsorption and extent of arsenic removal decreased with increasing salinity
for all adsorbents tested. The effect of pH on As adsorption by the three
adsorbents was determined by varying pH from 2 to 12. Activated carbon
adsorbs better in the acidic pH range (i.e., between 3 and 5) than at higher
or lower pH values. Alumina and bauxite both displayed adsorption maxima for
pH values between 4 and 7. Even at the pH of maximum adsorption by activated
carbon, alumina and bauxite demonstrated superior performance. Figure 8.1.4
depicts the effect of pH on the performance of activated carbon.
Adsorption equilibrium were described adequately by both the Langmuir amd
Freundlich isotherm models. The effect of solution composition on adsorption
equilibria for activated carbon is shown using the Langmuir model in
Figure 8.1.5. With regard to the ionic strength of the solutions and their
effect on adsorption, it was determined that the rates of adsorption were
slowest in seawater, yet the extent of adsorption was reduced by no more than
23
5 percent relative to freshwater. However, it must be noted that the
isotherm plots are based on only three data points.
Cadmium— In one study, adsorption of cadmium was batch tested using four
brands of activated carbon (GAG and PAC), as shown in Table 6.1.3. Stock
solutions were prepared to represent cyanide and fluoborate plating baths as
follows:
1. cyanide bath: 10~2 M CdO * lO"1 M NaCN (molar ratio of Cd:
CN = 1:10); and
2. fluoborate bath; 10"1 M cd (BF4>2 + 7.0 x 10~2 M NH^
BF, + 5.0 x 10~2 M iUBO, (molar ratios of 'Cd:BF:,NH, :
Hlo - 1:2.7:0.7:0.5).
8-16
-------
18
> 12
Adsorbent
A Aluffiitifi
& Alumina
* Bauxite
O Bauxic*
• Carfesn
o Carbon
Solvent
Water
See Water
Mater
Sea UateT
Hater
Sefi Vacer
Initial As
M H/l
53.4
53,4
53.4
53.4
26.7
26. J
pH
6.5
6.J
6.4
6.7
3.2
3.2
20
60 -
100
600
000
2000
40DO
Figure 8.1.3,
Adsorption of AslV) by different absorbents at
optimum pH values.
Source: Reference 23(
8-17
-------
6 5
TJ
O
JO
o 4
•ft
> 3
4 5
Final pH
Figure 8.1.4.
pH effect an adsorption o£ A.s(V) by activated carbon.
Source; Reference 23.
8-18
-------
300
250 -
£00 -
00
I
150
3000
Figure 8.1.5.
Langmuir isotherms of adsorption of arsenic(V) on, activated
carbon fq.=mM As(V) adsorbed/g of solids, C=mM of As(V)I.
Source: Reference 23.
-------
TABLE 8.1.3. TYPICAL SURFACE PROPERTIES OF ACTIVATED CARBONS
' DSED IN THIS STUDY*
Specific surface area
Carbon type
-------
The ionic strength of the waste water was shown to have a minimal effect
on adsorption, with the rate increasing only slightly with decreasing ionic
strength. However, as expected, the kinetics of cadmium adsorption onto PAG
is faster than that onto GAG since pore diffusion is probably the rate
limiting step for GAG. Conversely, external surface area contributes
significantly to total surface area for PAC, therefore pore diffusion is less
critical. Powdered activated carbons, in particular Nuehar S-N and Nuchar
S-A, had larger cadmium adsorption capacities than the granular forms tested
(Darco HD 3000 and Filtrasorb 400). Figure 8.1.6 sho«s that Nuchar S-A
+2
achieves 90-95 percent Cd removal in the neutral pH range which is
approximately three times that of granular carbons.
Since differences in specific surface area between the diferent carbon
+2
types are not large and the hydrated radius of a Cd ion is estimated to be
much smaller (4 angstroms) than the lower pore size values (10 to 1000
angstroms), the performance difference can be attributed to surface
chemistry. Powdered carbon has a low pH at zero charge (pH ) and
ZPC*
excellent adsorption capacity for cationic metal ions. Granular carbon,
having a high pH , is rather poor for metal ion adsorption. The pH,__
2PC> £ti\j
value reflects the nature of surface functional groups.
The distribution of cadmium species in solution is a function of pH. The
hydrogen ion concentration of the wastewater solution plays a critical role in
the extent of Cd*2 adsorption. For both the fluoborate and cyanide
wastewaters, the adsorption density was found to approach its maximum Level in
the neutral pH range. This is a positive feature of carbon treatment when
compared to alternatives such as precipitation, which requires a pH adjustment
to 10 or 11 for effective removal. Nuchar carbon was found to be particularly
effective in this regard, with little or no pH adjustment required after
14
addition to solution.
Following these experiments, a suspension-polymerization technique was
then used to aggregate one of the PACs, Nuchar S-A, to sizes suitable for
22
column packing. The beaded carbon was compared to a number of other
-activated carbons as listed in Table 8.1.4. All metal solutions were prepared
from reagent graae chemicals. Trie cadmium was a synthetic cadmium
fluoborate, Cd(BF^>2, plating wastewater. Strong acids such as
H SO HC1, and HC10 were used to regenerate the CdCID-laden activated
carbon beads.
8-21
-------
00
NJ
Originai Cd-conc. = 1 x 10"4M
Cd-Bf, Solution
Ionic Strength = 0.1 M as NaCIO
- o o- -
Room Temperature
Reaction Time = 2 hrs.
O Nuchar S-A
O Nuchar S-N
Filtrasorb 400
Darco HD 3000
IO
Figure 8.1.6. Comparison of CdCIl) adsorption capacity by granular
and powdered activated carbon, as affected by pH.
Source: Reference 14.
-------
TABLE 8.1.4. COMPARISON OF ADSORPTION CAPACITY BY VARIOUS TYPES OF
ACTIVATED CARBON8
Carbon
Filtrasorb 100
Filtrasorb 200
Filtrasorb 300
Filtrasorb 400
Darco 12 x 40
Pittsburgh HGR
Darco Granules (HD 3000)
% CdCIl)
Removed
20.5
20.5
20.5
17.0
26.0
5.5
25.0
Carbon
Darco 12 x 20
Darco 20 x 40
Nuchsr 722
Nuchar WV-G
Wuchar WV-L
Nuchar WV-W
Nucbar S-Ab
Nuchar S-N*
% Cd(II)
Removed
26.0
25.5
30.0
23.0
22.0
10.0
83.0
67.0
aBatch adsorption conditions: Cd(BF4>2 10~* M; Carbon 1 g/L;
pH = 7.00; 1 * 0.01 M HeClO^, reaction time - overnight,
bThe only two powdered activated carbons.
Reference 22.
8-23
-------
Preliminary batch studies were performed for m total of fifteen different
types o£ activated carbons. Table 8.1.4 shows the results of the preliminary
runs. The PACs, Nuchar S-A and Nuchar S-N, exhibited a greater Cd(II) removal
capacity than the granular carbons. Column experiments performed with the PAC
beads demonstrated that the superior eiectrophoretic properties of powdered
activated carbon could be combined with the manageability of GAG enabling use
22
of the same contact equipment.
Chromium—One study found that the removal of chromium from solution by
activated carbon occurred through two major interfacial reactions: adsorption
and reduction. This study investigated chemical factors, such as pH and
Cr concentration, that affect the magnitude of Cr adsorption. This
study used a commercial activated carbon, Calgon Filtrosorb 400, in a
continuous mixed batch system.
+6
The adsorption density of Cr increases with increasing pH to a
maximum value and then declines rather rapidly with further increase in pH
(see Figure 8.1.7). When the pH becomes greater than 10, no appreciable
adsorption is observed. The extent of adsorption also increases with Cr
concentrat ion.
Figure 8.1.8 demonstrates that Cr is also removed by reduction to
Cr in the presence of activated carbon. In the absence of activated
carbon, the Cr added remained in the hexavalent state. However, based on
the absence of Cr+3 in the supernatant, the researchers concluded that
reduction only occurs at pH less than 6. This conclusion is valid since
Cr is adsorbed to a lesser extent by activated carbon than Cr .
The results shown in Figure 8.1.9 demonstrate that cyanide is removed by
activated carbon with a maximum value oceuring around pH 8.
Polaroid Corporation, reported the successful application of activated
carbon for the removal of hexavalent chromium from an aqueous waste stream
generated by a slide film production facility. Several alternatives were
considered including ion exchange, electrochemical treatment, sodium
reetabisulfite reduction, ferrous sulfate reduction, and carbon adsorption. A
feasibility study and economic analysis resulted in selection of the activated
carbon system.
8-24
-------
30
25
o
ft
Ui
o
z
o
20
15
a
o 10
i I r~~i i 1 1 1 \ r
10 G/L FILJASORB 400
O.I M NaCI
0.6
0.5
0.4
0.3 o
ui
8
0.2 |
O.I
0.0
10 II
FiRure 8.1.7. The effect of pH and total Cr(Vl) on the adsorption of Cr(Vl).
Source: Reference 6.
-------
Removal of chromium(VI) from dilute aqueous solution
oo
NJ
CARBON VALENCE OF Cr
Figure 8.1.8.
The effect of pH on the state of chromium in the
presence and absence of activated carbon.
Source: Reference 6.
-------
0.8
0.7
0.6
0.5
t-
i-
UJ
0 0,4
I 0.3
to
o
0.2
O.I
"1IIIF
10 G/L CARBON
O.I M NoCi
0.87mMCN
Figure 8.1.9. The effect of pH and total CN on the adsorption of ON,
Source: Reference 6.
8-2?
-------
A schematic of the system is shown in Figure 8.1.10. The system
utilizes carbon for- adsorption of the chromium, which is believed to occur by
reduction with subsequent adsorption of Cr (Note: this mechanism is
different than that postulated in the previous example since different carbon
24
types are used). After exhaustion, the carbon is regenerated by treatment
with sulfuric acid which is then pH adjusted and filtered to remove
24
precipitated chromium.
During the pilot study, the feed pH and pretreatment (filtration) were
found to have a major effect on successful operation of the carbon system.
Adjustment of pH was necessary to extend the Life and capacity of the carbon.
If pH remained above 5.0, bed breakthrough occured 5 to 6 times more quickly
than with adjustment. Pre-filtratioo of the feed was required to prevent
hydraulic fouling of the bed since the film production effluent contained a
gelatinous component that easily plugged the carbon column.
In addition to technical success, Polaroid determined that a carbon
system would also be more economical than the other technologies considered.
Using the carbon on a once-through disposal mode (see Figure 8.1.10), as
opposed to 3 regenerative mode, resulted in the lowest capital and operating
costs in this particular application. A Carbon Service Agreement was selected
as the optimal arrangement. Under this agreement, the carbon manufacturer
leases an adsorber system, supplies carbon, disposes spent carbon, and
provides maintenance support.
Mercury—Experiments were conducted to evaluate the enhancement of .
mercury adsoption that could be realized through preliminary cbelation of
25
mercury ions and chemical treatment of activated carbon. The experiments
included laboratory evaluations of process variables (i.e., pH, chelaee type
and dose, and carbon dose), batch capacity and isotherm tests, and continuous
flow column studies. Synthetic solutions of mercuric chloride, which
simulated wastewaters from chlor-alkali industries, were used throughout the
studies.
Literature reviews performed by the investigators revealed the following:
• chelation of the mercury improves carbon removal capacity,
• reduction of Hg+* to the elemental state may proceed subsequent to
adsorption,
8-28
-------
OTHER WASTE
STREAM
GEL/Cr ^6
STREAM
PH
ADJUSTMENT
FILTER
CARBON ADSORPTION
BED
CLEAN
WATER
00
fo
SUMP
pH
ADJUSTMENT
TO MUNICIPAL
SANITARY SEWER
Figure 8.1.10.
Single usage carbon adsorption system as installed.
Source: Reference 24.
-------
o sulfurizing agents such as C$2 can be used to improve mercury
removal, and
o regeneration of activated carbon laden with adsorbed mercury may be
facilitated at high pH since its adsorption is enhanced by low pH.
Calgon Filtrasorb 300 effectively removed chelated Ammonium pyrrolidine
dithiocarbonate (APDC) mercury from dilute mercuric chloride solutions at both
pH 4 and pH 10, as shown in Figures 8.1.11 and 8.1,12, respectively. Howeverk
at the lower pH, carbon capacities were increased seven-fold in isotherm tests
and nearly 14-fold in column tests. The pronounced pH effect may be explained
in terms of changes in the carbon surface charge. At the mercuric chloride
concentration in these experiments, sufficient chloride ion was available to
complex the mercuric ion to its neutral or negatively charged chloride
complexes. Any soluble metal oxides present at high pH would be converted to
25
mercuric chloride species as the pH decreased. A shift in pH would change
the nature of the carbon surface. This carbon is an H-type carbon and possess
basic surface oxide groups. Since the H—type carbons readily adsorb hydronium
ions, the surface oxides could be neutralized at low pH, allowing pore
diffusion of the mercury-chloride complexes. At higher pHf .the basis surface
groups repel the neutral to negative forms of the complexes.
Carbon disulfide (CS_) greatly increased the removal of mercury by
activated carbon at pH 10 when the carbon was soaked with CS. and dried
prior to adsorption. At mercury concentrations of 1 ppm» CS_ treatment
25
resulted in a 50-fold increase in carbon capacity.
Column tests with granular activated carbon showed improved operation at
low pH. Figure 8.1.13 and 8,1,14 show the breakthrough curves for pB lOand
pH 4, respectively. Relatively poor results were obtained at the high pH;
breakthrough occurred within one day. At pH 4, a substantial improvement in
performance was observed, with no breakthrough evident after 5 days of
25
operation.
In summary, the data indicate that removal of Hg by activated carbon
treatment is feasible. Carbon removals from alkaline wastes may be enhanced
by chelation of the mercury with APDC, treatment of the carbon with CS , or
05 l
by lowering the pH of the wastewater,
8-30
-------
so-
10..
050
for
the lin* y - *xb
•r* •-0.187, b-1.07
APOCDOSC«1».4mg/l
I 5 10
FILTRATE MERCURY CONCENTRATION, ppb
Figure 8.1.11.
Freundlich isotherm demonstrating removal of mercury(II)
by APDC and powdered activated carbon at pH 4 and 20 C.
Source: Reference 25.
8-31
-------
10.0
5.0-
Oil
I?
I
10
OS~
Conttwita for
the line y.axb
, b-.81O
APOC DOSi =18.4 ma/I
10O 5OO 1000 2000
FILTRATE MERCURY CONCENTRATION, ppb
Figure 8.1.12.
Freundlich isotherm demonstrat Log removal of raercury(II)
by APOC and activated carbon at oH 10 and 20eC,
8-32
-------
9000
1000 ••
600
1
UJ
100
10--
OPERATING CONDmONSi
(Hg)in« 3mg/l
Row Rat** 2.3 gpm/tq tt
Empty B*d Contact Tirtm* fit min
Column H*igW • 107 cm
550 pm ActivM«dOvtanClwg«
0 1 23 4
OPERATING TIME, DAYS
Figure 8.1.13. Column run at pH 10 and 25°C - Carbon-only system.
Source: Reference 25-
8-33
-------
Empty B«d Con tad Urn* «TU mln
CoJ WTO Height 1107 cm
660 flnw Activated CwfeanDwvB
12 3
OPERATING TIME, DAYS
Figure 8.1.14
Column run at pH 4 and 25°C - Carbon-only system.
Source: Reference 25.
8-34
-------
In another series of laboratory experiments, columns packed with
Huchar 722 activated carbon were used to determine the tnechaniam responsible
9 &
for mercury removal. A 17 percent caustic solution originating from
mercury electrolysis cells was fed to one of the columns, while an aqueous
preparation containing methyl mercuric chloride was percolated through the
second. About 80 percent of the influent mercury was removed in the first
column and no mercury was detected in the discharge from the second. The
investigators concluded from the results obtained that organic mercury is
readily adsorbable, both on an absolute basis and relative to other forms of
mercury. They also postulated that filtration was the dominant mechanism in
the observed removal of finely divided metallic mercury from the caustic
stream.
Successful application of a full-scale activated carbon treatment system
9 ? 9 &
has been reported. * This particular system was devised for handling
small volume, pesticide manufacturing discharges containing organic mercury
compounds. In the process, suspended solids are removed by coagulation and
fioccuLetion with iron salts and polyeiectrolytes prior to carbon adsorption
in a series of packed beds. Mercury loadings of 0.05 kg per kg carbon are
readily attained and the spent carbon sorbent is thermally regenerated.
29
Cyanide—Kuhn patented a process where activated carbon is utilized
as a catalyst for cyanide oxidation. The process involves mixing air and
cyanide-bearing waste, at an alkaline pH, and pulsating the solution through a
bed of activated carbon. Calgon Corporation extended this cyanide
detoxification method by adding cupric ions to the wastewster along with
oxygen prior to passing the cyaniderbearing waste through GAC columns.
The cupric ions accelerate and increase the efficiency of the catalytic
oxidation of cyanide by granular activated carbon. The presence of cupric
ions results in the formation of copper cyanides, which are adsorbed more
readily than copper or cyanide alone.
Calgon demonstrated that the capacity of the granular carbon was limited
to 2-3 mg of cyanide adsorbed per gram of adsorbent when no copper was used.
However, the addition of copper increased the adsorption capacity to 25 mg/g.
In the presence of dissolved oxygen, adsorption sites were continuously
regenerated tnrough the oxidation of the cyanide.
8-35
-------
Based on Calgon's studies, a study was undertaken to investigate the
feasibility of a low-cost activated carbon treatment process for petroleum
refinery wastewater. The conceptualized process evaluated in this study
involves the addition of powdered activated carbon (PAC) and cupric chloride
directly into an activated sludge unit which is commonly used for secondary
treatment at petroleum refineries. Some of the potential benefits of adding
PAC to an activated sludge system include;
• improvement in BOD and COD removals;
* improved solids settling, decreased effluent solids and increased
sludge solids;
* adsorption of dyes and toxic components that are either not treated
biologically or are poisonous to the biological system;
• prevention of sludge bulking over broader ranges of feed to
microorganism;
* effective increases in plant capacity at little or no additional
capital investment; and
* more uniform plant operation and effluent quality, especially during
periods of widely varying organic or hydraulic loads.
From the results of batch tests, the following five parameters were
considered to be the major variable affecting the cyanide treatment using PAG
in activated sludge units: pH, mode of copper addition, carbon type, and
carbon and copper dosages.
Increased cyanide removal rate was observed at lower pH, as shown in
Figure 8.1.15. However, although low pH favors increased cyanide removal,
effluent copper levels were unacceptable. Further study demonstrated that at
pH values near neutral, 95 percent cyanide removal was achieved while
maintaining effluent copper levels of 0.05 mg/1 or below.
Copper salts can be introduced by two different techniques, either
directly into the aeration basin or by being adsorbed onto the carbon prior to
addition. The results did not demonstrate a significant difference between
these methods.
8-36
-------
0.4
03
I
U
C/1
UJ
Q
u
UJ
f-
•3
o:
h-
_j
u.
0.3
O.I
n - pll 10.2, Initial CN~
A - p» 7.9, initial CN~
O - pH 2.5, initial CN~
_*
0.53 mg/0.
0.48 mg/S.
0.48 mg/8.
each reactor contained 250 mg/j? lignite
based carbon and 1.5 mg/^ cupric tons.
*Fcrrocyanide, measured as CN .
0.25
I 6
TIME, HOURS (LOG SCALE)
24
48
Figure 8.1.15. pH effect on rate of cyanide removal.
Source: Reference 7.
-------
Two type of PAG, lignin-based (Agua-Nuchar) and lignite-based tHydarco C)
were evaluated throughout a range of carbon and copper concentration. Typical
results are shown in Figure 8.1.16. When utilizing potassium ferroeyanide,
the lignite-based carbon was superior at all carbon dosages tested, however
the overall improvement diminished as copper dosage was increased. Similar
tests performed with potassium ferricyanide as the cyanide source showed the
lignin-based PAC to be more effective. In both tescs, the equilibrium soluble
cyanide level was reduced as the carbon concentration was increased..
Copper dosage was found to have the greatest influence on cyanide
removal. The data presented in Table 8.1.5 demonstrate that as the carbon
dosage increases, there is greater copper removal in addition to greater
cyanide removal. Hence, concerns over excessive copper effluent levels can be
addressed by increases in carbon dosage.
The above results demonstrate that the addition of PAC/CuCl directly
into petroleum refinery activated sludge aeration basins can enhance cyanide
removal without any detrimental effect on the aicroorgmnisms, provided that
the copper concentration in the influent is maintained at less than 1 mg/1.
The addition of PAC also improves the removals of BOD, COD and TOC.
8.1.3 Cost of Carbon Adsorption
The cost of carbon adsorption treatment can be described in terms of
direct and indirect capital investment"* operation and maintenance costs. For
the small-scale system, direct capital investment costs include the purchase
of a waste storage tank, a pre-filter, carbon columns, a waste feed pump,
piping and installation. For the large-scale system, additional direct
capital investment costs include storage tanks for spent and regenerated
32
carbon and automatic controls. A model has been developed by 1C?, Inc.
33
for calculating carbon adsorption costs. Table 8.1.& presents equations
used to calculate direct capital costs as a function of carbon consumption
rate and storage volume.
Indirect capital costs include engineering and construction costs,
contractor's fee, startup expenses, spare parts inventory, interest during
construction, contingency and working capital. These costs are expressed as
percentages as summarized in Table 8.1.7. Direct and indirect capital costs
are assumed to be incurred in year zero.
8-38
-------
7
ID
0.4
to
0»
E
0.3 —
UJ
O
z
o
UJ
o
I
o
0.2 -
O.I -
O - 0.5 mg/8,
A - 0.5 rng/fc
- 1.0 rag/e
- 1.0 mg/£
Cu , 1 Ignln carbon
Cu , 1 ignite carbon
Cu , 1 Ignln carbon
Cu , 1 Ignite carbon
Initial terrocyanide concentration,
0.5 - 0.05 rag/ as CN~
100 200 300 400 500 600 700
CARBON CONCENTRATION, mq/Jt
800 900 1,000
Figure 8.1.16. Equilibrium cyanide levels as a function
of carbon concentration.
Source: Reference 7.
-------
-------
Reference 7.
TABLE B.I.5. FILTRATE COPPER LEVELS
Initial
copper
concentration
Average Filtrate Copper Levels
100 mg/L
carbon
250
carbon
1,000 mg/L
carbon
0.5
0.06
0.05
<0.02
2.0
0.08
0.05
0.02
1.5
0.19
0.09
8-40
-------
-------
TABLE 8.1.6- DIRECT COSTS FOR CARBON ADSORPTION*
Carbon consumption Direct capital Direct Operation and
rate costs maintenance cost"
(Ibs/day) ($) (*/yr)
less than 400 l,256(c)-603 + 140(s)-54 29(c)'6 + 350(c)(cp) +
619(c)'168(h) + 5(c)Cp)
greater than 400 14,231(c)•522 + 14Q(s)-54 58tc)'657 + 35(c)Ccp) +
105(c)-455(h) +
25,012-383(c)(p) +
1.49 106(c)(f)
where: c — carbon consumption rate in pounds per day
s = storage volume in gallons
cp = carbon price in dollars per pound (&0.8/lb)
h = hourly wage rate in dollars per hour C$14.56/hr)
p = power price in dollars per kilowatt-bour (40.05/KWh)
f •= fuel price (natural gas) in dollars per Btu ($6xlO~6/Btu)
aCost estimates were developed for three model treatment systems (three
small scale and three large scale systems). The cost estimates for these
systems were then used to develop a cost eouation in the form of a power
curve.
"The power requirement is derived from the equipment specifications.
Source: Reference 33.
8-41
-------
TABLE 8.1.7. INDIRECT COSTS FOR CARBON ADSORPTION
Item
Percent
of direct
capital
costs
Percent of the
sum of direct
and indirect
capital costs
Percent
o£ total
annual cost3
Indirect Capital Costs
Engineering and
Supervision
Construction and
Field Expenses
Contractors Fee
Startup Expenses
Spare Parts
Inventory
Interest During
Construction
Contingency
Working Capital
Indirect Operatio.n and.
Maintenance Costs
Insurance, Taxes,
General
Administration
System Overhead
12
10
7
5
2
10
0
0
15
18
10
aThe total annual cost is defined as the sum of the total capital cost
multiplied by the capital recovery factor and the total operation and
maintenance costs.
Refecence 33.
8-42
-------
Operation and maintenance costs also consist of direct and indirect
costs. Direct operation and maintenance costs include operating labor,
electricity, and carbon consumption. Table S.l.fi also contains the equations
used in the model to calculate direct operation and maintenance costs. All
costs are presented for four flow rates ranging from 100 to 2,500 gai/hr.
8«l«4 Overall Status o£ Process
Activated carbon is a widely used technology for treating waste streams
containing organic compounds. In contrast, the application of activated
carbon technology to the treatment of metal and cyanide containing
uastestreams is limited. However, the ability of activated carbon to treat
these wastestreatns has been demonstrated at bench, pilot, and full-scale
levels. Full-scale systems have been used commercially to treat chromium and
mercury wastestreams, but these applications are few in number. Performance
data is difficult to acquire due to confidentiality agreements between
activated carbon manufacturers and their customers, '
Environmental impacts can occur when the exhausted activated carbon must
be regenerated or disposed. The regeneration of activated carbons used for
the treatment of metals or cyanides is accomplished using a strong acid or
base. Regeneration is usually not performed unless there is an economic
34
incentive to recover the adsorbed metals, C<
is typically disposed of in a secure landfill.
34
incentive to recover the adsorbed metals. Consequently, the spent carbon
8-43
-------
TABLE .8.1.8. CARBON ADSORPTION COSTS3
100
Quantity processed
(gal/hr)
4-00
1,000 2,500
Capital Expenditures
Capital Cost Including Installation1*
($1,000)
Annual Operacion and Maintenance ($l,000)c
59
561
904
1,462
Energy
Labor
Carbon
Other
Capital Recovery
Total Annual Cost
Cost per 1,000 gald
2
23
7e
1
10
42e
210e
11
35
27
5
99
177
221
27
53
67
10
• 160
317
159
68
80
168
18 '
259
593
119
aCosts are based on the RCRA E1SK-COST ANALYSIS MODEL.33
"Capital costs for the 100 gal/hr system include waste storage tank,
prefilter, carbon columns, waste feed pump, piping and installation; the
other flow levels (400, 1,000, 2,500) include these units plus storage
tanks for spent and regenerated carbon, a multiple hearth furnace and
automatic controls.
"-These costs are based on the following data:
carbon price = $0.8/lb
hourly wage rage - $14.56/hr
power price = $0.05/kwh
fuel price (natural gas) = $6 x 10~&/Btu
capital recovery factor = 0.177
^Unit costs are based on 2000 hours of operation per year.
eModified to reflect a direct relationship between carbon requirement ar.d
quantity processed.
*Sote: 1984 dollars.
8-44
-------
REFERENCES
I, Weber, W. J., Physieochemieal Processes for Water Quality Control. John
Wiley & Sons. 1972.
2. Lyman, W. J., Applicability of Carbon Adsorption to the Treatment of
Hazardous Industrial Wastes. In: Cheretnisinoff, et al., Carbon
Adsorption Handbook, Ann Arbor Science. 1980,
3. U.S. EPA, An Investigation of Techniques for Removal of Cyanide from
Electroplating Wastes. U.S. Environmental Protection Agency Water
Pollution Control Research Series No. 12010 EIE" 11/71. 1971,
4. Stnithson, G.R., Jr. An Investigation of Techniques for the Removal of
Chromium from Electroplating Wastes. U.S. Environmental Protection
Agency, Water Pollution Control Research Series. No. 12010 EIE 03/71.
1971.
5. ICF Inc. Survey of Selected Firms in the Commercial Hazardous Waste
Management Industry: 1984 update. Final report to U.S. EPA, Section
II. OSW Washington, D.C. 1985.
6. Huang, C.P. and M.H, Wu. The Removal of Chromium(Vl) from Dilute Aqueous
Solution by Activated Carbon. Water Research 11:673-679. -Fergamon
Press. 1977.
7. Huff, J.E., E.G. Fochtman and J.h, Bigger, Cyanide Removal from Refinery
Wastewater Using Powdered Activated Carbon, In: Cheremisinoff and
Ellerbusch, Carbon Adsorption Handbook, Ann Arbor Science. 1980.
8. Wilk, Lisa et al. Alliance Technologies Corporation. Technical Resource
Document for Treatment of Corrosive Wastes. Prepared for U.S. EPA HWERL
under Contract Ho. 68-02-3997. October, 1986.
9. Slejko, F.L., Applied Adsorption Technology, Chemical Industry Series
Volume 19. Marcel Dekker, Inc. NY, NY. December 1985.
10. Huang, C.P., Chemical Interactions Between Inorganics and Activated
Carbon, In: Chereeisinoff and Ellerbusch, Carbon Adsorption Handbook,
Ann Arbor Science. 1980.
11. Jevtiteh, H.M. and D. Bhattacharyya. Separation of Heavy Metal Chelates
by Activated Carbon: Effect of Surface and Species Charge. Chem. Eng.
Conmun. 23:191-213. 1983.
12. Steenberg,. B. Adsorption and Exchange of Ions on Activated Charcoal,
Almquist and Wilksell, Uppsala. 1944.
13, Benjasin, M.M. and J.O, Leckie. Conceptual Model fcr
Metal-Ligand-Surface Interactions During Adsorption. Environmental
Science & Technology 15:1050-1057. 1981.
8-45
-------
14. Huang, C.P. and E.H. Smith. Removal of Cdt.II) from Placing Waste Water
by an Activated Carbon Process. In: Chemistry of Water Reuse, Ed:
Copper, W, Ann Arbor Science* L981.
15. Huang, C.P. and F.B. Ostovic. The Removal of Cadmiuro(II) from Dilute
Aqueous Solution by Activated Carbon Adsorption. J. Environ. Eng. Div.,
ASCE L04(BE5). 1978.
16. Langmuir, I. The Adsorption of Gaaea on Plane Surfaces of Glass, Mica,
and Platinum. Jour. Am. Chew. Soc., 40:1361-1403. 1918.
17. U.S. EPA Background Document for Solvents to Support 40 CFR Part 268,
Land Disposal Restrictions, Volume II. January 1986.
18. Rizzo, J.L. Calgon Corporation, Letter to Paul Frillici, Alliance
Technologies Corporation. May 5, 1986,
19. U.S. EPA. Activated Carbon Treatment of Industrial Wastewater-Selected
Papers. EPA-600/2-79-177. Robert S. Kerr Environmental Research
Laboratory. August 1974.
20. IT Enviroscience, Incorporated. Survey of Industrial Applications of
Aqueous-Phase Activated-Carbon Adsorption for Control of Pollutant
Compounds from Manufacture of Organic Compounds, Prepared for U.S. EPA
IERL; PB-83-200-188. April 1983.
21. Lyman, W.J. Carbon Adsorption, In: Unit Operations for Treatment of
Hazardous Industrial Wastes. Pollution Technology Review No. 47, Noyes
Data Corporation, Park Ridge, New Jersey. 1978.
22. Huang, C.P. and P.K. Hirth. The Development of an Activated Carbon
Adsorption Process for the Treatment of Cadmium(lI)-Flating Waste. Heavy
Metals in the Environment, Amsterdam, The Netherlands, 15-18,
September 1981.
23. Gupta, S.K. and K.Y. Chen. Arsenic Removal by Adsorption. Journal
WPCF. March, 1978.
24. Praino Jr., R.F. and R. O'Gorman. Technology Evaluation, Installation
and Performance of a Chromium Removal System for Aqueous Discharges.
Hazardous Waste 1:469-487. Publ: Mary Ann Liebert, Inc. 1984.
25. Humenick, Jr., M.J. and J.L. Schnoor. Improving Mercury(II) Removal by
Activated Carbon. J. Environ. Eng. Div., ASCE 100E6:1249-1262. 1974.
26. Smith, S.B.; Hyndshaw, A.Y.; Laughlin, H.F. and S.C. Maynard (1971)
Mercury Pollution Control by Activated Carbon: A Review of Field
Experience. Rept. M1002.01, Westvaco Corp., Covington, VA.
27. Eosenzwsig, M.D. (1975) Mercury Cleanup Routes - I. Chea. Ens., Vol. 82,
No. 2, pp. 60-61.
8-46
-------
28. TNO (undated) Purification of Mercury-Containing Waste Haters.
Apeldoorn, The Netherlands.
29. Kuhn, R.G. Process for Detoxification of Cyanide Containing Aqueous
Solutions. U.S. Patent 3,586,623. June 22, 1971.
30. Bernardin, F.E. Cyanide Detoxification Ueine Adsorption and Catalytic
Oxidation on Granular Activated Carbon. J. Water Poll. Control Fed,
34:221-231. 1973.
31. Adams, A.D. Improving Activated Sludge Treatment with Powdered Activated
Carbon. Presented at the 28th Annual Purdue Industrial Waste Conference,
Purdue University, Lafayette, Indiana. May 1-3, 1973.
32. U.S. EPA. Development Document for Effluent Limitation Guidelines and
Standards for Petroleum Refining Point Source Category.
EPA-440/1-82-014. 1982.
33. ICF, Inc. RCRA Risk-Cost Analysis Model, Phase III U.S. EPA, OSW.
March 1984.
34. Roy A., Calgon Corporation. Telephone conversation with D. Sullivan,
Alliance Technologies Corporation. March 5, 1987.
8-47
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8.2 ION EXCHANGE
Ion exchange has been used commercially to recover metal-containing
wastes from the metal finishing, electroplating, and fertilizer manufacturing
industries. These wastes contain dissolved metal salts which dissociate to
form metal ions. In conventional ion exchange, metal ions from dilute
solutions (e.g., plating rinses) are exchanged for ions which are held by
electrostatic forces to charged functional groups on the surface of the
exchange resin. An alternative design, the acid purification unit, adsorbs
acids from concentrated solutions (e.g., etchants) and allows metal
contaminants to pass through the system. In both cases, the adsorbed
constituent is subsequently removed by contacting the resin «ith a regenerant,
resulting in a potentially recoverable by-product stream which is highly
concentrated in the adsorbed constituent.
The major applications of ion exchange are water purification
(deionization) and selective removal of toxic heavy metal and metal-cyanide
complexes from dilute wastewater streams. Rinse water is reused and metal
contaminants are concentrated in the regenerant stream, allowing more
economical treatment and enhancing their recovery potential. As an end of
pipe application, ion exchange resins have been applied for selective removal
of toxic compounds, while allowing nontoxic dissolved ionic solids to remain
in solution. The acid purification unit (APU) has been successfully applied
commercially for recovery of steel pickling, aluminum anodizing, etchants, and
rack stripping operations.
8.2,1 Ion Exchange Process Description
General System.Description—
The ion exchange system may be operated in a batch or flow-through
(column) mode, the latter being generally preferred due to greater exchange
efficiencies! With the batch mode of operation, the ion exchange resin and
the. waste.solution are mixed in a batch tank. , Upon completion of the exchange
reaction (i.e., equilibrium is reached), the resin is separated from the
treated solution by filtering or settling, regenerated, and reused. Unless
8-48
-------
the resin has a very high affinity for the contaminant ion, the batch mode of
operation is chemically inefficient and thus has limited applications.
Flow-through operation involves the use of a bed or packed column of the
exchange material (resin). These systems are typically operated in cycles
consisting of the following steps:
i. Service (exhaustion) - Waste solution is passed through the ion
exchange column or bed until the exchange sites are exhausted.
2. Backwash - The bed is washed (generally with water) in the reverse
direction of the service cycle in order to expand and resettle the
resin bed.
3. Regeneration - The exchanger is regenerated by passing a
concentrated solution of the ion originally associated with it
through the resin bed or column; usually a strong mineral acid or
base.
4. Rinse - Excess regenerant is removed from the exchanger; usually by
passing water through it.
A flow-through (column) system can be designed with cocurrent or
countercurrrent flow of the waste and regenerant (steps 1 and 3 above). In
cocurrent systems, the feed and the regenerant both pass through the resin in
2
a downflow mode. Figure 8.2.1 illustrates the cocurrent flow process.
Each ion exchange unit consists of a cylindrical vessel having distributors or
collectors at the top and bottom. Resin is loaded into approximately half of
the vessel to accommodate resin expansion during the backwash cycle.
Cocurrent systems are only cost-effective for weak acid or base exchangers
which do not require highly concentrated regenerant solutions. However,
regeneration of strong exchangers (high exchange capacity) requires strong
acid and base solutions which can be more costly,
Often it is too costly to full^ regenerate a bed. In order to avoid
carry over of contaminants into the next service run, two or more sets of
fixed columns arranged in parallel series can be used. Similarly, to avoid
excessive downtime during the regeneration cycle, dual sets of fixed columns
can be used. While one set of columns is being regenerated, the second set of
columns will be switched on line permitting continuous operation of the
system. Improved regenerant efficiency can also be accomplisned by reusing
8-49
-------
i-n
O
WASTE
SOLUTION
CONTAMINATED
BACKWASH
THEATED
EFTLUtN f
WATER
BACKWASH
STRONG
ACID/ALKALI
— WATER
H t It
CONTAMIflATED
REGENERATE
REGENERATION
I „ CONTAMINATED
RINSE
RINSE
Figure 8.2.1. Cocurrent ion exchange cycle.
Source: Reference 2.
-------
3 3
the last portion of the regenerant solution. For example, if 5 Ib/ft
(80 g/L) of regenerant were used for the system shown in the figure, the first
50 percent of spent regenerant would only contain 29 percent of the original
acid concentration, whereas the remaining regenerant would contain 78 percent'
3
of the original acid. If the last portion of the regenerant is reused in
the next cycle before the resin bed is contacted with fresh HC1, the exchange
capacity would increase from 60 to 6? percent at equal chemical doses.
In addition to cocurrent designs, countercurrent systems are available
«hich result in a more efficient use of regenerant chemicals. They also
achieve a higher concentration of contaminant in the regenerant stream thereby
enhancing potential for further recovery. A widely used countercurrent design
for chemical recovery from plating rinses is the reverse or reciprocating flow
ion exchanger (RFIE), as depicted in Figure 8.2.2. Another variation of this
design, which uses countercurrent flow through a fluid bed, is called the acid
purification unit (APUJ, developed by Eco-Tech, Ltd. Instead of adsorbing
metallic species, the resin adsorbs- acid which is then regenerated by flushing
the bed'with water (see Figure 8.2.3).
Resin Selection—
The most significant design parameter in an ion exchange system is the
selection of an appropriate resin. Resin selection is based on the type of,
ion exchanger, flow volume and the resin's strength, exchange capacity and
selectivity. Resins can be classified by functional (reactive) groups and the
type of exchangeable ions present. Exchanger categories include strong and
weak, cation, strong and weak anion, and chelating ion exchangers. Some of the
more common reactive groups are:
Reactive Croup Exchangeable Ions
Strong acid (suifonic) Cations in general
Weak acid (carboxylic) Cations in general
Weak acid (phenolic) Cesium and polyvalent cations
Strong base (quaternary amine) All anioris, esp. used for anions of
weak acids (cyanide, carbonate,
silicate, etc.)
8-51
-------
RINSE
WATER
M SO
CO
(Jl
!
CATION
,
1
ANION
-
1
EXHAUST
(TO WASTE
FILTER
1
CATION
EXHAUST
t
CATION
t
i
TREATMENT]
t
ANION CATION
j
WATER
t
2 4
=^~^
S~^:
1
CATION
1
t
AMION CATION
t
WATER
--3==
Figure 8.2.2. Schematic of a fixed bed reverse flow ion exchange (RFIE) system
for the recovery of chromic acid from a dilute solution.
Source: Reference 3.
-------
I i I
UPSTROKE
COMPRESSED AIR
T
1
RESIN BED
*
i i i
T I T
SPENT
ACID
(FEED)
COMPRESSED AIR
DOWNSTROKE
.111
T T T
WATER
T
RESIN BED
T
SPENT ACID (FEED)
i
i A
PURIFIED ACID IPRODUC"
Figure 8.2,3, Baste operation of the acid purification unit (APU)
using a continuous bed RFIE system.
Source: Reference 1,
8-53
-------
Weak base (tertiary and secondary Anions of strong acids (sulfate,
amine) chloride, etc.)
• Chelating (varied, may be imino- "• Cations, especially transition and
diacetate or oxime groups) heavy elements
Cation exchangers have positively charged, mobile ions for exchange.
Strong acid cation resins are those containing functional groups derived from
a strong acid. Their behavior is similar to that of a strong acid in that
they can convert a metal salt to the corresponding acid. Both the hydrogen
form (used for deionization) and sodium forms (used for water softening) are
highly ionized. Due to the highly dissociated nature of these resin typea,
their exchange capacity is independent of solution pH.
Weakly acidic cation exchangers are resins derived from a weak acid.
These resins behave like weakly dissociated organic acids. The degree of
dissociation is strongly influenced by solution pH and they tend to
2
demonstrate limited capacity below a pH of 6 . Due to the pH limits, weak
acid resins are unsuitable for deionizing acidic wastes.
Strong and weak anion exchange resins behave in a fashion analogous to
cation exchangers. Strongly basic anion exchangers are highly ionized and can
be used "over a wide pH range. Weakly basic anion resins are strongly
influenced by solution pH and exhibit limited exchange capacity above a pH
of 73.
Chelating resins behave similarly to weak acid cation resins, but are
highly selective for heavy metal cations. This type of resin forms an
essentially non-ionized complex with divalent metal ions. Consequently, once
an exchanger group is converted to the heavy metal form, it is relatively
unreactive with other similarly charged ions in solution, regardless of
concentration. Chelating reains will effectively remove heavy metal cations
from solutions of pH 4- and above.
The exchange capacity of a resin is generally expressed as equivalents
per liter (eq/L, where an equivalent is equal to the molecular weight of the
ion, in grams, divided by its electrical charge or valence. For example, a
resin with an exchange capacity of 1 eq/L could remove 37.5 g of divalent zinc
* , molecular weight = 65 g) from solution.
8-54
-------
As noted above, solution pH can have a significant effect on exchange
capacity for weakly acidic, anionic, and chelacing resins. For example, the
effect of pH on the exchange capacity of Rohm and Haas Amber lite IRC-718, a
chelating resin specifically designed for selective heavy metals removal, is
quite dramatic. Because of the resin's affinity for hydrogen ions, the
capacity for most other ions falls off sharply below pH 4. Figure 8.2.4
compares the capacity of Amberlite IRC-718 when used to remove nickel from a
waste containing calcium chloride at pH 2 and pH 4. The data show that good
removal is realized for 200 bed volumes 11,500 gal/ft of resin) when
treating the stream at pH 4 whereas breakthrough occurred at pH 2 in less than
50 bed volumes,
Figure 8.2.5 illustrates the effect of pH on the capacity of Amberlite
DP-1 (a weak acid cationic resin) and Amberlite IRC-718 when these resins are
5 • +2
used for removal of cadmium at pH 2.1 and pH 8.0. In this example, Cd
was present at a concentration of 50 pptn with 1,000 ppni calcium chloride. At
a flow rate of 8 bed volumes/hour (.1 gpm/ft ) and a pH of 2.1, both resins
showed sharp breakthrough curves with end points less than 100 bed volumes.
Conversely, at pH 8,0, Amberlite IRC-718 showed less than 0.1 ppm leakage for
350 bed volumes while leakage from Amberlite DP-1 remained under 0.1 ppm for
520 bed volumes.
The metal removing performance of an ion exchange resin is also
influenced by its ionic form. For example, Amberlite IRC-718 is available in
hydrogen and sodium forms. Figure 8.2.6 demonstrates the difference in
exchange capacity for this resin for the removal of copper from a stream
+ 2 5
containing 50 ppm Ca and 1,000 ppm CaCl, at pH 4. Breakthrough
occurs much sooner for the sodium form of Amberlite IRC-718 despite the fact
that the two resins demonstrated comparable removal efficiencies.
Ion exchange reactions are stoicbiometric and reversible. A generalized form
of an ion exchange reaction can be described as follows:
R-A"1" + B+ ^ R-B* -•- A*
where R is the resin, A is the ion originally associated with the resin,
and B is the ion originally in solution.
The degree to which the exchange reaction proceeds is dependent on the
preference, or selectivity, of the resin for the exchanged ion. The
8-55
-------
CD
I
Amberllte 1RC-71B lor Nickel Removal
In the Pretence of (Jalclum
Influent—Nl+* SO ppm.
CeCI, 1000 ppm.
Flow Rate—8 bod volumes/hour
-» Nickel Leakage pH-2
-* Nickel Leakage pH-4
I
75
100
125
r
ISO
I
175
200
I
225
250
Figure 8.2.4.
Breakthrough curve demonstrating variable pH performance
of Amberlite IRC-718.
Source: Reference 5.
-------
6.0-
AmbeiTHe IRC-718 vs AmbeilHe DP-1
Cd« 50 ppm
CaCI, 1000 ppm
Flow Rale 8 bed volumes/hour
AmberlUe DP-1 pH 2.07
Amberlite IRC-718 pH 2.07
Amberlile DP-1 pH 80
Amberlile IRC 718 pH 8.0
Jt *.
100
Bed Volumes
300
400
500
800
Figure 8.2.5. Breakthrough curves comparing pll variable performance
of chclating (Amberlite IRC-718) and weak acid
(Amberlite DP-1) resins.
Source: ReEerencc 5.
-------
oo
6.0-
5.0 —
4.0 —
aj 3.0 —
c
.3
5
- 2.0—|
O
a 1.0^
Amberlite IRC-718 (Na+) vs. Amberlite IRC-718 (H+)
Influent
50 ppm
1000 ppm
4.0
8 bv/hr.
Amberiite IRC-718 (Na+)
Amberlite IRC-718 (H+)
PH
Flow Rate
100 200
Bed Volumes
300
400
500
600
700
800
Figure 8.2.6. Comparison of chelating resin performance (Hydrogen form
vs. Sodium form) using Amberlite IRC-718.
Source: Reference 5.
-------
selectivity coefficient, K,, expresses the relative distribution of ions when a
charged resin is contacted with solutions of different, but similarly charged
ions.. For example, in- the generalized ion exchange reaction presented above,
the selectivity coefficient (K) is defined as follows:
[B ] in resin lA j in solution
K - x
[A j in resin IB 1 in solution
The selectivity coefficient of a resin will vary with changes in solution
characteristics and the strength of the resin. Table 8.2.1 summarizes the
selectivities of strong acid and strong base resins for various ionic
species. ' Resin selectivity is dependent upon ionic charge and size. The
force with which an ion is attracted is proportional to its ionic charge and,
therefore, the counter ion of higher valence is more strongly attracted into
3
the exchanger. The preference o£ exchange resins for counter ions of
highest charge increases with dilution of the external electrolyte and is
strongest with exchangers of high internal molarity.
With regard to ionic, size, ions of smaller radius are preferentially
adsorbed. When the resin is in a polar solvent, such as water, the fixed ions
within the exchanger and mobile ions in both the resin and the solution tend
to hydrate, causing the resin to swell. Hydration of the ions exerts a
swelling pressure within the resin which is resisted by the cross-linked
polymer matrix holding the resin particle together. As a result, the resin
prefers the ion of smallest hydrated radius, since smaller ions can most
readily enter the matrix of the resin and react with its functional groups.
In general, tnultivalent hydrated ions are smaller in size than an equivalent
charge unit ol ions of lower valence.and are therefore preferentially
adsorbed. Within a given series of ions, the hydrated radius is generally
inversely proportional to the unhydrated ionic radius.
Another factor affecting resin 'selectivity- is the interaction of ions
within the exchanger and in bulk solution. The exchange resin prefers
counter-ions which associate most strongly with the, fixed ionic groups. If
the groups are sitnilsr in structure to precipitating or cotsplsxing agents for
a particular ion, the resin will prefer that ion. As a result of this
phenomenon,'aany resins containing ehelating functional groups show pronounced
selectivities for transition group metal ions.
8-59
-------
TABLE 8.2.1. SELECTIVITIES OF ION EXCHANGE RESINS IN ORDER OF
DECREASING PREFERENCES3
Strong acid
cation
exchanger
Strong base
anion
exchanger
Weak acid
cation
exchanger
Weak base
anion
exchanger
Weak acid
chelate
exchanger
Barium (+2)
Lead (+2)
Mercury (+2)
Copper (+1)
'Calcium (+2)
Nickel (+2)
Cadmium (+2)
Copper {-»-2)
Cobalt (+2)
Zinc (+2)
Cesium (+1)
Iron (+2)
Magnesium (+2)
Potassium (+1)
Manganese C+2)
Ammonia (+1)
Sodium (+1)
Hydrogen (+1)
Lithium C+l)
Iodide (-1)
Nitrate (-I)
Bisulfite (-1)
Chloride (-1)
Cyanide 1^1)
Bicarbonate (-1)
Hydroxide (-1)
Fluoride (-1)
Sulfate (-2)
Hydrogen (+1)
Copper C-i-2)
Cobalt {+2}
Nickel (+2)
Calcium (+2J
Magnesium (+2)
Sodium (+2)
Hydroxide (-1) Copper (+2)
Sulfate (-2) Iron (+2)
Chromate (-2) Nickel (+2)
Phosphate (-2) Lead (+2)
Chloride (-1) Manganese (+2)
Calcium (+2)
Magnesium (+2)
Sodium (+1)
aValence number is given in parentheses,
Source: References 3 and 6.
8-60
-------
The major disadvantage of a high degree of selectivity in an exchange
reaction is Che reluctance of the resin to release the ion during
regeneration. Figure 8.2,7 illustrates the elution curves for zinc from a
chelating resin and a weak acid cation resin with a 10 percent HC1 regenerant
and a flow rate of 8 bed volumes/hour Cl gpm/ft ). The weak acid cation
resin, Amberlite DP-1, gives a sharper elution curve, demonstrating the «
relative ease with which it is regenerated. Conversely, the highly specific
cbelating resin, Amberlite IRC-718, requires nearly twice as much regenerant.
Although chelating resins clearly offer superior selectivity for metals
removal, a weakly acidic cation exchange resin in the sodium form can
sometimes exhibit equal or superior capacity and regeneration efficiency when
treating heavy metal waste streams.
Operating Parameters—
Operating parameters vary considerably depending or. the particular
application. The following factors will influence the selection of a resin
type, pretreattnent requirements, flow rates, cycle times, and the sizing of a
system for a particular application:
» Types and concentrations of constituents present in the feed;
• rate of metal salt accumulation in the bath;
» flow rate; and
• number of hours of operation.
The types and concentrations of constituents present in the spent
solution will determine the type of resin selected. Weak cation exchangers
•can be used for spent solutions containing low concentrations of metal ions.
For solutions containing'high concentrations, a strong anion exchanger may be
preferred. The constituent concentrations and waste volume will also
determine the resin volume needed to treat the stream. Commercially available
systems are aole to process wastes at throughput rates ranging -from 38 to
o
6,700 liters/hour. Cycle times for RFIE systems generally range from 5 to
O Q
15 minutes ' , whereas for cocurrent systems, they can be as much as 1 to
2 hours because of the time needed to regenerate the column. As a result,
dual sets of columns are typically used in cocurrent systems to avoid
excessive downtime.
8-61
-------
14,000
12,000
Eluilon ol Znf* Frem Arnbertiie DP-1
and Amberlile 1RC-718
Regene/anl 10% NCI
Flow Rale 8 Bed Volumes/Hour
• Amberlile DP-t
Amb*rlile IRC-718
1 2
Bed Volumes
Figure 8,2.7. Begeneracion performance of Amberlite DP-1 (weak
acid resin) versus Amberlite IRC-718 (cbelating resin),
Source: Reference 5.
8-62
-------
Pretreatment Requirements—
Pretreatment of the waste stream (visually via filtration) is often
necessary to remove many constituents which would otherwise adversely affect
the resin. Certain organics (e.g., aromati.cs) become irreversibly sorbed by
exchange resins, and oxidants, such as chromic or nitric acid, can damage the
resin. Sodium metabisulfite, which converts hexavalent chrome to its
trivalent state, can be added to the solution to prevent damage to the
resin, Eco-Teeh claims that resin degradation is less of a problem with
the RFIE process due to the short duration of contact (e.g., 1.5 min) between
the acid and the resin. '
High concentrations of suspended solids, which can foul the reain bed,
are typically pretreated through some form of'filtration; e.g., activated
carbon, deep bed, diatomaceous earth precoat, and resin filters. The filters
eventually become clogged with particulates, and are replaced when overall
cycle time increases to unacceptable levels due to excessive head loss.
For large volume systems which require frequent changing of filter cartridges,
it may be more cost-effective to use a multimedia sand filter with a
baekwashing regeneration system. Although initial capital costs are higher,
o
significant savings in filter replacement costs can be realized.
The use of weak acid or base exchangers for treating wastes will require
additional pretreatment. The exchange capacity of weak acid exchangers is
generally limited below pH 6.0, and weak base exchangers are not effective
4
above pH 7. Therefore, a pH adjustment system must be incorporated prior
to feeding the waste stream to the exchanger.
Ion exchange using cocurrent flow is not economically suitable for
removal of high concentrations of exchangeable ions; i.e., above 2,500 mg/L,
expressed as calcium carbonate equivalents. Above this level, the resin
material is rapidly exhausted during the exchange process and regeneration
becomes prohibitively expensive, ' However, the reverse is true for acid
purification units since they are capable of recycling the regenerate. In
addition, higher acid concentrations in the waste feed solution will improve
APU removal efficiencies.
8-63
-------
Post-Treatment Requirement-s-—
Overall savings in treatment and disposal costs can be realized through
the use of ion exchange since, being a separation process, the total volume of
wastes generated is reduced. Waste streams from ion exchange include spent
regenerant solution, wash, and filtered solids. Cocurrent ion exchange
generates an additional waste stream as a result of the need to backwash and
expand the resin bed prior to regeneration (see Figure 8.2.1). Spent
solutions from cocurrent operations are generally combined and managed through
neutralization, precipitation, and disposal of the resulting metal sludge.
Recovery of the regenerant or metals may not be economically justified since
regeneration is conducted relatively infrequently in applications where this
process is typically used; e.g., polishing treated effluents. Conversely,
RFIE units generate more highly concentrated regenerant solutions which are
more amenable to metal"recovery (e.g., electrolytic recovery) and regenerant
reuse.
The only waste product generated from an APU is a moderately acidic
metallic salt sludge. This also may be amenable to metal recovery
techniques. The recovered acid stream is generally reused as make-up in the
processing bath which was being treated.
For all units, filtrate from prefiltering systems can generally be land
14
disposed without further treatment. Otherwise, these can be managed
through dewateriug and solidification prior to landfilling. The quantity of
sludge generated will depend on the types and concentrations of suspended
solids present in the waste solution.
8.2.2 Proces s Performance
The performance of an ion exchange system will be predominantly
influenced by the characteristics of the waste stream being treated including:
types and concentrations of constituents present, acidity of the epent stream
relative to that of the fresh stream, and required effluent quality. Factors
which must be considered when evaluating system performance include: the
quantity of residuals generated, cycle time,, product concentration, process
modifications required, attainable flow rate,"system size, and overall
processing costs.
8-64
-------
14 15
A comparison of ion exchange systems is shown in Table 8.2.2. *
Coeurrent flow units have the lowest capital costs but also the highest
operating costs per unit of contaminant removal. RFIE units are generally
more cost-effective than cocu'rrent fixed-bed systems for wastes with
appreciable contaminant content (e.g., plating bath rinses). They use smaller
resin volumes, minimizing capital costs and space requirements, nave lower
operating costs as a result of regenerant reuse, are capable of handling
higher volume flows, and generate more consistent effluent quality. Examples
of industrial applications for the various types of Lon exchange units,
including APUs, are discussed below.
Effluent Polishing: Plating Facility—
Polishing of effluents from conventional treatment svstems using ion
exchange has been applied successfully at a number of commercial
installations. As an example, the Mogul Corporation designed a 2-stage, fixed
bed polishing system for a client that could not meet effluent standards for
4
Zn, Ni, Cu, and Cr« -The plating facility originally used sodium bisulfite
chromium reduction and hydroxide precipitation to batch treat four segregated
heavy metal plating waste effluents. Ion exchange was selected as the ideal
choice to polish the combined waste discharge at this facility for several
reasons. With a centralized treatment facility in place, no additional
chemical destruction systems were needed to treat regenerant solutions.
Similarly •, no investment was needed for sophisticated pH control systems,
flocculant feed systems, clarifiers and other process ecuiptnent. Finally, ion
exchange units are compact and easy to automate and, therefore, were not,
difficult to incorporate into the existing system.
Rohm and Haas Company recommended using Amberlite XE-318 cation exchange
resin for this application since, as a result of its strong chelating
functionality, it is selective for removing transition metals in the presence
of alkali or alkaline earth cations. Laboratory tests indicated that optimal
removal was obtained by using the resin in both its hydrogen and sodium forms
in a two-stage system. Since the selectivity of the resin is less for calcium
ions than for sodium ions, lime was substituted for sodium hydroxide in the
first-stage treatment. Resin column breakthrough tests were then performed to
determine the quantity of resins needed to handle the 500 GPD volume of
plating wastes.
8-65
-------
TABLE 8.2.2. COMPARISON OF ION EXCHANGE OPERATING MODES
Criteria
Cocurrent
fixed bed
Couti t ere ur rent
fixed bed
Countercurrent
continuous
Capacity for high feed
flow and concentration
Effluent quality
Regenerant and rinse
requirements
Equipment complexity
Least
Fluctuates with
bed exhaustion
Highest
Simplest; can use
manual operation
Equipment for continuous Multiple beds,
operation single regenera-
tion equipment
Relative costs (per
unit volume);
Investment
Operating
Least
Middle
High, minor
fluctuations
Somewhat lesa
than coeurrent
More complex;
automatic con-
trols for
regeneration
Multiple beds,
single regener-
tion equipment
Middle
Highest'
High
Least, yields
concentrated
regeneration
waste
Most complex;
completely
automated
Provides con-
tinuous service
Highest chemicals
Lesa chemicals,
and labor; highest water, and labor
resin inventory than cocurrent
Highest
Least chemicals
and labor;
lowest resin
inventory
Source: References 14 and 15.
8-66
-------
Figure 8.2.8 shows a schematic ot the upgraded treatment system. The
primary system will continue to batch treat the segregated wastes which will
then pass through the' ion exchange system. Zinc pit wastes were judged to
require pretreattnent in a second H -form resin "roughing" column before
4- +
entering the 2-stage columns (H -form followed by Na -form resin). This
"roughing stage" column will require regeneration after each use. However the
other wastewaters (Ni, Cr) only require treatment in the 2-stage column
system. This system has the capacity to handle a full week's volume of
wastewater before each column must be regenerated.
Resin regeneration procedures are briefly described as follows:
1. Backwash each column with city water for a minimum of 5 -minutes at
60 gpia to reclassify the resin bed. Backwash water will dump into
the chrome floor sump,
2, Regenerate the resins fully using 300 gallons of 10 percent sulfuric
acid on the H -form columns, and a two-step regeneration of
100 gallons of 10 percent sulfuric acid followed by 50 gal Ions-of- a
5 percent solution for the Na —form column.
3, A final 5 minute rinse of 30 gpn city water for each column.
The effluent from the ion exchange system flows by gravity into a
15,000 gallon underground retention tank. A 24-hr composite sample of the
effluent is collected daily by the sampling pump and an aliquot checked for
compliance with the effluent limits. Table 8.2.3 compares the final effluent
levels achieved by the upgraded treatment system with prior, discharges and
permissible Legal limits.
8-6?
-------
-------
Clly Wot.r
00
I
00
Slurfgi
12,000 //nil
Pre^urn Dual Ion Exchange
Rogulolcr Package
Wllh RflgnneratiDn
11 ll'irj
Landfill
Clly 5>»r<
Manhol.
Ml,.r
Figure 8.2.8. Schematic o£ Lhe enhanced treatment system.
Reference: 4
-------
-------
TABLE 8.2.3. EFFLUENT QUALITY COMPARISON FOR UPGRADED TON EXCHANGE POLISHING
oo
cr
Parameter
pll
Color, pt-co
TI1S, mg/L
COD, mg/L
Cc'idmium, mg/L
Chromium (T) , mg/L
Copper, mg/L
Iron, mg/L
Lead, mg/L
Nickel, mg/L
Zinc, mg/L
State
(Avg.-
6.5
12
15
20
0.01
0.05
0.05
0.50
0.05
0.10
0.10
Limits
-Max.)
- 8.5
- 20
- 30
- 0.01
- 0.10
- 0.10
- 1.0
- 0.10
- 0.30
-0.30
Original
8.0
40
15
200
<(
0.02
0.1
0.2
<(
2.0
1.0
Effluent Range
With Ion
L System Exchange Upgrade
- 9.0 6.9 - 11.6
- 200 0
- 150 <1.0
- 500 210 - 928
3.01 <0.005
- 1.8 <0.02
- 20.0 <0.05
- 0.3 <0.05
3.2 <0.05
- 4.5 <0.05
- 125.0 <0.02
Adapted from Reference No. 4.
-------
The sorption filter design was slightly more expensive than the
multi-media system, but would produce * better quality effluent. Since this
system requires a proprietary media, it has a potentially high built-in
uncontrollable operating cost tied to a single supplier. In addition, the
solid product contains far acre filter media than metal hydroxide.
The precoat filtration option vould utilize three 150 gpm automatic
diatomaceous earth precoat filters in place of the multi-media filters as
described above. The sludge from this system would be mainly diatomaceous
earth containing metal hydroxides.
The ultrafiltration design consisted of pretreatment followed by four
100 gpm trains la parallel, each consisting of a 1,500 gpm pump and
40 membrane modules, Ultrafiltration essentially replaces the multi-media
filters with ultrafiltration units, while everything else remained virtually
unchanged. However, this system required considerable pumping with its
corresponding power costs and maintenance, and the membranes are susceptable
to organic contamination. Ultrafiltratioo units are often used upstream of
ion exchange systems to reduce the process load. In this case, however, the
low inlet metal concentrations eliminated the need for both systems.
The. sorption filtration design was a proprietary polishing system,
yielding low metal effluent concentrations. It consisted of pH control,
sodium sulfide addition, and filtration through a "sorption filter" precoated
with diatomaceous earth. The advantage of this filter was that the media
could be hydraulically pumped from the filters and reinstalled several times
for each fresh charge, thereby reducing media consumption.
The proposed ion exchange system consisted of twin carbon towers for
removal of trace organics, followed by twin sets of dual bed ion exchange
columns manifolded for two-pass flow. The two-pass arrangement insured
against breakthrough of poor quality water and capacity to handle variations
in process load. Since the ion exchange regenerate solutions could average
between 30 and 60 gpm, a second concentrate treatment system was required,
The ion exchange system was the least expensive to operate, but its
capital cost was nearly twice that of the multi-media or sorption filter
systems. However, since the exchange system recycles 80 to 90 percent of the
processed water, it realizes significant savings in water/sewer charges. In
addition, the recycled water is already warm and'does not.nave to be heated
8-70
-------
from 40°F to 70°F, as does once-through water. Another advantage of the ion
exchange system is that it concentrates the metals, thereby decreasing the
size of the treatment system, and increasing the efficiency of the reducing
agents. Concentrates, however, often require an additional stage of treatment
to reach the low effluent levels mandated by environmental regulations.
Effluent Polishing: Chlor-Alksli Plant—
Akzo Chemicals Company' of the Netherlands developed a process for the
removal of mercury using Rohm and Haas Duolite GT-73, a weak acid cation
exchange resin with a high degree of specificity for mercury. The Duolit*
GT-73 resins utilize thiol (-SE) functional groups,•which tend to form very
strong bonds with ionic mercury. The process, as installed at a
chlor-alkaii plant, involved the following steps:
* Oxidation—Since the resin reacts only with ionic mercury, metallic
mercury must be converted to the ionic form. To accomplish this, an
oxidation step is required, with solution pn maintained at 3 to
prevent iron precipitation. To prevent clogging, metal hydroxides
and unreacted mercury are filtered with sand or cloth filters.
"*>•
* Dechlorination—The resin's thiol groups are readily oxidized,
therefore, removal of chlorine is essential to retain resin
activity. 'The Akzo process employs a two-stage dechlorination
step. First, the stream is reacted with NaHSC^, Na£S03 or
S02, and then it is passed through an activated carbon column.
* Ion Exchange—Two.resin beds are used in series operating in a
counter-current mode. One bed acts as a roughing stage and the
second unit as a scavenger.
Figure 8.2.9 presents a schematic of the Akzo process for mercury removal in
the treatment of chlor-alkali wastewater. It has been demonstrated to produce
3
a mercury concentration well below 5 ppb at a flow rate of 1.25 gpia/ft .
Figure 8.2.10 presents a typical breakthrough curve for chlor-alkali brine at
pH 2 and a feed concentration of 20 to 50 mg/L mercury.
The performance of che Duolite GT-73 resin can be further demonstrated by
its Freundlich isotherm, as shown in Figure 8.2.11. Figure 8.2.12 illustrates
trie capacity of the resin as a function of mercury concentration in the feed.
Finally, Figure 8.2.13 presents a typical elution curve for Duolite GT-73,
showing that it is readily regenerated using concentrated HC1.
8-71
-------
Regeneration Liquid to Brine Cfrcuil
OXIDATION FILTRATION OECHLQRINATiON
DUOL1TE 1 Ef!tuen*< Bppb Hg
GT-75 I
I , J
i Rs^CfiftrolsOfi Liquid
PIgura 9.2.9. Afczo process for mercury removal,
Source: Reference No. 17.
8-72
-------
1000
ioo
MERCURY CONCENTRATION
IN EFFLUENT,ppfc
DETECTION LIMiT
DUOLITE GT-73
BREAKTHROUGH
EFFLUENT VOLUME
1000
5000
10,000
Figure 8,2.10. Typical breakthrough curve cblor-alkali plant
brine, pri 2, mercury concentration in feed 20-50 mg/I
Source; Reference No. 17.
8-73
-------
IOO
ao
«o
M«rcury C
in Risin a
CM 0.6 0,8 i.o
Figure 8.2.11. Freundlich Isotherm of Duclite GT-73,
Source: Reference No. 17.
8-74
-------
IOO.OOO
80,000
60,000
40,000
zo.ooo
10.000
6,000
*,OQO
CAPACITY
volume of wos»e Woic
FEte CQNCENTB4TION
or MEHCunr
J L
Fipure 8.2.12. Capacity of Duolite GT-73 relative to the
concentration of mercury in the feed.
Source: Reference No. 17.
8-75
-------
20
15
QHg/iiter HCI 35%
Number of bedvolumes
in3 HCI/m3 resin
Figure 8.2.L3. Regeneration of Duolite GT-73 with concentrated
HCI. Regeneration rate I m^ HGl/ffl^ resin hour
mercury concentration of resin prior to
regeneration 35 a He/Liter resin..
Source; Reference No. 17.
8-76
-------
Effluent Polishing: Printed Circuit Board Manufacturer--
Honeywell Corporation upgraded its printed circuit board facility
wastewater treatment plant to accomodate production increases and reduce water
consumption. The existing treatment system consisted of neutralization
followed by automatic precoat pressure filtration with ion exchange as a rinse
water recycling step. Several alternatives were considered to polish the
effluent sufficiently to meet discharge levels while simultaneously reducing
rinse water consumption from a projected 300 to 400 gpra to less than 75 gpm.
Rinses which could be recirculated after treatment included those from
combined rinses, alkaline etchants, ammonium persulfate deoxidizer, sulfuric
acid and copper plate, and tin-lead and solder strip. Concentrates to be
treated included the residuals created by rinse treatment, plating bath dumps,
and spills.
Technologies considered for process enhancement included multimedia
filtration, sorption filtration, automatic precoat filtration,
ultrafiltration, and ion exchange. All process options would require
2
6,000 ft of floor space, except the multi-media system which needed
5,000 ft . All of the systems met the effluent discharge limits of less
than 1 ppm Cu and Ni and 0.5 ppm Pb. These alternatives are described below
and their economics are summarized in Table 8.2.4.
The" multi-media design would consist of two 400 gpm lined carbon steel
pressure vessels charged with various filtration media. The rinse water was
to be pH controlled and reducing agents added. The filter would require
baekwashing every 4 to 8 hours which would be pumped to a sludge conditioning
tank. The final product would be a metal hydroxide sludge that contained
little or no filter aide. The dewatering filter presses used for the
concentrates would need to be increased in size-since they also had to handle
the main filter backwash.
The multi—media filtration system is the simplest, most compact and least
capital intensive of the evaluated systems. Second only to the ion exchange
system, it would have been the least expensive to operate. This type of
filter can withstand higher feed metal concentrations than the other systems,
and should produce an effluent between 0.4 and 1.0 Dum cooper. The laree
amounts of backwash would require a large sludge conditioning tank and
dewatering filter press. However, no filter aide is necessary.
8-77
-------
TABLE 8.2.4, ECONOMIC COMPARISON - WASTEWATER POLISHING ALTERNATIVES*
Ion Multi- Sorption Preeoat Ultra-
exchange media1 filter filter filter
Water consumption 10-20 103 108 108 108
million gal/yr
Concentrate treatment
Combined rinse treatment
Existing ion exchange
Control
Misc
Capital:
Depreciation (10 years)
Water ($1.87/k gal)
Heat water to 70°Fa
Laborb
Electric ($0.10/kwh)
Media/resinc
Chetnicalsd
Sludge disposal6
Annual operating cost:
$1,000 gallon
$ 150
1,102
250
100
250
$1,852
$ 185
37
0
250
43
220
81
10
$ 826
$ 7.65
$ 150
200
350
75
250
$1,025
* 103
201
378
250
22
68
96
10
$1,128
$ 10.44
$' 150
402
350
75
100
$1,077
$ 108
201
378
250
22
261
50
60
$1,330
4 12.32
$ 150
650
350
75
100
$1,325
$ 133
201
378
250
22
162
96
120
$1,362
fc. 12.61
$ 150
630
350
75
250
$1,455
$ 146
201
378
250
87
98
98
10
$1,268
$ 11.74
S250 Btu/gal, 415/M Btu.
^Supervisor, foreman, and three operators.
cResin life of 3 years, regenerating every day,
dNaBH4 at 4l,500/drum.
e*150/ton.
*A1I values in $1,000.
Source: Reference No, 16.
8-78
-------
Acid Purification Unit: Pickling Liquor Recovery—
Acid purification systems using RFIE have been commercially demonstrated
to be effective in the recovery of acids from aluminum anodizing solutions,
acid pickling liquors, and rack-stripping solutions. Acid purification
systems are the most effective form of ion exchange for recovering acids which
have high concentrations of metal ion contaminants. •
An APU was installed at the Continuous Colour Coat, Ltd, plant in
1 8
ReKdale, Ontario, to recover sulfuric acid from a steel pickling process,
As an alternative to neutralization and disposal of spent baths,' Che plant
employed an APU to remove iron build-up so that the solution could be recycled.
At a flow rate of 19 gal/min and' a temperature of 119 F, iron
concentration in the reclaimed acid was reduced by 80 percent and acidity
&
losses were minimal. Occasional replenishment of the bath was necessary,
but draining of the tank (.an expensive process) was no longer required. Also,
improvements in product quality were noted due to more uniform bath
consistency. Savings were realized in reduced neutralization and disposal
requirements and net reductions in labor requirements for batlr,maintenance.
An economic evaluation of the system (see Table 8,2.5) showed an estimated
payback period for the unit of less than 2 years.
Acid Purification Unit: Aluminum Anodizing Solution Recovery—
Another common application in which the APU works effectively is the
recovery of acids from aluminum anodizing solutions. An APU system was
installed at the Modine Manufacturing Company in Racine, Wisconsin to recover
19
nitric acid from an aluminum etching process. The APU was connected
directly on-line, which allowed for continuous process operation. It
generated a more concentrated solution of recovered nitric acid than it was
fed; thus e slightly lesser volume of acid was returned to the tank.
Table 8.2.6 presents a summary of the results and operating parameters of
the APU at the Mediae plant. Improvements in product quality and savings in
neutralization, disposal, and fresh acid makeup were noted by Modine
19
personnel. Ar. economic evaluation ef this syste~ 'Table S.2.1) shc-vs
nitric acid recovery with an APU to be very cost-effective.
8-?9
-------
TABLE 8.2.5. ECONOMIC EVALUATION OF THE APU INSTALLED AT
CONTINUOUS COLOUR COAT, LTD.
Item
Cost
CAPITAL COSTS3
(includes costs for equipment & installation)
OPERATING COSTS
Resin Replacement
(every 4 years at $58/liter)
Utilities
(0.5 KW x 16 hrs/day x 250 days/yr x J0.055/KWH)
Taxes and Insurance
(1% of TIC)
TOTAL OPERATING .COSTS
COST SAVINGS
Reduction in Acid Purchase § |92.4Q/ton
Reduction in Neutralization Costs (Lime) @ $80/ton
Reduction in Sludge Disposal Costs.
TOTAL COST SAVINGS
NET COST SAVINGS
(Gross Savings - Operating Costs)
PAYBACK PERIOD
$100,000
$3,770/year
$ 110/year
$l,000/year
$4,88Q/year
$25,875/year
$18,OQO/year
$20,OQO/year
S63,875/year
$58,995/year
1.7 years
aCapital Equipment included APU Model No. AP30-24, multimedia sand filter,
water supply tank, and piping.
Source: Reference 9' (icotech cost quote, August 1986).
8-80
-------
TABLE 8.2.6. TYPICAL OPERATING PARAMETERS AND RESULTS DURING TESTING
OF THE APU FOR RECOVERY OF NITRIC ACID AT MODINE
MANUFACTURING COMPANY IN RACINE, WISCONSIN
Parameter Result
Feed to APU from etch tank 6,2 N
Product returned to etch tank 6.5 N
By-product going to waste treatment 0.6 N
Level of aluminum contamination:
Coming into APU from etch tank 793 mg/L
Returning to the etch tank 231 me/L
Average cycle time 12.7 min
Volume of water removed from etch tank/APU cycle 0,89 gal
Mass balance
Equivalents of nitric acid into APU from etch tank 251
Equivalents returned to etch tank and waste 257
Source; References 19 and 20.
6-81
-------
-------
TABLE 8.2.7. ECONOMIC EVALUATION OF THE APU INSTALLED AT MOD1NE
MANUFACTURING COMPANY IN RACINE, WISCONSIN FOR
THE RECOVERY OF NITRIC ACID
Item
Cost
CAPITAL COSTS
(includes costs for equipment and installation)
COST SAVINGS
Reduction in nitric acid purchase
Reduction in neutralization costs
Reduction in disposal costs
Reduction in labor
TOTAL COST SAVINGS
OPERATING COSTS
Resin replacement
(every 4 years at JS8/liter)
Utilities
(0.5 KW x 16 nrs/day x 300 days/yr x $O.Q55/KWH)
Taxes and insurance
(1% of TIC)
TOTAL OPERATING COSTS
NET COST SAVINGS
(Gross savings - Operating costs)
PAYBACK PERIOD
137,234
|2Q,064/year
$ 6,276/year
$ 7,236/year
$ 2,400/year
i35,976/year
$ 1,305/year
$ 132/year
$ 372/year
I 1,809/year
$34,l67/year
1.1 years
Sourcei References 8 and 19 using August 1986 cost data.
8-82
-------
-------
Another full—scale APU demonstration of aluminum anodizing solution
recovery was performed at Springfield Machine and Scamping, Inc. of Warren,
20
Michigan. Typical operating parameters and results during the 6 month
testing period are summarized in Table 8.2.8. The system proved to be
cost-effective for recovery of the sulfuric acid solution due to high aluminum
removal efficiency, retention of acid strength, and reductions in raw material
purchase, disposal and labor costs.
TABLE 8.2.8. TYPICAL OPERATING PARAMETERS AND RESULTS FOR THE APU
INSTALLED AT SPRINGFIELD MACHINE & STAMPING, INC.
IN WARREN, MICHIGAN FOR SULFURIC ACID RECOVERY
Parameter Feed Product Byproduct
Flow rate (liters/hr)
Sulfuric acid concentration (g/L)
Aluminum concentration (g/L)
298
183.8
12.2
296
175.0
4.2
175
13
12
.0
.0
Source: Reference 21.
Acid Purification Unit: Electroplating Pickle Liquor Recovery—
A pilot-scale unit for recovering hydrochloric acid from an
electroplating pickling liquor was tested at Electroplating Engineering, Inc.
in St. Paul, Minnesota. The results were not as successful as with the
cases presented previously since reduced metal removal efficiencies and high
acidity losses were experienced.
primary contaminants in the spent solution included iron (1,650 mg/1),
zinc (4,283 mg/1), and nickel, copper, and chromium in the ppm range. A
different system configuration was required for this application because zinc
was present in the form of chloride complexes. As described in Section 8,2.1,
the resin used in an APU shows a preferential affinity for acid anions as
opposed to metal cations, which causes metals to pass through the resin while
the acid is retained. However, instead of passing thrsugh the resir., sine
chloride complexes are also retained by the anion resin.
8-83
-------
In order to remove both the zinc chloride complex and Che iron
contaminants, it was necessary to operate the system in two stages.
Initially, the spent solution was passed through one resin to remove Bine.
This is termed the inverse mode since the acid ions are not retarded. Then
the solution is passed over a second resin in the normal node of operation,
retaining the acid while allowing iron ions to pass through the resin. As
with typical APU processes, the acid is recovered during the regeneration
cycle. Figures 8.2.14 and 8.2.15 illustrate these'two modes of operation.
Three different HC1 pickling liquors were used to test the performance of
the APU. Analysis of these spent solutions yielded the following:
Parameter i RangeofConcentrations
Acidity
{CaCC>3 equivalents) 77,000 - 284,000 mg/L
Zinc content 640 - 52,000 mg/L
Iron content 1,100 - 7,000 mjj/L
Several test runs were performed using the two-stage APU system, as
summarized in Table 8.2.9. The results show that good zinc removal
efficiencies (99,3 percent) were achieved during the inverse mode of operation
with minimal losses in acidity (3.5 percent). However, during the normal mode
of operation, an average of only 60 percent iron removal was achieved and
acidity losses were high (averaging 38 percent). The results showed that
increased iron removal could only be achieved at the expense of greater
reductions in acidity. It was determined that the ratio of iron to acidity in
the feed had to approach 1:15 in order to achieve effective performance. The
iron to acidity ratio for the feed used during these tests was 1:67, which
contributed to the poor performance results.
It is possible that the intermediate byproduct solution generated after
the inverse mode may be of sufficient quality to be returned to the pickling
bath. Iron content of the byproduct solution was comparable to the iron
concentrations measured in the bath during its intermediate solution stage.
Additional testing would be required in order to determine whether bath
quality would be acceptable under these conditions. Based on the results of
these pilot-scale tests, it was determined that the APU could only be.
cost-effective if a large volume of spent solution is processed.
8-84
-------
STEP ONE - WATER DISPLACEMENT
.•RESIN
IX-.BED
SPENT ACID
Compressed ai r
Spent Acid (SA) displaces
water from resin void
volume.
STEP TWO - INTERMEDIATE BYPRODUCT GENERATION
.Intermed iate .Byproduct (IB) to reservoir
'_ Compressed ai r
Intermediate Byproduct (18) is
SPENT'*,C ID routed to IB Reservoir.
Water r.ecleni shed
'
WATER
RESIN
:%BEC
SPENT AC
Water reservoir is refilled
at beginning of step.
STEP THREE - SPENT ACID DISPLACEMENT
Compressed sir-
.RESIN
•. BED
Water displaces
Spent Acid from resin void
volume.
SPENT ACID
STEP FOUR - INTERMEDIATE WASTE PRODUCT GENERATION
Compresses a i r
,— —
? '
WATER
! 1
i 4
; 1
••RESIN
(" BED
Ll^j
I
r-=— .
SPENT ACID
<
Intermediate Waste Praesuct (IW)
Intermadiale Wasts Product (itf)
is producec.
Spent Acid reservoir is
refilled at beginning
of step.
Spent Acid Replenished
LEEENO
Open va]_ve __^
Closed valve ^_^
Direction of flow
Figure 8.2.14. schematic of invarse mode of operation.
Source: Reference 10.
8-85
-------
STEP ONE - WATER DISPLACEMENT
WATER
t
.RESIN
::>.BED
I
— =• I
INTER-
MEDIATE
BYPRODUCT
(IB)
Compressed air
Intermediate Byproduct (IB)
displaces water from resin
STEP TWO - FINAL WASTE GENERATION
Fina! waste (FW) to final
Water Replenished
~p
WATER
RE
:;.t
i i
SIN
Ep
waste reservoir.
L_Compressed air
— * — i Final Waste Product (FW) is
INTER- routed CO FU reservoir.
MEDIATE
BYPRODUCT Water reservoir is ref filed
STEP THREE - INTERMEDIATE BYPRODUCT DISPLACEMENT
Corap res sect a i r
_
___!
WATER
f.Ri
LJ
:SIN
i.ED
• i <
INTER-
MEPIATE
BYPROPUCT
(IB)
1 I
• p*»"«.
Water displaces Intermediate Byproduct from
STEP FOUR - RECLAIMED ACID GENERATION
Compressed air
-1
• ""•* •
WATER
•RESIN
:>JfD
=i=f
INTER-
HEDIATE
BYPRODUCT
(IB)
i
Rsc fa i med Ac i d
(RA)
Reclaimed Acid (RA) is
Intermediate Byproduct
reservoir is refilled.
Intermediate
Byproduct
Repleni shed
Figure 8,2.15. Schematic of normal mode of operation.
Source: Reference 10.
8-86
-------
TABLE 8.2.9.
SUMMARY OF APU RESULTS ON HCL PICKLING
LIQUOR RECOVERY PERFORMED AT
ELECTROPLATING ENGINEERING, INC.8
Parameter
VOLUME TREATED/ GENERATED (LITERS):
Spent acid
Intermediate by-product
Reclaimed acid
Intermediate waste product
Final waste product
INVERSE MODE LOADINGS TO RESIN:
Zinc (grams/cycle)
Volume (bed volumes/cycle)
Feed rate (liters/hour)
NORMAL MODE LOADINGS TO RESIN:
Acidity (grams CaGOj/cycle)
Volume C beei volumes/cycle)
Feed rate (liters/hour)
STREAM CONCENTRATIONS OF ACIDITY:
(expressed as g/L CaCQ3 equivalents)
Spent acid
Intermediate by-product
Reclaimed acid
Intermediate waste product
Final waste product
STREAM CONCENTRATIONS OF ZINC Cmg/L) :
Spent acid
Intermediate by-product
Reclaimed acid
Intermediate -waste product
Final waste product
Preliminary
runs
(No.2,3,4)b
98
86
70
64
67
34.2
4.05
11.4
43.4
0'.40
6.0
242
217
116
39
90
40,333
34,667
19,233
13,000
7,366
Preliminary
runs
(No.5A,5B,6)c
174
174
162
96
121
4.4
4.05
11.4
35.2
0.40 ,
6.0
156
151
94
15
57
1,100
8
IS
2,000
0.61
Final
runs
(No. 7A-7P)d
189
189
4.6
106
3.3
2.6
4.05
11.4
* 14.6
0.34
5.5
77
.74
46
5.9
30
640
1.4
5,47
1,200
0.25
(continued)
8-87
-------
TABLE 8.2.9 (continued)
Preliminary
runs
Parameter (No.2,3,4)b
Preliminary
runs
(No.5A,5B,6)c
Final
runs
(No. 7A-7P)d
STREAM CONCENTRATIONS OF IRON Cmg/L):
Spent acid 4,700
Intermediate by-product 4,400
Reclaimed acid 1,450
Intermediate waste product 1,277
Final waste product 3,367
STREAM CONCENTRATIONS OF CHROMIUM (mg/L):
2,600
2,433
920
357
1,733
aThe results of Run 1 were discarded due to improper installation.
''The zinc loadings for Runs 2, 3, and 4 were above the recommended
18 grams/cycle maximum loading recommended for the system.
cThe objective for Runs 5A, 5B, and 6 were to process a sufficient
quantity of acid for reuse and to optimize loadings.
dThe objective for Runs 7A through 7P were to optimize loadings for
the normal mode of operation.
Source: Reference No. 10.
1,100
1,100
439
120
728
Spent acid
Intermediate by-product
Reclaimed Acid
Intermediate waste product
Final waste product
43
42
12
4.6
28
2.7
2.6
1.6
0.51
3.4
NA
NA
NA
NA
NA
8-88
-------
In summary, available performance data suggest that the technical and
economic.feasibility of acid purification systems will mainly depend on the
types and concentrations of metal ions present. These systems work well in
recovering solutions with highly positively charged contaminant ions
(e.g., aluminum, iron) because these ions pass rapidly through the strong base
anion exchanger resin. Solutions containing low concentrations of contaminant
ions are not efficiently recovered using the APU. Recommended minimum
concentrations for efficient results are presented in Table 8.2.10. Although
lower concentrations may be treated, removal efficiencies will be low unless
larger systems are employed, compromising cost-effectiveness due to increased
capital costs. A summary of•demonstrated applications is provided in
Table 8.2.11,
8.2,3 Process Costs
An economic evaluation of countercurrent (RPIE) systems is presented in
this section since this is the only ion exchange system which is directly
applicable to the treatmnet of California List metal/cyanide wastes.
Cocurretit flow methods will only be technically and economically feasible for
the treatment of California List wastes which have been diluted by mixing with
spent rinse water.
Factors that affect the costs of RFIE units include: quantity and
quality of constituents recovered, production rates, volume of spent solution
to be treated, concentration of metal salts present in the spent solution,
rate of build-up of metal ions in the bath, concentration of the bath, and
number of hours of process operation.
Capital costs, which include equipment, installation, and peripheral
costs, increase with system size. These costs a're offset by savings which are
realized through reduced volumes of wastes requiring post-treatment
(e.g., neutralisation) and disposal, and reduced purchase requirements for
bath reagents. Operating costs will include replacement of filter cartridges,
resin replacement (approximately every 5 years), and utilities.
8-89
-------
TABLE 8,2.10. RECOMMENDED MINIMUM CONCENTRATIONS (g/L) FOE EFFICIENT
METALS REMOVAL USING THE EGO-TECH APU
Total
Solution Iron Einc Aluminum Copper metals
Hydrochloric acid 30-50 130-150 -
Sulfuric acid 30-50 5 20
Nitric/hydrofluoric acid - - _ - 30
Nitric acid rack stripping - - - 75-100
Note: The APU can be used for solutions with lower concentrations of these
metals, but the metal removal efficiencies will be lower unless a
larger unit is used. Metal removal efficiencies average 55% for
typical systems.
Source: Reference No. 14.
8-90
-------
TABLE 8.2.11. DEMONSTRATED APPLICATIONS OF ECO-TECH ACID
PURIFICATION UNIT USING RFIE
Application/
bath components
Sulfuric acid
Aluminum
Sulfuric acid
Iron
Nitric acid
Nickel and copper
Sulfuric acid
Hydrogen peroxide
Copper
Hydrochloric acid
Iron
Nitric acid
Hydrofluoric acid
Iron
Nickel
Chromium
Sulfuric acid
Sodium
Typical bath
concentration
190
10
127
36
514
99
128
41
13.3
146
34
150
36
29
7.02
7.33
61.3
7.8
Typical product
concentration
(g/L)
182
5.
116
10.
581
47.
113
35
5.
146
25
139
28.
8.
2.
2.
54.
0.
5
5
5
9
8
7
1
2
9
8
Typical by-produc
concentration
(g/L)
13
6
10
21
10
70.
18
7
9.
10
15
4.
7.
20.
4.
5.
5.
5.
8
2
5
2
3
9
1
88
56 '
Source: Reference No. 11 (Based on July 1986 Ecotech cost data).
8-91
-------
Capital costs for acid purification systems typically range from $15,000
to $180,000 depending on the throughput, as shown in Table 8.2.12. These
costs include installation, equipment and -peripherals, and a prefliter
system. Capital costs presented in this table are for the recovery of
sulfuric acid from aluminum anodizing solutions and may be slightly higher for
22
other applications .
Typical operating costs are presented in Table 8.2.13. Finally,
Table 8.2.14 presents an economic evaluation of several hypothetical systems.
8.2.4 Process Status
Cocurrent ion exchange systems are generally not employed for direct
treatment of concentrated metal wastes. Cocurrent systems using weak
exchangers have inefficient exchange capacities for these wastes and are
generally only used as polishing systems following other treatment
operations. Cocurrent systems using strong exchangers are technically
feasible for the treatment of metal-containing rinses and other wastes, but
they are not typically cost-effective because of the high costs for column
regeneration.
Ion exchange systems, using the reverse or reciprocating flow mode
(countercurrent), have been shown to be effective in the treatment of metal
wastes. The process has been demonstrated commercially for chemical recovery
from acid copper, acid zinc, nickel, cobalt, tin, and chromium plating baths,
as well as for purification of spent acid solutions (i.e., the APU).
Chemical recovery systems using fixed bed RFIE have been used to recover
chromic acid and metal salts» It has also been used to deionize mixed-metal
rinse solutions for recovering process water and concentrating the metals for
3
subsequent treatment. Commercial units are available from several vendors.
Acid purification systems using continuous RFIE have been used to remove
aluminum salts from sulfuric acid anodizing solutions, to remove; metals from
nitric and rack-stripping solutions, and to remove metals from sulfuric and
hydrochloric acid pickling solutions. The APU is primarily used for
recovering aluminum anodizing solutions• Acid purification systems are
more cost-effective for retaoving high concentrations of contaminants than
. 8-92
-------
TABLE 8.2.12. TYPICAL CAPITAL COSTS FOR ECO-TECH APD
Item
APU Model No.
Flow rate
Capital cost
Sma 1 1
unit
AP-6
38 L/hr
$14,000
Medium
unit
AP-24
500 L/hr
$37,000
Medi-ani
unit
AP-54
800 L/hr
$116,000
Large
unit
AP-72
6700 L/hr
$184,000
Notes: Capital Costs include equipment, installation, peripherals, and
cartridge-type prefilter system.
Costs presented in this table are for application to recovery of a
sulfuric acid anodizing solution. Costs for other applications may be
slightly higher.
Twelve different size units are available from Eco-Tech, Ltd. The
model numbers, which indicate bed diameters, for these units are:
AP-6, AP-12, AP-18, AP-24, AP-30, AP-36, AP-42, AP-48, AP-54, AP-60,
AP-66, and AP-72.
Source: Reference No. 22 (Ecoteeh quote July 1986).
8-93
-------
TABLE 8.2.13. TYPICAL OPERATING COSTS FOR ACID PURIFICATION USING
CONTINUOUS COUNTERCURRENT ION EXCHANGE (RFIE)
Item
Cost
Filter cartridges for prefilter system
Utilities:
(0.5 KW x 16 hrs/day x 20 days/month .
^x 0.055 $/KWH)
Resin replacement
(specific cost depends on system size)
&lO.OO/month
$8.80/montb
$58/liter every 4 years
Source: References 8 and 10 (Based on August 1986 cost data).
8-94
-------
TABLE 8.2.14. ECONOMIC EVALUATION OF ACID PURIFICATION PROCESS
30,000 gpy 100,000 gpy 500,000 gpy
Description throughput throughput throughput
Case 1 - Purification of Sulfuric Acid Anodizing Solution: Previous
approach used caustic acid neutralisation- New approach uses
APU with caustic neutralisation.
Appro*. APU Cost 4 6,000 $11,000 ' $ 25,000
Previous treatment cost 4 9,690 $32,300 4161,500
Previous acid cost 4 2,349 4 7,830 S 39,150
Annual ssvings 4 8,427 $28,891 $140,455
Payback (months) S 5 2
Case 2 - Purification of Sulfuric Acid Anodizing Solution: Previous
approach used lime neutralization: New approach uses APU with
lime neutralisation.
Approx. APD cost 4 6,000 411,000 A 25,000
' Previous treatment coet 5 2,250 S 8,500 i 36,500
Previous acid coat 4 2,345 & 7,830 S 39,150
Annual savings J 3,216 410,731 4 53,655
Payback (months) 22 12 ,5
3 ~~ Purification of Sulfuric Acid Anodisirtg Solution: Previous
approach used waste haulage: New approach uses AFU wich
caustic neutralization.
Approx. APU cost & 6,000 $11,000 ' $ 25,000
Previous treatment cost " 4 3,000 510,000 $ "50,000
Previous acid cost S 2,349 S 7,830 S 39,150
Present treatment cost 4 2,90? S S,690 4 48,450
Annual ssvings S 1,737 S 3,791 $ 28,955
Payback (months) 61 23 10
Case A - Purification of Sulfuric Acid Anodizing Solution: Previous approach
used waste haulage: New approach uses APU with lime neutralisation.
Approx. APD cost $6,000 411,000 4 25,000
Previous treatment cost $3,000 410,000 & 50,000s
Previous acid cost 42,349 i 7,830 feS 39,150
Present treatment cost - 4 675 i 2,250 S 11,250
Annual savings 43,969 413,245 S 66,155
Payback (months) 18 10 5
Case 5 ~ Nitric Acid Recovery: Previous approach u-sed caustic neutralisation:
New approach uses APU with caustic neutralisation.
Apprpx. APU cast
Previous treatment cost
Previous ' ac id cost
Totsl previous cose
Annual savings
Payback (months)
49,400
£?,575
48,775
i!6,3SO
S- 9,810
1 i
511,300
430,300
435,100
465,400
439,240
3
S IS',400
4 50,500
4 58,500
4109,000
4 65,400
3
-Scares: References 11, 12, 22, sr.d 21. (Eased or, July 15S6 Ecocac'h cooC dat
3-95
-------
other ion exchange systems. Although the use of ion exchange for acid
purification is currently under investigation by several ion exchange vendors
(e.g., Alpha Process Systems; Illinois Water Treatment Company; Ionics, Inc.;
etc.)» Eco-Teeh, Ltd. is the only vendor with commercial units currently in
10,22,24
operation.
8-96
-------
REFERENCES
1. Wilk, Lisa et al. Alliance Technologies Inc. Technical Resource
Document: Treatment Technologies for Corrosive-Containing Wastes.
Prepared for U.S. EPA HWERL, Cincinnati, OH. under Contract
No. 68-02-3997. October, 1986.
2. GCA Technology. Industrial Waste Management Alternatives And Their
Associated Technologies/Processes. Prepared for the Illinois
Environmental Protection Agency, Division of Land Pollution Control,
Springfield, Illinois. GCA Contract No. 2-053-C11 and 2-053-012.
GCA-TR-80-80-G. February 1981.
3. U.S. EPA, Industrial Environmental Research Laboratory, Cincinnati,
Ohio. Summary Report: Control and Treatment Technology for the Metal
Finishing Industry - Ion Exchange. EPA-625-8-81-007. June 1981.
4. Yeats, A,R. Ion Exchange Selectively Removes Heavy Metals From Mixed
Plating Wastes. In: Ind. Waste Cont. Proc, 32, 467-76, 1978.
5. Waitz, Jr., W.H., Rohm & Haas Company. Ion Exchange in Heavy Metals
Removal and Recovery. Amber-hi-lites No. 162. 1979.
6. U.S. EPA, Industrial Environmental Research Laboratory, Cincinnati,
Ohio. Sources and Treatment of Wastewater in the Nonferrous Metals
Industry. EPA-6QQ/2-80-074, April 1980.
7. Weber, W.J. Physicochemical Processes for Water Quality Control.
John Wiley & Sons. 1972.
8. Fontana, C., Eco-Tech, Ltd. -Telephone Conversation with L. Wilk,
GCA Technology Division, Inc. Re: Acid Purification Unit.
August 21, 1986.
9. Dejakj M. Acid Recovery Proves Viable in Steel Pickling. Finishings
10(1): 24-27.' January 1986.
10. Pace Laboratories, Inc. Final Report: Reclamation and Reuse of Spent
Hydrochloric Acid, Hazardous Waste Reduction Grant. Prepared'for the
Minnesota Waste Management Board on behalf of Electro-Plating Engineering
Company, Inc. February 14, 1986.
11. Eco-Tech, Ltd. Product Literature; Acid Purification Unit (APU).
Bulletin No. ET-4-84-5M, Received July 1986.
12. Eco-Tech, Ltd. Product Literature: Ion Exchange Systems. Bulletin
No. ET-11-83-3M. Received July 1986.
8-97
-------
13. GCA Technology. Corrective Measures for Releases to Ground Water from
Solid Waste Management Units. Prepared for U.S. EPA-OSW Land Disposal
Branch, under EPA Contract No. 68-01-6871, Work Assignment No. 51.
GCA-TR-85-69-G. August 1985.
14. Fontana, C., Eco-Tech, Ltd. Telephone Conversation with L. Wilk,
GCA Technology Division, Inc. Ee; Acid Purification Unit.
August 26, 1986.
15. U.S. EPA, Office of Research and Development, Washington, B.C.
Treatability Manual, Volume HI: Technologies for Control/Removal of
Pollutants. EPA-600-8-80-042e. July 1980.
16. Van Dyke, Jr., B.H., Gonoby, J.F., and C. Alderuccio. Innovative
Hazardous Waste Stream Reduction Alternatives. In: Proc. of the Third
Annual Hazardous Materials Management Conference. June 1985.
17. Rohm & Haas Company. Duolite GT-73 Ion Exchange Resin Product Bulletin.
August 1986.
18. Chemical Processing Staff. Spotlight: Pickling Acid Recovery Unit Saves
$40,000/year, Purifies Spent Sulfuric Acid. Chemical Processing,
49(3): 36-38. March 1986.
19. Robertson, W.M., James, C.E., and J.Y.C, Huang. Recovery and Reuse of
Waste Nitric Acid From An Aluminum Etch Process. In: Proceedings of the
35th Industrial Waste Conference at Purdue University. May 13-15, 1980.
20. Brown, C.J., Davy, D., and P.J. Simmons. Recovery of Nitric Acid from
Solutions Used for Treating Metal Surfaces. Plating and Surface
Finishing. February 1980.
21. Brown, C.J., Davy, D., and P.J. Simmons. Purification of Sulfuric Acid
Anodizing Solutions. Plating and Surface Finishing. January 1979.
22. Fontana,, C., Eco-Tech, Ltd. Telephone Conversation with L. Wilk,
GCA Technology Division, Inc. Re: Acid Purification Unit. July 7, 1986.
23. Parcy, E., Ionics, Inc. Telephone conversation with J. Spielman,
GCA Technology, Inc. August 14, 1986.
24. Jain, S.M., Ionics, Inc. Telephone conversation with J. Spielman, GCA
Technology Division, Inc. August 12, 1986.
8-98
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'8.3 DEVOE-HOLBEIN TECHNOLOGY
8.3.1 Process Description
DeVoe-Holbein Technology uses coordinating compounds covalently hooded to
the surface of an inert carrier material to capture metal ions. In waste
treatment applications, the reactants are used in equipment similar to that
employed for ion exchange resins.
The technology was originally developed by DeVoe-Holbein as an adaptation
of biological mechanisms in which living cells selectively extract a variety ••
of metal nutrients (e.g., Na, K, Mg, Ca, Cu, Zn, Co, Fe, Se, and Mn) from
their environment. Cells can acquire target metals by means of specialized
molecular sites on their surfaces that recognize and bind only that species,
Examples of such selective reactants are the nonprotein iron-binding
2
molecules, collectively known as siderophores.
Siderophores generally fall into two classes of molecules, hydroxomates
and phenolate-catecholates. DeVoe-Holbein covalently linked microbial"
siderophores which belong to each of the two classes, Entetobactin and
Desferrioxamine (see Figures 8.3.1 and 8.3,2), to porous glass bead
supports. In subsequent biological experiments, these particulate
compositions were used successfully as a fixed-bed, iron-retrieval system,
DeVoe-Holbein has since synthesized a series of metal-capturing
compositions with catechol, or substituted catechols, as the active
component. Such compositions have similar properties to those of
Enterobactin. Catechol was covalently bound to solid surfaces with
bifunctional linking agents of defined lengths (Figure 8.3.3), Hiehlv porous
glass is the solid substrate which has been found to be most practical for the
composition synthesis.
According to DeVoe-Holbein, the resulting compositions proved to be
highly efficient, typically achieving 99 percent or higher removal rates, and
are selective for individual or groups of metals. The rapid adsorption
kinetics minimizes required contact time and the compounds are mechanically
and chemically stable. In addition, the compositions are regenerable,
3
8-99
-------
Enterobactin
o-c
HCK
Figure 8.3.1. Microbial siderophore Enterobactin. Reference I.
H
I
H —N
\
"sJs
N-C
1 II
-0 O
X V
Linear ferrloxamine
CONH CONH
/ \
(CH,)3 (CH,).
\ /
N-C
I 11
-0 O
CH,
I
N-C
1 II
-O O
Figure 8.3.2. Microbial siderophore Ferrioxamine« Reference 1.
8-100
-------
Figure 8.3.3.
Graphic display of metal-capturing composition. Fork~like
symbols represent siderophores iramobilized through bridging
ager.ts to a solid surface represented by the continuous
line. Reference 1.
8-101
-------
The synthesized compounds were employed for waste treatment applications
4
in a manner similar to classical ion exchange, as shown in Figure 8.3.4.
The media is contained in a fixed bed, and the metal-laden solution is passed
through the bed during the service cycle. Following saturation of the media
with metals, the bed is backwaahed and the bound metal is displaced by an
4
appropriate regenerant; e.g., 2 N HCi. Co-current and counter-current
fixed bed systems have been developed. The basic modular system which is now
commercially available, can be expanded or realigned to correspond to end
users' varying throughput requirements and spatial limitations.
DeVoe-Holbein adsorption units are only able to treat contaminants in
solution. Similar to ion exchange, high concentrations of suspended solids
which can foul the adsorbent bed are typically pretreated through some form of
filtration. Uaste streams from the adsorption process include: contaminated
regenerant and filtered solids from the pretreatment system. Filtrate from
the pre-filtering system can generally be land disposed without further
treatment. The regenerant may require treatment (e.g., neutralization,
precipitation, dewatering) and disposal if not amenable to recycling.
One of the reported advantages of the DeVoe-Holbein system is that it is
capable of yielding a more highly concentrated regenermnt than ion exchange.
Several options for downstream utilization of the concentrated metal
regenerant are therefore possible. When it is compatible with the parent
solution bath and metal concentrations are sufficiently high, the regenerant
stream may be reused directly. If higher metal concentrations are required,
an intermediate recovery step can be employed. For example, metal may be
recovered from the regenerant electrolytically, recycling the regenerant to
the adsorption process and selling the metal as scrap.
8.3.2 Process Performance
The performance of a DeVoe-Holbein system will be influenced by the
characteristics and quantity of the waste stream being treated. Parameters
which need to be considered when evaluating the applicability of the system
for a particular waste streaa include: rypes and concentrations of
constituents present in the waste stream, required effluent quality, ,and
8-102
-------
WASTEWATER
T
OO
I—1
o
fDTIS:
1) OM.Y 3 BED VOUJ-t OF DILUTH)
f^GQIT f€QUlPH) TO REGOOATE
DEVC£-HOLBEIN CO'Y^OSITION.
2) FRESH WATER RINSE CMSES OUT
CONCENTTWTH) I-ETAL TO RECOVERY
FOR REUSE.
80*
RETURN
TO TAW
FOR-REUSE
CONCENTRATED
METAL
TREATED
WATER
Figure
Wastewater process unit column showing regeneration. Reference
-------
options for managing the regenerant stream. DeVoe-Holbein says its
compositions, trademarked Vitrokele, meet the following criteria;
o The ability to capture all or virtually all of a specific, target
metal, even in the presence of very low concentrations of that metal
or in the presence of competing metals.
o The ability to withstand harsh physical and chemical treatment
without losing structural or functional integrity,
o The ability to allow eaay displacement of the metal, permitting
metal concentration and regenerant solution volume reduction, reuse
of the composition and, possibly, reuse of the captured metal,
o The ability to capture substantial quantities of metal per unit of
composition while maintaining high capture efficiency.
o The use of non-toxic agents; i.e., the process will not add trace
toxic components to the solution from which the netal is being
captured.
o The capability for being produced at a low cost, enhanced further by
regenerabiiity.
DeVoe-Holbein compositions all display very high metal capture
3
efficiencies. Table 8.3.1 illustrates the high capture efficiencies
obtainable in the laboratory with test metal solutions of importance to metal
finishing and hydro-metallurgical operations. High capture efficiencies
are demonstrated up to the capacity of the particular compound, with a sharp
breakthrough curve occurring after saturation. A typical breakthrough curve
is shown in Figure 8.3.5, where the DeVoe-Holbein DH-520 has been used for Cu
removal from a relatively concentrated metal solution.
Selectivity and specificity of a particular composition are, in part,
functions of the pretreatments used with the particular composition, and the
conditions under which the metal solution or wastewater are treated. In many
instances, selectivity and specificity can be altered (broadened or narrowed)
4 6
to meet specific requirements of metal extraction, ' Figure 8,3.6
demonstrates the selectivity o£ DeVoe-Holbein composition DH-506 (F-l) for
iron relative to sodium, cadmium, and cobalt.
8-104
-------
TABLE 8.3.1. THE EFFICIENT CAPTURE OF SOME TOXIC HEAVY METALS OF
IMPORTANCE TO THE HYDROMETALLURGICAL AND METAL
FINISHING INDUSTRIES BY DEVOE-HOLBEIN COMPOSITION
loxic
metals
Cadmium
Chromium
Copper
Lead
Mercury
Nickel
Zinc
DeVoe-Holbein Influent
composition concentration
DH-516
DH-524
DH-520
DH-501
DH-573
DH-507
DH-508
674 pptti
694 ppm
38 ppm
42 ppm
12 ppm
0. 10 pptn
6 . 5 ppm
Effluent
concentration8
<1.0
<0.01
<1.0
< 42 . 0
<1.0
<1. 5
<0.8
ppb
pptn
ppt
ppb
ppb
ppt
ppb
Capture
efficiency (X)b
^99.99
•>« O 0 Q Q
>_99.99
>_99,99
>_99.99
>99.99
>99.99
aEffluent concentration at or below normal detection limits using either
radioactive tracer or atomic absorption spectrophotometric determinations,
"Capture efficiency determined as percent reduction in influent concentra-
tion; values.are greater or equal to those shown due to detection limits
of effluent metal concentration.
Reference 3.
8-105
-------
280 i
_ 240-
i •
a
— 2004
| .160-
,** •
I 120-
2
o
o
80-
40-
0 40 80 120 160 200 240 280
Volume (ml) of Cu solution (2SO opm)
applied to 4s composition
Figure 8.3.5. Cbpper removal breakthrough curve. Reference 1.
20
40 60
Effluent Volume (mi)
100
Figure 8.3.6. Selectivity of DH-506 (F-l) for iron. Reference 3, 6.
8-106
-------
The specificity of DeVoe-Holbein compounds can be demonstrated by the
capture of a test metal from a complex solution containing a'number of other
metals. Figure 8.3.7 demonstrates the ability of DeVoe-Holbein composition
506 (F—I) to selectively remove iron from sea water. The figure shows
differences between the highly selective DeVoe-Holbein composition and
competitive adsorbatits, in this case, a selective ion exchange resin and a
. . , 3
strong cationic exchange resin.
Resin regenerability is of prime importance in determining overall
economic viability of metals adsorption processes. Figure 8.3.8 illustrates
the regetierability of DeVoe—Holbein compositions with a comparable cation
exchange resin. This test was performed with identical bed sizes and similar
materials were captured. Regeneration of the cation exchange resin required
nearly five times the regenerant solution volume to achieve comparable
regenerat ion.
Case studies of DeVoe-Holbein adsorption applied to metal/cyanide wastes, as
adapted front DeVoe-Holbein, are presented below.
Case Study #1—
DeVoe-Holbein technology was evaluated for removal of chromium from three
representative chromium wastewater streams wastewaters: boiler blowdown
water, chrome plating waste precipitator effluent, and cooling tower
wastewater. Waste stream metal concentrations are summarized in Table 8.3.2,
DeVoe-Holbein composition DH-524 was used for'this analysis.
High recovery efficiency was demonstrated by using composition DH-524, as
shown in Table 8.3.3. Both chromium removal and regeneration efficiency were
essentially complete. In addition, the regenerant volume reqired was only 3
bed volumes and showed no loss in efficiency over 15 cycles.
Case Study #2—
The DeVoe—Holbein treatment process was employed to treat a 5 gpm
counter-current rinse effluent from a zinc chloride electroplating line. It
demonstrated high zinc removal efficiencies in the presence of other cations.
8-10?
-------
^ 100-,
o
2
2 80-
e «
o =
^2 C
a —
60-
40-
o rf
|-
S 20
20
40
60
80
100
59,
Volume (ml) Fs solution
Figure 8,3.7.
Demonstrated ability of DeVoe-Holbeln DH-506
-------
TABLE 8.3.2. PARTIAL ANALYSES OF DIFFERENT CHROMIUM WASTEWATERS
Element3
Chromium (VI)
Sodium
Calcium
Magnesium
Silicon
Strontium
Zinc
Boron
Iron
Boiler
blowdown water
(ppm)
694.3
432.0
155.3
64
5.2
0.8
9.9
0.4
—
Chrome plating"
waste effluent
(ppis)
3.8
14.2
88.4
6.8
10.6
0.3
0.1
—
—
Cooling tower bleed
(pptn)
7.6
57.4
176.4
42.1
35.2
0.9
0.8
0.3
0.5
aElemental analyses carried out by Inductively Coupled Plasma Emission
Spec ttotne try.
^Wastewater following conventional chemical reduction and precipitation
of chrome plating rinse water.
Reference 3.
8-109
-------
TABLE 8.3.3. RECOVERY OF Cr FROM VARIOUS WASTEWATERS WITH DH-524
Type of
wastewater
Boiler blowdown
Plating waste
precipitator
effluent
Cooling tower
Influent
CR VI
(ppm)
694.3
3.8
7.6
Treated3
effluent Cr VI
(ppm)
NDC
ND
ND
Cr removal
efficiency (%)
>99.99
>99.99
>99.99
Capacity mg Cr VI
per kilogram
composition
~ 20,000
^ 20,000
~ 20,000
Regeneration
efficiency (%)b
100
100
100
aFlow rate of 20 bed volumes/hour in a fixed bed of DH-524.
percent of bound Cr displaced in approximately 3 bed volumes of regenerant.
Fully regenerable over 15 cycles of use so far tested.
CND = not detectable by atomic absorption spectrophotometry .
Reference 3.
-------
The wastewater process was designed to operate on a 16 hour feed cycle, with
the influent zinc concentration ranging from 50 to 300 ppm. At the end of the
operating period, the process unit is regenerated and1 reconditioned. Figure
8.3.9 shows the results of the operation of the wastewater process unit over a
2-week period. Depending on plating activity, the inlet zinc concentration to
the process unit varied significantlv, from 10 ppm to as hieb as 280 ppm. Yet
zinc concentration in the treated effluent consistently remained below
1 ppm.
The adsorption unit was regenerated daily followinR the 16 hour
processing period. Less than one—third of a bed volume of regenerant was
applied, at a flow rate of 0.5 bed volumes/hour,, followed by a similar volume
of rinsewater. The regenerant, with typical metal concentrations of 50,000
ppm (as high as 100,000 ppm) is directed to a storase tank for further
3
recovery.
Case Study #3—
In another example, a large job shop operating four different processes,
and using at least eight different metals, recently installed a DeVoe-HoLbein
treatment system. Prior to this, spent baths, acids and soaps were sent to a
landfill and aqueous effluents were pH adjusted in a neutralization pit and
discharged directly to the sewer.
The new wastewater process successfully treats both the individual and
combined wastewater effluents which contain nickel, zinc, brass, chrome,
precious metals, and possibly cyanide. Although combined effluents are
treated efficiently, treatment of individual rinse lines offers several
advantages over treatment of the combined effluent. Smaller treatment systems
are required for individual rinsewater effluents, and these can be operated in
4
a closed loop cycle, reusing treated rinsewater and recovered metals. To
demonstrate the efficiency of the DeVoe-Holbein wastewater treatment system, a
treatment study of the nickel rinse effluent from an automatic rack plating
operation was undertaken.
3-111
-------
00
I
300i
200-
£X
Q.
100-
c
M
>. ^
10-
0:
M
I I
I /
I /
» /
I /
I '
\/
\\
30 60 90 120
Elapsed Process Time (h)
150
' i
'. f' \
180
WASTEWATER INFLUENT (•--•), TREATED EFFLUENT (••—•)
Figure 8.3.9. Removal of Zn from electroplating wastewater. Reference 4.
-------
Figure 8.3.10 shows the results of operation of the adsorption
treatment system over two complete cycles of loading and regeneration. The
nickel concentration in the rinsewater was reduced from 130 to 180 ppm in the
feed to an average of less than 1 ppm nickel in the treated effluent. When
the nickel concentration in the effluent reached 5 ppm, the unit was
regenerated. Less than two- bed volumes of regenerant was required, resulting
4
in a stream that was highly concentrated in nickel (7,000 to 8,000 ppm).
Case Study #4—
For cyanide complexes, DeVoe-Holbein media have demonstrated high
efficiency but, in some cases, result in relatively low capacity. This has
been successfully overcome by pretreating cyanides using a destruct process.
For example, following cyanide oxidation with chlorine, both cadmium and zinc
were efficiently captured (99 percent) with capacities in excess of 12 grams
of ainc or cadmium per liter of DeVoe-Holbein media.
Silver-contaminated effluents containing thiosulfate or cyanide complexes
do not require destruction of the cyanide prior to extraction of silver.
Using a selective and regenerable DH media, silver can be efficiently
(99.99 percent) removed from solutions to concentrations Less than 10 ppb.
Although -the capacity of this medium for silver is somewhat dependent on the
nature of the effluent, capacities as high as 20 grams of silver per liter of
media have been realized,
8.3.3. Process ^Costs
DeVoe-Holbein offers several different VITROK.ELE compositions, each of
varying selectivity and metal capture capacity (depending upon the environment
in which the metal must be captured and recovered). The costs associated
with DeVoe-Holbein treatment systems is subject to the particular VITROKELE
composition sought and the volume ordered, and total system costs are assessed
9
on an individual basis. As a general guide per liter of D-H composition,
the prices range from $10 to $50. System costs, however, were
unobtainable in conversations with DeVoe-Holbein representatives.
3-113
-------
REGENERATION
LOAD-t
JRINSE||RINSE|
LOAD-2
REGENERATION
JRINSEHRINSEI LOAD-3
00
i
8OOO-
£EGOOO-
CL
Q4OOO-
— -»
Z 2OOO-
2O-
'
1
2O 4O 6O 8O 1OO 12O
bed volumes
Figure 8.3.12. Removal of Ni from electroplating rinsewater, and regeneration of wastewater
treatment unit.
Source: Reference 4.
-------
8.3.4. Process Status
DeVoe-Holbeia technology is protected under U.S. Patent No. 4,530,963 and
a number of pending patent applications throughout the world. DeVoe-Holbein
International N.V. holds the worldwide rights to the technology and is
commercializing various aspects through subsidiaries and joint ventures.
The process appears most applicable to the selective removal of valuable
metals (e.g. silver) from waste streams. Although the process appears
promising, further information concerning selectivity and capacity is! needed.
8-115
-------
REFERENCES
1. DeVoe, I.W., and B.E. Holbein. "A New Generation of Solid-state Metal
Cofflplexing Materials: Models and Insights Derived from Biological
Systems." In: Newer Methods for the Removal of Trace Metals from
Aqueous Solution. Amual. Chem. Congress, Royal Society of Chemistry,
University of Warwick, England. April 1986.
2. J.B. Neilands. Ann. Rev. Biochem. 50:715. 1981,
3. Holbein, B.E,, DeVoe, I.W., Neirick, L.G., Nathan, M.F., and
R.N. Arzonetti. "DeVoe-Holbein Technology: New Technology for
Closed-Loop Source Reduction of Toxic Heavy Metal Wastes in the Nuclear
and Metal Finishing Industries. Presented at: Massachusetts Hazardous
Waste Source Reduction Conference Proceedings.- October 17, 1984.
4. Brener, D., Greer, C.W., and S. G. Nelson. "Novel, Synthetic Compounds
for the Efficient Removal of Nickel and Zinc Prom Plating Effluent."
Presented at: 7th American Electroplaters Society, U.S. Environmental
Protection Agency Conference on Pollution Control for the Metal Finishing
Industry. 1986.
5. Greer, C. W. and A« C« Huber. Synthetic Compositions: Effective
Treatment of Heavy Metal Containing Wastewaters from the Metal Finishing
Industry. DeVoe-Eolbein Canada, Inc., Montreal, Quebec, Canada.
6. DeVoe, I.W., and B.E. Holbein. "New Technology for the High Affinity
Capture of Radioactive Metals from Water. Presented: 4th Annual
Conference Canadian Nuclear Society. June 1983.
7. Neirick, L.G., and B.E, Holbein. "Removal of Heavy Metals from Waste
Streams with Novel High-Affinity Selective and Regenerable Media.
Presented at: Annual Meeting American Electroplatera Society, Detroit,
Michigan. 1985.
8. DeVoe-Holbein Technical Brochure.
9. 0. D'Sylva. Telephone conversation with David Sullivan, Alliance
Technologies Corporation. May 4, 1987.
10. T. Resch, Devoe-Holbein. Personal communication with David Sullivan,
Alliance Technologies. April 1, 1987.
11. U.S. Patent No. 4,530,963. Insoluble Chelating Compositions, I.W. DeVoe
and E. Holbein. July 23, 1985.
8-116
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SECTION 9.0
ELECTROLYTIC PROCESSES
9.1 PROCESS DESCRIPTION
The electrolytic cell is the basic device used in electroplating
operations. The cell consists of an anode and a cathode immersed in an
electrolyte. When current is applied, dissolved metals in the electrolyte are
reduced and deposited on the cathode. This process is attractive for
pollution control because of its ability to remove specific contaminants from
the waste stream without the addition of chemicals which produce large
quantities of sludge. In addition, it is often possible to reuse the metal
which is removed from solution, thereby making the technology a recovery
1 2
process as opposed to an end-of-pipe treatment process. '
A commonly used configuration for electrolytic recovery is to connect the
electrolytic unit to the dragout tank that follows metal plating or etching
baths and precedes the running rinse (see Figure 9.1.1). The solution in the
dragout tank, which contains diluted plating chemicals, is circulated through
the electrolytic reactor and hack into the dragout tank. In this way, the
concentration of metals in the dragout tank is maintained at a low level.
Instead of being carried into the running rinse and eventually into the
wastewater treatment system, the metals are recovered by the electrolytic
reactor.
Electrolytic treatment is not effective in removing all contaminants. It
is most effective in removing the noble metals such as gold and silver. These
metals have^igh electrode potentials (see Table 9.1.1) and are easily reduced
and deposited on the cathode. Other metals, such as aluminum and magnesium,
cannot be removed by this type of process because their electrode potentials
favor oxidation rather than reduction. Compounds such as cadmium, tin, lead
and copper can be removed, but a greater amount of current is required,
particularly when the metal concentration is low; e.p. less than 1,000 ppm.
In additzc" tc t^s type cf tr.Etal, ths type of solution slso has sn effect
on the practicality of ectrolytic recovery. Extremely corrosive solutions
(e.g., certain etcnants) may pose problems for electrolytic recovery because
9-1
-------
PLATING BATH
DRAGOUT
TANK
RUNNING RINSE
WORK
VO
to
ELECTROLYTIC
CELL
TO
DRAIN
RINSE
WATER
Figure 9.1.1. Typical electrolytic recovery system configuration.
-------
TABLE 9.1.1. ELECTRODE POTENTIALS AT 25°C
Metal
Gold
Platinum
Silver
Copper
Lead
Tin
Nickel
Cadmium
Steel or Iron
Zinc
Aluminum
Magnesium
Cathode
Aur * 3e-
Pt^* 4- 2e-
Ag •*- e-
Cu2* + 2e-
Pb2* + 2e-
Sn2+ + 2e-
(j£2+ + 2e-
Cd2+ + 2e-
Fe24 + 2e-
Zn2+ 4- 2e-
Al''+ + 3e-
Mg2"1" + 2e-
reaction
Au(e)
Pt(e)
Ag(s)
CuCa)
Pb(s)
Sr(s)
Ni(s)
Cd(s) ,
Fe(s)
Zn(s)
Al(s)
MgCs)
Electrode potential
(volts)
1.5
1.42
0.8
0.345
-0.126
-0.136
-0.25
-0.40
-0.44
-0.76
-1.68
-2.37
Source: Reference No.
9-3
-------
the metal that is plated on the cathode is etched off as quickly as it is
plated. In addition, solutions with chelated metals, such as electroless
copper plating solutions, may be more difficult for electrolytic recovery than
solutions containing free metal ions such as acid copper electroplating
solutions.
For dilute metal~eontaining solutions, electrolytic recovery can be
extremely difficult, particularly when using standard flat plate electrodes.
One of the primary limitations of this type of electrode is that high
mass-transfer rates are difficult to achieve. When plating metals from a
solution, the layer of solution next to the cathode becomes depleted in metal
ions. Since there are fewer ions present in dilute solutions, diffusion into
and across the depleted layer is much slower and the layer becomes thicker- and
more depleted . Mass transfer rates can be enhanced both by agitation and
by increasing the effective surface area of the electrodes, particularly the
cathode. Both of these actions will increase the rate of movement of metal
ions to the cathode, which is'equivalent to an increase in the current passed
between the electrodes.
Since most rinsewaters requiring treatment contain metals at
concentrations of less than 1,000 ppm, a number of electrolytic reactors have
been designed with electrodes that either enhance mixing or have large surface
areas. Some of the electrode designs are:
* Concentric cylinder;
• Parallel, porous plates;
* Rotating cylinder;
* Packed bed;
* Fluidized-bed; and
* Carbon fiber.
The electrodes used in these reactors may be more effective in removing
metals from solution, but their design may also make it difficult to remove
the metal once it has been plated onto the cathode. For example, the use of a
reactor with parallel stainless steel cathodes generally allows for the
9-4
-------
production of a compact:, adherent layer of metal. This can be mechanically
removed and sold as scrap. Conversely, the use of a reactor with a high
electrode area results in the deposition of metal within pores of a cathode
which may be comprised of carbon fibre, carbon granules, metal mesh, or metal
sponge. In this case, mechanical removal of the metal is generally not
feasible. Therefore, recovery o£ the metal must be accomplished by leaching
the deposited metal out of the cell as a concentrate, by either corrosion or
anodic dissolution. Alternatively, the cathode material may be disposed, or
in the case of precious metals, sent to a refiner.
Both of these methods have drawbacks. Disposal does not allow reuse of
the cathode, and leaching may not be practical in all situations. For
example, precious metals such as silver, gold, arid platinum are difficult to
dissolve by corrosion, and aggressive solutions may damage cell components.
Anodic dissolution involves reversing the polarity of the electrolytic cell
which may also damage electrode materials. Therefore, increasing mass
transfer by using high surface area electrodes may come at the expense of
reuse of the deposited metal and/or cathode.
Electrolytic cells have also been used to treat plating solutions and
rinsewaters containing cyanide. In this case, cyanide is oxidized at the
anode forming cyanate as an intermediate product and carbon dioxide, nitrogen,
and ammonia as the end products. In some cases, it is even possible to
concurrently deposit metals at the cathode and oxidize cyanide at the anode.
Optimal destruction usually requires temperatures between 150 and 200"P.
Sodium chloride can be added to provide a source of chlorine which also acts
7 B
as an oxidant to enhance cyanide destruction. * However, with conventional
reactors, it is difficult to treat solutions containing low concentrations of
cyanide; e.g., less than 1,000 mg/L. Therefore, another treatment process,
such as alkaline chlorination, is often used after electrolytic oxidation to
destroy residual cyanide.
A pilot scale electrolytic device called a Trickle Tower Electrochemical
Reactor, has been developed which is reportedly capable of acheiving low
9
cyanide levels. The column is comprised of alternating layers of an
electrolytically conductive packing (e.g., carbon Rasehig rings) separated by
Voltage apDiied to plane eiectrooes at eitner
-------
end of the column results in each of the conductive layers of the column
becoming a bipolar electrode; i.e. one face positive and the other negative,
leaving a neutral center zone. The cower is regenerated by filling with
anodic solution and reversing the polarity. The concentrated metal solution
which deveiopes can be recycled to the plating bath.
Copper is catalytic for the oxidation of cyanide due to the intervention
of Che Cu (CN ) complex in the solution phase. Cyanide is oxidized as
it passes the anodes and, when excess cyanide has been eliminated, copper is
deposited on the cathodic surfaces of the bipolar layers. Cyanide levels have
been reduced from over 200 ppra to near zero in 60 minutes during bath
treatment of copper/cyanide solutions. In comparison, electrolytic oxidation
9
of free cyanides (i.e., without metals present) takes roughly twice as long.
9.2 PRETREATMENT AND POST-TREATMENT REQUIREMENTS
Electrolytic processes are generally used at the source of waste
generation. The aqueous effluent is then either reused directly (e.g., bath
make-up) or further treated to remove other contaminants or to be
neutralized. Therefore electrolytic recovery is itself somewhat of a
pretreatment process. However, in many cases, it is necessary to filter the
8 10
wsstewater prior to feeding it through the electrolytic reactor. ' This
is particularly true with reactors that utilize porous or packed bed
electrodes since particulates can potentially clog the reactor.
Adjustment of pH is also sometimes necessary as a pretreatment measure
since the waste pH affects metal special: ion. At a low pH, free metal ions
predominate. These exhibit a higher mass—transfer rate to the cathode than do
metals at higher pH. However, when treating wastewater containing chelated
metals, the pH of the solution will not have a significant effect unless pH
goes below 3.
As discussed previously, post-treatment may be required to recover metals
from the cathodic repenerent solution if this cannot be reused directly ss
bath make-up. Recovered wastewater will also require eventual disposal or
treatment due to build-up of organics and other impurities present in the
bath. Finally, stripped metals and cetal Isdea caihcdss car. be shipped •
offsite to smelters or reclamation facilities.
9-6
-------
-9,3 PERFORMANCE OF ELECTROLYTIC RECOVERY SYSTEMS
The performance of several electrolytic reactors on specific
metal/cyanide wastes^is summarized in Table 9.3.1. However, these data
reflect performance of a particular electrolytic reactor on a specific waste
stream and, thus, should not be taken as a general indicator of performance.
Performance can be assessed in terms of rate of metal removal from
solution, or current efficiency. Rate of metal removal can be determined
either by doing a metal mass balance on inlet and outlet streams or by
weighing the amount of metal which has deposited on the cathode. Current
efficiency compares the actual amount of metal (or other contaminant) removed
to the amount that could be theoretically removed for a given current. In
practice, a high current efficiency is not necessarily equivalent to a high
rate of removal since removal rate increases with current.
The paragraphs below discuss, in detail, the case studies of electrolytic
treatment that were summarized in Table 9.3.1.
9.3,1 Concentric Cylinder Reactor
The reactor employed in this study consists of a central post-type anode
surrounded by a cylindrical cathode with a diameter of 8 inches and a height
of 6 inches. The waste solution is rapidly recirculated through the annular
space between the electrodes to provide a constant supply of metal ions. The
cathode material is stainless steel and the anode material depends on the type
of solution being treated. For acid copper solutions, the aateriai is
titanium; for corrosive fluoborate solutions used in solder (tin/lead) plating
niobium is preferred.
At this facility, the reactors were being used to treat rinsewater from
copper and solder electroplating. Four reactors were connected in parallel to
the copper plating dragout tank, and rinse solution was circulated through
each of the reactors at a rate of about 16 gal/min. This setup is shown
schenatically in Figure 9.3.1. The four reactors removed approximately
20 grams of copper per hour at an average current of 5 azps. The
9-7
-------
TABLE 9.3.1. SUMMARY OF PERFORMANCE DATA
Reactor type Flow
Concentric 16 gpm
cylinder
Waste type
Current Voltage (contaminant)
5 amps 7.5 volts Acid Copper
300-400 mg/L Cu
Removal
efficiency
or rate
5.7 g/hr
Current
efficiency
80 - 90 %
Carbon Fiber 46 gpm
46 gpm
Parallel Porous 2.0 gpm
Plate
Fluidized Bed —
Packed Bed
10 gpm
300 amps
Cadmium Cyanide-
300 mg/L Cd
800 mg/L cyanide
89-98%
0.8-2.4 g/min
30-94%
1.47-5.34 g/min
7.5 volts Electroless Copper- 80-85%
60-120 mg/L Cu
Cadmium Cyanide-
mg/L Cyanide
40-50 kg 20-30%
Cadmium per year
175 amps 7.7 volts Chromate - 17 mg/L 99.7%
of heaxavalent
chromium
conversion to
trivalent chromium
10 gpm
560 amps 40 volts Cyanide at 80 mg/L 65%
-------
DIRECTION OF BOARD
TRAVEL
ELECTROPLATING
BATH
CATHODE
750 GALLON
DRAGOUT BATH
750 GALLON
"SECONDARY
RINSE"
ELECTROLYTIC REACTOR
SAMPLING LOCATIONS
RINSEWATER DISCHARGE
TO WASTEWATER SUMP
2 GPM RINSEWATER
Figure 9.3.1. Schematic of electrolytic recovery system using concentric cylinder reactors.
Source: Reference 12.
-------
concentration of copper in the rinse discharge was reduced from 3,000 to
4,000 mg/L to between 50 to 100 mg/L to yield a current efficiency of
90 percent. Copper was recovered from the stainless steel cathodes once each
week. Removal was easily accomplished by hand, producing a metal foil that
could be sold as scrap metal.
Three reactors were connected to the dragout tank that followed the
solder electroplating bath. However, initial tests were not successful in
recovering tin or lead. Personnel at the facility indicated that these mee«ls
are sometimes recovered, but occasionally problems result due to the
corrosiviey of the solder plating solution. This corrosivity nay cause
etching of tin and lead from the cathode, thereby negating any electrolytic
removal. To overcome this effect, it would be necessary to increase current
to the reactor. The corrosivity may also cause occasional breakdown of the
pump that is used to circulate the solution through the reactor. For these
reasons, this type of solution is not always amenable to electrolytic
12 13
recovery. *
o
9.3.2 Carbon Fiber Cell
The data presented in Table 9.3.1 for the carbon fiber cell were
generated using a unit developed by HSA Reactors Ltd. of Rexdale, Ontario and
currently marketed by Metal Removal Systems, Inc. of Melville, NY. The
electrolytic cell module contains a carbon fiber cathode and an anode of
titanium coated with a rare earth oxide. The cathode1s'carbon fibers have a
diameter of only 5 to 15 microns and 1 gram of fiber has a surface area of
ti
6
f *J
2.6 x 10 cm . This is reportedly 1,000 times greater than the surface
area of other porous electrode materials.
Electrolytic recovery was used to recover cadmium and destroy cyanide in
a cadmium-cyanide barrel-plating line. A reactor was operated in a
closed-loop circuit connected to a dragout/recovery tank. Plating solution
drag-out was carried into the recovery tank in volumes of 1.5 to 8 liters for
each barrel plated. The contaminated rinse solution was then pumped through a
700 liter electrolytic reactor at a rate of 175 L/tnin.
9-10
-------
As shown in Figure 9.3,2, the rinsing of a barrel raises the cadmium
concentration in the rinse from 80 to 300 mg/L, Within approximately two
minutes, the electrolytic reactor restores this concentration to its original
level. The rate of removal then levels off, so that by the time another
barrel is ready for rinsing, the concentration is approximately 60 mg/L. This
curve clearly shows the dependency of rate of removal on metal concentration.
In this case, removal drops off sharply when the cadmium concentration falls
below 100 nig/L. Overall, between 89 and 98 percent of the cadmium washed out
in the dragout tank was removed by the electrolytic reactor, corresponding to
a recovery of up to 24 g/min of cadmium. As a result, the average cadmium
concentration in the facility's effluent decreased from 4.0 to 1.0 jng/L.
An electrolytic reactor was also used to destroy cyanide in the drag-out
tank. To enhance cyanide destruction, 80 R/L sodium chloride was added to the
system. The sodium chloride acted both as a source of chlorine for cyanide
destruction and as an electrolyte. Initially, cyanide destruction was
achieved in the sane electrolytic reactor as was used for metal removal. To
increase destruction rate, a holding tank and a second, smaller electrolytic
reactor were added to the system. This unit provided extended retention time
for increased cyanide destruction, resulting in a 93 percent destruction
efficiency, even when cyanide loadings were high (5.74 g/min). This
destruction rate maintained the concentration of cyanide in the recovery tank
at just over 1,000 mg/L. Before the modifications, the cyanide concentration
averaged 4,000 to 8,000 mg/L.
Operation of the electrolytic reactor requires "a. 19 hour cycle of which
161 hours is for cadmium removal and cyanide destruction and the other 3 hours
are for removal of the cadmiuta from the carbon fiber cathode. The latter- is
accomplished by pumping a high-strength cyanide solution through the reactor,
allowing cadmium oxide to form and dissolve in solution. This product is then
reused as cadmium-cyanide plating solution make-up.
9.3.3 Parallel Porous Plate Electrodes
The RETEC heavy metal recovery system used in this application consists
of a parallel plate electrode configuration, in which the cathodes are a metal
sponge-like material (as shown in Figure 9.3.3) formed by depositing copper on
a polyester foam. The result is a porous, flow-through cathode with a much
9-11
-------
300
250
200
. 150
u
100
50
Sorrel Entry
468
Time, minutes
10
Figure 9.3.2
Reduction Curve for Cadmium.
Source: Reference No. 6.
9-L2
-------
ELECTRODE c'"^:---^.
ISOLATOR
ANODE
CATHODE
BUS
RETICULATE
CATHODE
OUTLET
CATHODE CONNECTOR
CELL COVER
FLOW FILTER
r~ ANODE BUS
/W/ CONNECTOR CLIPS
CELL BOX
INLET
AIR SPARGER
Figure 9.3.3. Retec cell.
Source: Reference No. 14,
9-13
-------
higher surface area than a standard flat plate, stainless steel cathode. The
applicability of this system to different types of metal containing solutions
is summarized in Table 9,3.2. The majority of commercial applications have
been in acid or electroless copper plating rinses.
A recent study evaluated the performance of this unit on an electroless
copper rinse solution from a printed circuit board manufacturing facility.
The RETEC unit contained 25 copper-plated polyester "sponge" cathodes
alternated by 26 titanium-coated anodes, each measuring 45 by 38 cm, spaced
0,64 cm apart. Influent copper concentration ranged from 8 to 1,100 mft/L, and
flow rate ranged from 0.5 to 2,5 gpm. Testing of the unit yielded the
following conclusions:
» Copper removal efficiency averaged 80 to 85 percent.
* pH had little effect on copper removal over the pH range of 3 to 11.
* Removal efficiency was independent of influent copper concentration
above 50 fflg/L but removal efficiency decreased significantly below
this concentration.
* Removal efficiency was best at low flow rates - removal efficiency
increased by 15 percent when the flow rate was decreased from 25 to
0.5 gpm.
* Recirculation did not affect removal efficiency.
At a second facility, an identically designed but larger unit (50
cathodes as opposed to 25) was used to treat 10 gpm of combined copper plating
and etching rinses. Inlet copper concentration was 100 to 200 ppra and the
discharge was from the RETEC unit was below 1 pptn copper.
9.3.4 Fluidized Bed
The Chemlec "Cell, a fluidi£«d bed developed by the Electricity Council
Research Centre in England (see Figure 9.3.4), consists of a set of apertured
expanded metal-mesh electrodes immersed in a bed of small glass beads. The
bed is fluidized to about twice its packed depth by pumping rinseuater upwards
through a distributor and the bed. The glass beads impinge'on the electrodes
and provide a simple means of agitation and mixing. The electrodes are
•9-14
-------
TABLE 9,3.2. RETEG APPLICATIONS
Metal
Electrolyte
Comment s/cond it ions
Cadmium
Cadmium
Alkaline
Acid
Cadmium
Cyanide
Chromium
Any
Copper
Copper
Dilute Acid
Elect'roless
Copper
Strong Acids
Bright Dip
Above pH 10.5 cadmium is soluble.
Theoretically it can be removed at
efficiencies and to levels equal to
cadmium cyanide.
Cadmium can be reduced with a RETEC
unit but not to Federal compliance.
Cadmium can be reduced to
approximately 50 ppnt in dilute acid
media (pH 2-5) (lab data). Strong
acid would require bulk cathodes to
achieve the high cathode potentials
needed to protect the substrate.
Cadmium can be removed easily from a
cyanide solution to compliance
levels. 300 ppm at 3 gpra and 500 amp
should operate at approximately 30%
efficiency (field and lab results).
Neither trivalent (Cr+^)nor
Hexavalent (Cr*^) is removable at
2 grams/liter (gpl.) In theory
chrome could be removed from a more
concentrated bath .but chi's has noc
been achieved with RETEC (lab and
field).
Copper can be removed easily from
acid copper plating, bases (lab and
field).
Copper can be removed from
electroless copper rinses less easily
than acid baths but still at
acceptable levels. High levels of
chelate in the rinse will reduce
anode life (lab and field).
'Soluble copper can be removed from
almost any acid if the current
density is raised sufficiently to
counter the corrositi^ity of the acid
(cathodically protect). This
generally means elevated 1000 amp
....._._.*..<- £ -.*- __-'J -'.»... AC fit . "D. , 1 1r
«ur*c.i.i~ ~CT dwJ-— i. ^ y . w3 LS . ~— i r~
cathodes and/or elevated current for
(continued)
9-15
-------
TABLE 9.3.2 (continued)
Metal
Electrolyte
Comments/conditions
Copper (continued)
Copper
Ammoniacal Etches
Copper
Cyanide
acidity <5N. For acidity >5N a
RETEC-50 cell would have to be
modified to carry sufficient
current. Higher current densities
could be achieved in a RETEG Jr. Due
to the lower surface area, efficient
tnetal removal cannot be achieved with
bulk cathodes with metal
concentrations < 500 ppm. Long
recycle times would be necessary, if
the acid could not be reused with
this metal loading.
If the acid is not going to be reused
neutralization is usually a more
efficient alternative.
Copper can be removed from ammoniaca1
etches at a lower efficiency than in
acid media. Acidification will help
efficiency but it will not be as
efficient as straight acid.
In order co remove copper from
ammonicacal baths, it is necessary to
raise the current density (and the
cathode potential) to counter the
action of the etch out.
Complete copper removal from very
concentrated aramoniacal baths cannot
be achieved (lab).
Copper can be removed easily from
copper cyanide rinses. Efficiencies
are less than in acid baths (field
and lab).
Lead
Acid Baths
Lead can be removed efficiently from
acid baths, lead fluoborate will
disassociate to fluoride and attack
DSA anodes. Lead sulfamate baths are
preferred. Lead and lead complexes
are not very soluble. Lead above
pH 4 will form insoluble .lead oxide
(continued)
9-16
-------
TABLE 9.3.2 (continued)
Metal
Electrolyte
Comments/condLtions
Lead (continued)
Mercury
Nickel
Palladium
Ruthenium
Selenium
Silver
Silver
Tin
Tin
Aqueous
Watts Rinse
NiCl Rinse
JNi Acetate
Pd G1-NH4
Ale-Acid
Dilute Tab
Cyanide
Thiosulrate
Acid
Alkaline
and in the presence of sulfate (from
copper sulfate plating for instance)
will form insoluble lead sulfate.
(lab and field).
Mercury can be removed
electrolytically but will not adhere,
it will roll off the cathode
and deposit in the bottom of the cell
(beaker testing).
Removal at extremely low efficiency
at low current densities from cone.
solution.
Inefficient removal from dilute
solution with bulk cathodes at
approximately I gpl (lab).
Efficient removal from Pd bath
(lab).
Good removal from HC1 alcohol
solution (lab).
Removal to approximately 20 ppm.
Better results theoretically
achievable.
Removable electrochemically at high
efficiency (lab).
Removable electrochemically at high
efficiency. May react with copper on
cathode to cause premature cathode
deterioration (lab and field).
Removable to approximately 40 ppm in
RETEC eel 1C lab and field).
Should be removable to approximately
1 ppm.
(continued)
9-1?
-------
TABLE 9.3.2 (continued)
Metal
Electrolyte
Comments/conditions
Zinc
Zinc
Zinc
Acid
Cyanide
Alkaline
Not removable to compliance level.
Zinc should be reduced to 100-200 ppn
(lab).
Zinc removal electrochemically to
compliance levels (lab and field).
Low levels, should be achievable.
Source: Reference No. 14.
9-18
-------
LtCUID- J
MESH
DISTRIBUTOR-
rfc?J
BUSBAR
PLATING SCIUT1ON
IHJl CTIDN- (CnAG- CUT)
pi | J.NB CH
RESERVOIR
PIN'S £ 7AM10
Figure 9,3.4. The Cheraelee cell.
Reference No. 6,
9-19
-------
commonly titanium, but when treating alkaline cyanide solutions, both anode
and cathode may be made of mild steel. Also, since the metal is deposited on
the cathode mesh rather than on the bed medium (e.g., as for the carbon fiber
reactor), it is easy to remove the cathode from the cell and place it in the
plating solution where it can be used as an auxiliary anode.
At one facility, the Chemeiec ceil was used to recover cadmium from an
alkaline cyanide plating drag-out tank. The concentration of cadmium in the
plating bath waa 20 g/L and the Chemeiec cell maintained the dragout tank
concentration between 100 and 400 mg/L. The current efficiency ranged between
20 and 30 percent at an applied current of 30 amps. In the first 2 years of
use, the cell recovered 33.7 and 59.7 kg cadmium, respectively, despite
lengthy stretches of downtime due to low workloads. The maximum capacity of
2
the 1 m cell is estimated to be 200 to 300 kg/year. Power consumption for
electrolysis for these 2 years was 13 and 11 kw-hr/kg cadmiun recovered,
respectively.
Under these operating conditions, it was necessary to remove the anodes
from the Chemeiec cell every 2 weeks to collect the cadmium deposit. The
cathodes are lifted out of the cell and hung by copper hooks from the anode
rail in the plating bath. After about one 8-hour shift, the cadmium from the
coated cathodes has been dissolved. In this way, there is complete recycling
, . . 16 '
of cadmium.
9.3.5 PackedBed
The packing in a packed-bed electrolytic reactor can be any of a number
of different types of material. In the case discussed here, the packing
material is inert carbon particles. These are loosely packed in a cell with
the major electrodes in a parallel configuration, spaced 6 inches apart. The
major electrodes act as baffles to create an overall bed length of
100 inches. The loose packing of the carbon particles imparts a
semi-conductive nature to the bed. Potentials are maintained between
particles and/or agglomerates of particles creating many anodic and cathodic
sites within the major cell.
9-20
-------
In this application^ the electrolytic reactors are used to reduce
i
hexavalent chromium1 to the trivalent state and to oxidize cyanide. Chromium
reduction is accomplished Sy four electrolytic cells used in parallel, each
having a nominal capacity of 10 gpm. The system was designed to treat
120 mg/L of hexavalent chromium, although, the actual concentration fed to the
reactors averaged only 17 mg/L. As a result, the length of travel through the
cells was greater than required and the cell discharge averaged 0.06 nsg/L
chromium, well below the design target of 0.2 tog/L. This was equivalent to a •
99.7 percent reduction efficiency. At higher feed concentrations (100 to
150 mg/L), the cell discharge was closer to the design value of 0.2 mg/L.
Other conclusions of the testing are as follows:
* Control of waste pH entering the cell is essential; at hexavalent
chromium concentrations of 50 mg/L, a pH of 1.8 to 2.1 is adequate; •
at 150 mg/L, a pH of 1.5 to 1,6 is preferred.
* Power consumption varies with chromium concentration; for a 20 gpm
flow at 17 mg/L, 1.4 kw (175 sop, 7.7 volts) of power is consumed;
at the same flow rate, but 156 mg/L, 2.0 kw of power was consumed.
• The deposition of chromium onto the carbon particles will result in
a very high initial removal rate, however, after this initial
period, a steady state will exist for electrolytic removal of
hexavalent chromium.
It should be noted that the low and narrow pH range (1.5 - 1.6) observed
during testing is extremely difficult to achieve on a consistent basis. Pilot
plant testing with an emphasis on process control procedures is recommended
prior to any full scale implementation.
The electrolytic treatment system for cyanide consisted of three cells
used in parallel, each with a nominal capacity of 10 gpm. This system was
designed to treat cyanide-bearing rinses at a hydraulic loading of 30 gpm and
cyanide concentrations of up to 30 mg/L. However, actual cyanide concentration
was found to exceed 30 mg/L 98 percent of the time, with an average
concentration of 80 mg/L. As a result, it was only possible to achieve an
overall destruction efficiency of 65 percent; (to 28 mg/1), which was not
sufficient to meet the design goal of 2 mg/L of cyanide in the cell discharge.
Further evaluations indicated that there was straight-line relationship
between distance of travel along the packed bed and removal of cyanide;
namely, 0.5 mg/1 of cyanide is removed for each inch of travel. Therefore, to
achieve the design goal of 2 mg/L of cyanide, it would be necessary to
.increase the bed length from 100 to 156 inches.
9-21
-------
Other results of the study include:
* Cyanide was completely oxidized to CC>2 and t^.
• Electrical power use averaged 22.4 kw (560 amps at 40 volts) which
was calculated to be equivalent to 45 kw/kg of cyanide removal.
* It was not demonstrated -that complexed metals were destroyed by the
electrolytic process.
9.4 SYSTEM COSTS
The major cost associated with electrolytic treatment is usually the cost
of the reactor itself. This can range from S3,500 for a reactor with a
1-ft stainless steel cathode, up to $89,000,for A reactor with a high
n t ~t
surface area, carbon fibre cathode.° »'•' Additional capital costs will
include items such as a rectifier and electrical connections, pumps and
plumbing, and installation labor. These items may represent 15 to 25 percent
811
of the cost of the reactor. * Operating costs for electrolytic recovery
include electricity, maintenance or replacement of electrodes, labor, and
chemicals for oxidation of cyanide; e.g., NaCl. Chemicals may also be used in
some cases to strip metals from cathodes.
The best way to illustrate the cost of electrolytic treatment is to
compare its costs with those of other treatment methods; e.g., precipitation.
Electrolytic treatment can be very cost effective since it usually permits
recovery of metal from the waste solution and also precludes the generation of
metal-bearing sludges that require subsequent management as hazardous waste.
Cost comparisons of electrolytic recovery versus other treatment methods
are presented in Tables 9.4.1 and 9.4.2. Table 9.4.1 shows the costs of using
a carbon fibre electrolytic reactor versus the costs for alkaline chlorination
and precipitation for a waste stream containing cadmium and cyanide. The
annual operating costs for electrolytic recovery are $25,000 less than for the
chemical treatment alternative. This would permit a- payback time for tne
D
higher cost reactor of less than a year.
The chemical costs for electrolytic treatment shown in Table 9.4.1 are
for sodium chloride (NaCl), which is used to aid cyanide destruction, and
sodium cyanide (NaCK) and oxygen, which are used to remove cadmium from the
-------
TABLE 9.4.1. COSTS FOR CASBQN FIBER ELECTROLYTIC TREATMENT
VERSUS CHEMICAL TREATMENT
Electrolytic treatment Alkaline chlorination/
(carbon fiber) precipitation
Annu'al operating costs (£)
Electricity
Chemicals
Sludge disposal
Labor
Maintenance
Total:
Capital costs (£5
Reactor
Electrical and plumbing
Installation
Total:
900
11,150
0
1,150
5,950
19,150
87,500
21,900
1,200
110,600
NA
33,800
5,700
4,600
0
44,100
'NA
NA
NA
90,000
NA = Not available
Source: Reference No. 8.
9-23
-------
TABLE 9.4.2. COSTS FOR ELECTROLYTIC TREATMENT USING ARETEC
CELL VERSUS CHEMICAL TREATMENT
Electrolytic treatment
(Retec cell)
Sulfide precipitation
Annual operating costs (Jj)
Electricity
Labor
Maintenance
Recovered copper
Total Operating Cost (1)
Total Capital Costs:
875
1,250
7,750
(625)
9,250
44,000
NA
NA
NA
0
52,000
140,500
NA = Not available.
Source: Reference No. 15.
it
9-24
-------
carbon fiber cathode. The maintenance costs are for replacement of the
anodes, which is required every 2 vears, and for filter cartridges which are
Q
used to remove particulate matter from the waste influent.
Table 9.A.2 compares the costs of using electrolytic treatment versus
suifide precipitation to treat a copper-bearing waste stream generated from
the manufacture of printed circuit boards. In this case, two RETEC
electrolytic reactors are reauired to treat the 10 gpra waste stream which
contains 200 mg/L copper. The major operating cost of electrolytic recovery
is the maintenance cost to replace the cathodes. However, the overall
operating costs are only one-sixth of that incurred using suifide
precipitation. The latter costs were not itemized in the reference but
probably result from the purchase of sodium suifide and the cost for disposing
the metal-bearing sludge that results from precipitation of cadmium.
The capital costs for electrolytic removal are one-third those for
suifide precipitation. The capital costs for suifide precipitation include
the purchase of tanks, pumps, a clarifier, and a filter, plus a building
addition to hold this equipment. The equipment required for electrolytic
treatment is much more compact and does not require the construction of
additional building space in which to place tanks and clarifiers. If these
assumptions are accurate, the capital costs for electrolytic treatment are
significantly lower than for suifide precipitation.
Figure 9.4.1 presents an operating cost comparison for treatment of
electroplating wastewater containing hexavalent chromium with two different
treatment technologies. One of the technologies is a packed carbon bed
electrolytic reactor and the other is chemical reduction with sodiutu
bisulfite. Both of these technologies are followed by addition of NaOH to
precipitate trivalent chromium. Figure 9.4.L shows that chemical reduction is
*6
less expensive than electrolytic reduction at low Cr concentrations, but
becomes more expensive as the concentration increases. Electrolytic reduction
becomes more efficient at high metal concentrations; at a concentration of
17 tng/L Cr , approximately 25 kwh of power are required to reduce 1 kg
wnereas, at a concentration of 156 kR/*l, only 2 kwh are required. Conversely,
the amount of sodium sulfite required for reducing Cr increases almost
linearly -as the concentration of chromium in the wastewater increases from 17
to 156 mg/ 1.
9-25
-------
I
hJ
a.
15,000
I2.E.OO
•» 10,000
in
o
o
a:
UJ
o_
o
7,t>00
5,1)00
2,noo
a
LEGEND
O= CHEMICAL REDUCTION + PRECIPITATION
D= ELECTROLYTIC REDUCTION and PRECIPITATION
1
1
20
4O 60 80 100 120
HEXAVALENT CHROMIUM CONCENTRATION, mg/L
140
160
Figure 9.4.1. Operating Cost Comparison for Treatment of llexavalent Chromium Rinse.
Source: Reference No. 11.
-------
The case studies discussed above illustrate the fact that electrolytic
recovery can be more economical than other treatment methods in certain
situations. However, as indicated in the final case study above, the unit
cost of reducing hexavalent chromium electrolytically is greatly affected by
the inlet concentration. If the concentration is low, it may be more
economical to use chemical reduction.
Another major factor to consider is the cost of disposing the
raetal-bearing sludge generated by most chemical treatment methods. The land
disposal ban is likely to cause further increases in secure landfilling
costs. Thus, situations for which recovery is not yet economically feasible
may be more cost-effectively managed through recovery in the future,
9.5 PROCESS STATUS
Electrolytic recovery is applicable for certain metal/cyanide waste
streams. It is a particularly attractive process for metal-bearing waste
streams because it allows for metal recovery, thereby precluding the
generation of petal-bearing sludge.
A number of different types of electrolytic reactors are currently
manufactured. Simple, parallel-plate reactors can be used to recover noble
metals such as gold and silver. More complex units with porous or granular
electrodes may be required to remove roetals such as copper, tin, and lead,
particularly when these metals are present in low concentrations; e.g., less
than 100 mg/L.
9-27
-------
REFERENCES
1. Bishop, P.L., and R.A. Breton. Treatment of Electroless Copper Plating
Wastes. In; Proceedings of Che 38th Annual Purdue Industrial Wasce
Conference.
2. USEPA. Environmental Pollution Control Alternatives; Reducing Water
Pollution Control Costs in the Electroplating Industry.
EPA-625/5-85-016. 1985.
3. Centec Corporation. Navy Electroplating Pollution Control Technology.
Prepared for Naval Facilities Engineering Laboratory. February 1984.
4. Snoeyink, V.L. and D. Jenkins. Water Chemistry. John Wiley & Sons, Inc.
1980.
5. Swank, C.A. Electrolytic Recovery from Rinse Waters. Printed Circuit
Fabrication. 5(5). 1982,
6. Walsh, F.C., and D.R. Gabe. Electrochemical Cell Designs for Metal
Removal and Recovery. In: Proceedings of the Symposium on
Electroplating Engineering and Waste Recycle New Developments and
Trends. Cleveland, OH. August-September 1982.
7. Patterson, J.W, Wastewater Treatment Technology. Ann Arbor Science
Publishers, Ann Arbor, MI. 1975.
8. Vachon, D.T. et al. Evaluation of Electrochemical Recovery of Cadmium at
a Metal Finishing Plant. Plating and Surface Finishing.- pp. 68-73.
April 1986.
9. Allen, R. Electrolytic Recovery from Dilute Solutions. Finishing
Industries. September, 1982. pp 27.
10. Kitn, B.M., and J.L. Weininger. Electrolytic Removal of Heavy Metals from
Wastewaters. Environmental Progress, 1(2), pp. 121-125- May 1982.
11. Warner, B.E. Electrolytic Treatment of Job Shop Metal Finishing-
Kastewater. Prepared for U.S. EPA Industrial Environmental Research
Laboratory. EPA-600/2-75-028. September 1975.
12. Alliance Technologies Corporation. Case Studies of Existing Treatment
Applied to Hazardous Waste Banned from Landfill. Phase II - Case Study
Report for Facility F. Prepared for U.S. EPA Hazardous Waste Engineering
Research Laboratory. Contract No. 68-03-3243. June 1986.
13. Rosenbaum, Wayne. ETIGAM, Inc. Warwick, RI. Telseon with Mark Ariemti,
Alliance Technologies Corporation. February, 1986.
14. Omnipore, Inc., Sugar Land, TX. Technical Literature for the Retec Heavy
Metal Recovery System. March 1987.
9-28
-------
15. Kula, D. The Cost Benefit of Metal Recovery in the Plating Shop Instead
of Waste Treatment Uhile Attaining Compliance. Presented at tbe
Interconnections Packaging Circuitry CIPC) Fall Meeting, Denver, CO.
September 1983.
16. Tyson, A.G. An Electrochemical Cell for Cadmium Recovery and Recycling.
Plating and Surface Finishing. pp. 44-47. December 1984.
17. Pierce, R, International Circuit Technology (ICT), Lynchtmrg, VA.
Telecoms with Mark Arienti, Alliance Technologies Corporation.
February 18-19, 1986.
9-29
-------
SECTION 10.0
CHEMICAL TREATMENT/REMOVAL PEOCESSES FOR hETALS
The treatment processes discussed in this section are based on physical/
chemical methods of separation and removal of metallic contaminants in the
waste feed stream. Processes discussed are:
10.1 Precipitation
10.2 Coa£ulation and Floccuiation
10.3 Chemical Reduction
10.4 Flotation
All of these processes are used to some extent for the treatment of
wastes, but differ in their applicability to various types of waste and their
need for pretreatment and post-treatment procedures. The physical/chemical
treatment processes land the other treatment processes discussed in the
following sections) are considered within the fraaework of four major areas:
(1) Process Description including pretreatment and post-treatment
requirements; (2) Demonstrated Performance in Field and Laboratory; (3) Cost
of Treatment; and (4) Overall Status of the Technology.
10.1 PRECIPITATION
All precipitation processes operate under the same fundamental chemical
principles and utilize similar types of eauipment and process configurations.
Additionally, pretreattnent requirements and residual post-treatment options
are comparable, regardless of the specific precipitation method under
-i^\f^s«;!rir*2*"i.oTJt "j'Ha^g^^'-a s isi13 r ^spscts of ^r^Gi-^Ltst^-QH systiQ"'(s vi.^^ b *•*
addressed prior to discussion of specific reagent/waste combinations.
Section 10.1.1 serves as introduction to the basic theory of precipitation
10-1
-------
chemistry and proceeds to identify considerations in pretreatment
requirements, process equipment, process configurations, post-treatment and
disposal of residuals. The remaining subsections (Sections 10.1,2 through
10,1.4) address specific precipitation reagents. These highlight the unique
aspects of eachs including compatible waste types, treatment costs, sludge
generation and special considerations in equipment design and reagent handling
practices. The reagents are;
• Hydroxides;
• Sulfides;
• Carbonates.
Each reagent subsection covers the following topics:
,« General process description including typical operating
characteristics;
• " Performance data which identifies operating parameters, processing
equipment, and system configurations;
• Capital and operating costs;
• Status of the technology.
10.1.I General Considerations
10.1.1.1 Precipitation Theory—
The principal mechanism of precipitation involves the alteration of the
ionic equilibrium of, a metallic compound to produce an insoluble precipitate.
Typically, an alkaline reagent is used Co lower the solubility ,of the metallic
constituent and thus, bring about precipitation. In certain cases, chemical
reduction (Section 10.3) may be needed to change the characteristics of the
metal ions (i.e., valence state) in order to achieve precipitation. In
general, precipitation reactions form a salt and an insoluble metal complex,
as illustrated in the following reaction between nickel sulfate and sodium
10-2
-------
NiS04 * 2 NaOH = N«2S04 + Ni(OH>2 (s)
ntckei sodium sodium nickel
sulfate hydroxide sulfate hydroxide
Chemical precipitation normally depends on several variables :
Maintenance of an alkaline pH throughout the precipitation reaction
and subsequent settling.
Addition of a sufficient excess of treatment ions to drive the
precipitation reaction to completion.
Addition of an adequate supply of_ sacrificial ione (such as iron or
aluminum) to ensure precipitation and removal of specific target
ions.
Effective removal of precipitated solids.
Control of pH is essential for precipitation of many metals, as
illustrated by the solubility curves for selected metal hydroxides and
sulfides shown in Figure 10.1.1. Hydroxide precipitation is effective in
removing arsenic, cadmium, chromium (+3), copper, iron, manganese, nickel,
lead, and zinc. Sulfide treatment is superior to hydroxide (and carbonate
treatment) for removal of several metals. As shown by theoretical
solubilities of hydroxides and sulfides of selected metals (Table 10.1.1),
sulfide precipitation is highly effective in removal o£ cadmium, cobalt,
copper, iron, mercury, manganese, nickel, silver, tin, and zinc. Estimated
achievable maximum 30-day average concentrations of several heavy metals under
different chemical precipitation and solids removal technologies are shown in
Table 10.1.2. The estimated achievable concentrations are based on the
performance data reported in literature.
Another factor that effects precipitation reagent performance is the
presence in solution of chelator/complexing agents. A list of common agents,
together with their structures, is given in Table 10.1.3. These cheiator/
completing agents prevent the complete precipitation of heavy metal hydroxides
by competing with the hydroxyl ion for possession of the heavy metal, e.g.,
Zn(NH_)/"M' + 20H = 2n(OH)7(s) +• 4SH
34 23 Caq)
10-3
-------
Pb(OH).
10
10
10
-12
01 2 3 « 5 67 E S 10 11 12 13 14
pH
Figure 10.1.1. Solubility of metal hydroxides and sulfides
as a function of pH.
Source: Reference 1.
10-4
-------
TABLE 10.1.1. THEORETICAL SOLUBILITIES OF HYDROXIDES AND SULPIDES
OF SELECTED METALS IN PURE WATER
Solubility of metal ion, mg/L
Metal
Cadmium (Cd-*-+)
Chromium (Cr+++)
Cobalt (Co++)
Copper (Cu++)
Iron (Fe++)
Lead (Pb+-O
Manganese (Mn++)
Mercury (Hg++)
Nickel (Ni++)
Silver (Ag+)
Tin lSn++)
Zinc (2n*+)
As
2.
8.
2.
2.
8.
2.
1.
3.
6.
13
1.
1.
hydroxide
3 x 10~5
4 x 10~
2 x 10~1
-2
2 x 10
9 x 10~l
1
2
-~4,
9 x 10
-3
9 x 10
.3
-4
1 x 10
1
As carbonate
1.0 x 10~4 6.
No
1.
5..
3.
-3
7.0 x 10 3.
__ n
-2
3.9 x 10 9.
• -1
1.9 x 10 6.
-1
2.1 x 10 A 7.
. "\
7.0 x 10~4 2.
As sulfide
7
x
10
-10
precipitate
0
8
4
8
I
0
9
4
8
3
x
x
X
X
X
X
X
X
X
X
10
10
10
10
10
10
10
10
10
10
—8
-18
-5
-0
j
-3
-20
— S
-1 2
4, £m
-8
-7
Source: Reference 1.
10-5
-------
TABLE 10.1.2. ESTIMATED ACHIEVABLE MAXIMUM 30-DAY AVERAGES FOR THE APPLIED TECHNOLOGIES
o
I
Lime ppt*
followed by
sedimentation
Ant imony , Sb
Arsenic, As
Beryllium, Be
Cadmium, Cd
Copper, Cu
Chromium, CrC+S)
Lee'id, Pb
Meicury, Hg(+2)
Nickel, Ni
Silver, Ag
SeJenium, Se
Thallium, Tl
Zinc , Zn
0.8
0.5
0.1
0.1
0.05
0.0
0.3
0.2
0.4
0.2
0.2
0.5
- 1.5
- 1.0
- .0.5
- 0.5
- 1.0
-0.5
- 1.6
- 1.5
- 0.8
- 1.0
- 1.0
- 1.5
Final concentrations (mg/L)
Ferrite Soda ash Soda ash
Lime ppt Sulfide ppt coprecipitation addition addition
followed by followed by followed by followed by followed by
filtration filtration filtration sedimentation filtration
0.4
0.5
0.01
0.05
0.4
0.05
0.05
0.1
0.2
0.1
0.1
0.4
- 0.8
- 1.0 0.05 - 0.1
- 0.1
- 0.1 0.01 - 0.1 0.05
- 0.7 0.05 - 0.5 0.05
- 0.5 0.01
- 0.6 0.05 - 0.4 0.20 0.4 - 0.8 0.1 - 0.6
0.01 - 0.05 0.01
- 0.5 0.05 - 0.5
- 0.4 0.05 - 0.2
-0.5
- 0.5
- 1.2 0.02 - 1.2 0.02 - 0.5
*ppt = precipitation
Source: Reference 1.
-------
TABLE 10,1.3. STRUCTURES OP CHELAT1NG AGENTS SEPARATED
NTA
(Nitrilo triscetic acid)
X.CH2C02H
H02CCH2-N ^
^~~
EDTA
' (Ethylene dinitrilo tetraacecic acid)
EBTA
(Ethylene bis(oxyethylenenitrilo)tetraacetic acid)
CDTA
1,2 diamino cyclohexane tetraaeetic acid
N(CH2C02H)2
N(CH2C02H)2
Source: Adapted from Reference 3.
10-7
-------
The equation indicates that solutions which contain dissolved ammonia tend to
drive the reaction to the left, thereby preventing removal of zinc as the
hydroxide. Calculations show that for a solution containing 100 ppia of
dissolved NH.,, at a pH of 8.0, nearly 3.0 ppm of 2n will remain
. . 2
unprecipitated. All cotoplexing agents solubilize certain heavy metals in a
2
fashion similar to that given in the above example.
10.1.1.2 Pretreattaent Requirements —
Pretreatment of metals containing wastes prior to precipitation typically
consists of gross solids removal (e.g., filtration), flow equalization,
neutralization, or treatment of individual waste streams prior to combination
with other process wastes. These treatments of segregated wastes result in
economic benefits from reduced reagent costs and smaller equipment sizing.
Other common pretreatment processes include cyanide destruction, chromium
reduction, and oil removal.
Cvanide destruction;—Cyanide wastes cannot be mixed with metal-containing
wastes due to the formation normally stable organo-metallic complexes or the
possible evolution of toxic hydrogen cyanide gas. Instead, cyanide is
typically oxidized to carbon dioxide and nitrogen gas through a chemical
oxidation process. In two-stage chlorination, pH is typically maintained
around 11.0 in the first reaction vessel and 8.0 to 8.5 in the second vessel
through addition of NaOH, as required (see Section 14.1),
Chromium reduction—Chromic ac-id wastes may contain hexavalent chromium
which must be reduced to the trivalent form prior to precipitation. Reduction
typically occurs at pH 2.0 to 3.0 through addition of acid (e.g., sulfuric)
and a reducing agent (e.g., sulfur dioxide, ferrous sulfate, sodium
metabisulfite, or sodium bisulfite). However, alkaline reduction (pH 7 to 10)
using ferrous iron has also been demonstrated. It has proven to be
cost-effective for highly buffered alkaline waste and the treatment of mixed
metal wastes containing less than 10 mR/L of hexavalent chromium (see
Section 10.3).
10-6
-------
Neutralization—Neutralization consists of adjusting an acidic or
alkaline waste stream with the appropriate reagent to a final pH of 6 to 9,
which meets surface water discharge requirements established under the Clean
Water Act. However, it is sometimes only necessary to adjust the pH to
approximately 5 to 6 (i.e., partial neutralization) to achieve certain
treatment objectives. In other applications it may be necessary to neutralize
an acid to pH 9 or higher to precipitate metallic ions or to completely
clarify a waste for acceptable discharge. These techniques are called under-
4
and over-neutralization, respectively.
Table 10.1.4 identifies several o£ the more prevalent neutralization
reagents and their characteristics. The selection of the appropriate reagent
for wastewater neutralization processes is site-specific and dependent on the
following considerations: wastewater characteristics, reagent costs and
availability, speed of reaction, buffering qualities, product solubility,
costs associated with reagent handling, and residual quantities and
characteristics. Typically, the first step in reagent selections is to
characterize the wastewater. General parameters of interest include flow
(rate, quantity), pH, pollutant loading, physical form of waste, and
waste/reagent compatibility. These characteristics narrow the range of
reagents and treatment configurations available for consideration.
Following the selection of candidate reagents, the quantity of reagent
required to neutralize the waste to the desired end point must be determined.
Reagent quantity is usually calculated by developing a titration curve for
each candidate reagent using representative wastewater samples. ' These
data determine the quantity of reagent required to bring the sample volume of
wastewater to the desired pH.
The next step in the experimental procedure is the preparation of
reaction rate curves and development of kinetic rate equations for each
candidate reagent. Reagent reactivity is an important factor in determining
retention time and consequently the size of the treatment facility, the final
effluent quality, and the ease or difficulty of process system control. These
parameters, in turn, will affect both capital and operational co5ts associated
with the wastewater treatment system. Reaction rate curves for various
quantities of residual reagent (,L.e.» excess above stoichioraetric
10-9
-------
TABLE 10.1.4. ACID/ALKALINE NEUTRALIZATION AGENT CHARACTERIZATION
o
o
Holecul.
Higli C. It nit. C.CO,
High C.lciun, C.COII),
HydrAleJ Lime
Hi 1(1, C.lcium CaO
l/iicbl line
Carbon CO2
Hydt.led Line HgO
Dolorailic C.O-HgO
Quicklime
Sod. Ath N.2CO]
Cflu.tic 3od. N.OH
H.gneii. HgO
Su If uric "2s0!.
Acid
Hurt.cfc IfCL
Acid
C«lclu> 100. 1
C.lelu* 74.1
HyJro.jJ.
C.Kiu. 56.1
Glide
C.rbon 54.0
Honn.l 114.4
Do 1 am i t i c
C.leiun 96.4
Sodiu. 10S-0
C.rbon.t.
Sodiua 40.0
HyJraiide
H.gneilun 40.11
Oxi Je
Sulfuric 98.1
AclJ
Hydro- )«.)
chloric
,c.J
Cannon fan*
Fowler
granule.
Powder
72-74J C.O
Pebble
91-98Z ClO
Pouder
46-48Z C.O
31-341 HgO
Pebble
55-58Z C.O
38-4 IJ Me"
Fovder
Liquid
13< H.OII
Pouder
Liquid
77Z lljS04
93Z H2S04
20'Be
Bulk Solubility
2000-2BOO O.OOI47?i
400-640 O.I530
to Ca(GII)2
416-666 0. 1530
(S 30-C)
801-1165 Converted
to Ca(OII)}
.rtJ H|t(ull>2
560-1041 17. 6.10
(S JO'C)
CS 100-C)
1017.9 0.008610
(0 30'C)
1704-1634 Complete
M57-IIJJ Couplets
Alk.ll He>itr>- H.y b. — Tank car 200 ]7
fcoui Mue g«.
iolublfl
handling
Acid Heutra- Illghty High Coat Bagged 0.929 ]65 39?. 89 70.16
1 iiflt ion i«act i va
inenpenilv« « i t It calcium
•Reference 3.
-------
-------
requirements) are determined by plotting pH as e function of time. Other
variables which should be monitored include temperature rise, agitator speed,
density, viscosity, color, sludge volume, and eettleability.
From the titration and reaction rate curves, kinetic rate equations can
be developed. Methods for determining the, kinetic rate equation
(e.g., integral or statistical analysis) are discussed in the literature.
The reader is also referred to standard engineering texts for reactor and
8
costing methodologies based on flow parameters and kinetic rate equations.
In the final selection, the optimal reagents and reagent/waste feed ratio
will be those which incur the least overall cost, including not only the cost
of the reagent itself, but also the cost of purchasing -and maintaining the
reagent and neutralization systems, and the costs associated with residual
handling. The combination of all such factors may make a slightly more
expensive reagent less expensive overall.
Oil removal—Removal of oil through enmleion breaking, dissolved air
flotation, skimming or coalescing may also be needed prior to precipitation.
Traditionally, emulsified oils have been treated at low pH (e.g., pH of 2,0)
with alum. Hovever, this form of treatment is giving way to the use of more
effective emulsion breaking coagulants such as cationic polymers and other
3
specialty chemicals.
Flow equalization—The most prevalent form of pretreatment is flow
equalization. It is generally used in facilities which experience wide
variations in the wastewater flow or pollutant concentrations.
Figure 10.1.2 illustrates a number of ways that flow equalization can be
achieved. In all methods of flow equilization, care must be exercised during
the wastewater analysis to completely characterize any peak flows or
concentrations that night overload the system. In addition, flexibility in
system design should be provided for any future expansion, change in location,
or deviation in flow rates.
10.1.1.3 General Precipitation Equipment—
Ine neart of a wastewater precipitation process is the pH control system,
which must bring the waetewater pH to the level required to precipitate the
optimal Quantity of contaminant metal salt. As previously discussed, pH
10-11
-------
-------
(a)
Legend:
FT = flow transmitter
FC = flow controller
UT = level transmitter
tor |___j>.
Figure 10.1.2
Alternative concepts for wastewater equalization:
(a) batch reactor system, (b) batch equalization
continuous processing, (c) side-scream equalization,
and (d) flowtnrough equalization,
Source: Reference 7.
10-12
-------
control is achieved through the use of a neutralization system.
Neutralization may occur either in the precipitation reactor or in a separate
tank. A wide variety of treatment options and configurations are available.
However, fully engineered component neutralization/precipitation systems
generally consist of the following equipment:
• Neutralization/Precipitation System
TankCs)
Mixer(s)
— pH control instrumentation
• Chemical Feed System
TankCs)
Mixer(s)
-• Level instrumentation
- Metering equipment
* Miscellaneous
- Flow monitoring
- Effluent pH recorder
— Electrical and mechanical fit-up
— Incremental engineering requirements
In addition, there is a need for facilities and equipment to collect and
segregate the wastewaters, transport the wastewaters to equalization sumps,
pump the wastewaters to the treatment system, perform liquid/solid separation,
and convey the treated wastewaters to the point of discharge.
Precipitation tanks are fabricated from a wide range of construction
materials such as masonry, metal, plastic, or elastomers. Corrosion
resistance can be enhanced with coatings or liners which prevent the premature
decomposition of tank walls. For example, concrete reactors susceptible to
corrosion can be installed with a tup-layer coating" of a 6.3 mm base surface
(glass-reinforced epoxy polyaaide) covered by a 1.0 mm coating of polyurethane
elastomer to extend service lifetimes.
Vessel geometries can be either cubical or cylindrical in nature with
agitation provided overhead in line with the vertical axis. While cubical
tsnks need no bsiili-g, cylindricsl vessels sre typically car;3:ruered with
suitable ribs to prevent swirling and maintain adequate contact between the
10-13
-------
reactance. A general rule of thumb in the design of precipitation reactors is
that the depth of the liquid should be roughly equivalent to the tank diameter
-j u 9 ' .
or width.
Reactors can be arranged in either single- or multi-stage configurations
and operate in either batch or continuous mode. Multi-stage continuous
configurations are typically required to neutralize and precipitate
concentrated wastes with variable feed rates. In these units roost of the
neutralization reagent is, added in the first vessel with only final pH
adjustments (polishing) and precipitation agent addition made in the remaining
reaction vessels. This is particularly true when using reagents which require
extens-ive retention time. Single-stage continuous or batch precipitation is
suitable for most applications with highly buffered solutions or dilute
wastewaters not subject to rapid changes in flow rate or pH.
An adequate retention time is required to provide time for the
precipitation reaction to go to completion. This factor is especially
critical where a dry feed (lime or ferrous sulfide) or slurry is used as the
control agent. In these systems, the solids must dissolve before they react,
increasing the required retention time and tank capacity. For example, liquid
reagents used in continuous flow operations generally require 3 to 5 minutes'
of retention time in the first tank. Three minutes corresponds to the minimum
time for adequate mixing. In comparison, solid-based reagent systems such as
lime or ferrous sulfide typically require approximately 30 minutes of
retention -time. ' '
The pH control systems for batch precipitation processes can be quite
simple with only on-off control provided via solenoid or air activated
valves. Control system designs for continuous flow precipitation systems are
more complicated because the wastewater feeds often fluctuate in both flow and
concentration. Systems currently available include; proportional, cascade,
feedforward, or feedback pH control. Each system has distinct advantages and
disadvantages which are discussed in detail in the literature. ' ' ' The
pH control equipment usually consists of a pH probe, monitor, and recorder.
In addition, there is typically a control panel with an indicator, starters
arid controls for metering pumps, all relays, high/low pH alarms, switches, and
Chemical feed apparatus consists of storage tanks, agitation} level
instrumentation, and metering pumps. Storage tanks should be sized according
to maximal feed rate, shipping time required, and quantity of shipment. The
10-14
-------
total storage capacity should be more than sufficient to guarantee a chemical
supply while awaiting delivery. Storage containers must be suitable for the
reagent being used. For example, hygroscopic reagents such as high calcium
quicklime must be stored in moisture-proof tanks to prevent atmospheric
degradation.
In addition to the chemical feed and neutralization/precipitation
systems, both flow monitoring and effluent pH recording equipment are
necessary to prevent discharge of insufficiently treated waste resulting from
surges or upsets. Also, spare parts such as pH probes, pH controller circuit
board, metering pump ball valves, o-rings, and strainers should be kept on
hand to prevent any excessive downtime.
10.1.1.4 Clarification and Sludge Consolidation—
Clarification and sludge consolidation unit operations are typically
applied as post-treatments to the majority of aqueous metals containing waste
treatment systems. Figure 10.1.3 illustrates a general treatment/
post-treatment approach for aqueous metal/cyanide bearing waste streams.
Usually, wastewaters undergo chemical treatment and enter a clarifier
where the flow is decreased to a point that allows solids with a specific
gravity greater than that of the liquid settle to the bottom. For
liquid/solid mixtures with a slight density difference, an organic polymer
(flocculant) can be added to allow the solids to agglomerate and improve the
14
settling characteristics (see Section 10.2). The supernatant in the
overflow is drawn off and residual trace organics or solids are removed in a
final polishing step such as carbon adsorption, ultrafiltration, or ion
exchange. The solids in the underflow can then be discharged to a holding
tank for subsequent dewatering.
In addition to differences in the quantity of sludge generated, each
reagent imparts to the sludge variable settling characteristics, thereby
affecting the sizing parameters of downstream equipment. For example,
lime neutralized sludge exhibits a granular nature that settles fairly rapidly
2
and dewaters effectively (4 to 20 Ib of dry solids/hr/ft yielding a 3/16 to
3/8 in. cake). Conversely, sodium hydroxide sludge results in a fluffy
gelatinous precipitate witn low settling rates." Figure 10.1.4 snows the
10-15
-------
DESTRUCTION
TREATMENT
AQUEOUS
METAL [X
WASTE ^
DESTRUCTION
i.e. ALKALINE
CIILORINATION
O
I
PRECIPITATION
LIME, SULFIDE,
SULFATF, CHROMIUM
REDUCTION, ETC.
FILTRATION
MIXED-MEDIA
CARBON, SAND,
ULTRAFILTRATION
STABILIZATION
SOLIDIFICATION
Figure 10.1.3. A general treatment approach for aqueous metal/cyanide bearing waste streams.
Source: Reference 13.
-------
-------
It)
411!
HE4TER
WASH
* ! inif cra
fer MifliR
* «r. l-ld
* 4BBitn) f
* Gil Fine S
lirot—Mi
il™mt. US.
(O
llf
FRfHUKK
VliSH
1C !5
1C 1C 3i
Figure LO.1.4, Settling rate curves.
Source: Reference 16,
10-17
-------
-------
results of three settling tests conducted on power plant effluents with both
lime and sodiun hydroxide. In all three cases sodium hydroxide settled more
slowly, and in subsequent filtration tests, dewatered about half as
effectively. However, the use of lime or calcium carbonate generates
greater sludge weight and volume. This is primarily due to insoluble acid
salts and calciun sulfates formed when precipitating metal sulfate containing
wastes such as acid plating baths. Therefore, as landfill and hauling, coats
become more significant, sodium hydroxide becomes more competitive with lime
and limestone as a precipitation agent.
Few, if any, sludges settle at a rate sufficient to utilize only
clarifiers or thickeners to accumulate sludge for disposal on land.
Therefore, the underflow from the clarifier is typically concentrated through
the use of mechanical dewatering equipment such as centrifuges, rotary vacuum
filters, belt filters, drying ovens, and recessed-plate filter presses. The
obtainable degree of cake dryness can be determined by bench-scale tests by
the equipment vendor to identify the suitability of a particular dewatering
device (see Table 10.1.5). The low solids content of sodium hydroxide after
1 R
Sedimentation (3 to 10 percent) requires the use of a filter press.
Conversely, suspended solids removal from lime neutralized sludges can be
accomplished through 'use of a wider range of equipment including rotary vacuum
or continuous belt filters.
10.1.1,5 Land Disposal of Residuals—
Installation of a metals precipitation system inevitably results in the
problem of sludge disposal. The cost of hauling the sludge to a licensed
hasardous waste Landfill will depend on the volume of sludge, the distance
hauled, and the sludge composition. Sometimes it is possible to dispose of
calcium-based reagent sludges through agricultural or acid pond liminfc. In
one neutralization/precipitation application, over 200,000 Ibs/acre of
lime-treated waste pickle liquor sludge was applied onto Miami silt loam to
, ,, 19
improve overall crop yields.
ID-IS
-------
o
I
Reproduced fror^ ~m^
best available copy
-^ 9? TABLE 10
Grnvity
I'n riimc t e r (lowpressureJ
"Jake sol ids X lf> - 2U
v,i rinb les feed
-Polymer
-Belt apeed
-Depth of ft ludpe
in cylinder
Ad vantages -Low energy &
* cnpi tnl coat
-Low spnoe
rcqtii rcment 8
— Require** little
operator skill
|iis stl vant,iE''fl -Limited capacity
-Low nol ids
concent rat i on
-Requi ren large
quantity of
condit inning
chemicals
.1.5. SUMMARY OF SLUDGE DEWATERING DEVICE CHARACTERISTICS
Basket
contri tige
20 - 30
-Time; at full
e peril
flkiinining
-Sludge feed
rate
— Same1 machine
for thicken ing
-Uoru flaviklo
very r ICKID le
attention
-Unit in not
coul inuoun
-Digit cotio of
capital coat
to capacity
-Rcqiiircfl com-
p lex controls
control
Solid bnwl
UgP
.30 - l\2
di [ Cerent ial
speed
-Sludge feed
rate
-Low space
ret|tii re men t
|l 'k' "
dcvatering
-lliph rate of
Cced
-Can operate on
highly vari-
able feeds
-Requi res
prescreen ing
-Very noisy
with higli
vihrat ion
-High power
conauinpt ion
maintenance
akillfl
acuum
30 - 40
of H20
-Drum apeed
-Conditioning
chemicals .
-Filter media
operation
— 1 Jing med m 1 L f e
y P
—High power
requirement
—Vacuum pumps
are noisy
-Requires at
least 3X feed
aolida for
opera
Belt filter
press
36 - 46
-Belt tension
-Uaahwater flow
-Belt type
-Polymer
cond itioner
press produces
vibrat ion
-Continuous
operation
-Very sensitive
to incoming feed
-Short med io. 1 iCe
-Greater
operational
attention and
polymer dosage
50 - 60
-Filtration time
-Use of p re co at
-Filter cloth used
-Hi gh sol ids capture
y
ency
Vdf'M° mC^ t" """
requiremen
-High capital cost
-Batch discharge
-II igli polymer usage
-Media replacement
cost s are high
Source: Adapted from Reference 18.
-------
Another option is to treat the waste by immobilizing the waste
constituents for as long as they remain hazardous. This method of treatment,
based on fixation or encapsulation processes, is a possibility for some metals
containing wastes. Certain of these residuals could be found hazardous; their
heavy metal content may lead to positive tests for EP toxicity. In such
cases, encapsulation may be needed to eliminate this characteristic.
The following discussions will summarize available information concerning
immobilization techniques, namely solidification/fixation or encapsulation.
Chemical fixation involves the chemical interaction of the waste with a
binder; encapsulation is a process in which the-waste is physically entrapped
within a stable, solid matrix.
Solidification technologies are usually categorized on the basis of the
principal binding media. These media include: cement-based compounds,
lime-based pozzolanic materials, thermoplasts, and organic polymers
(thermosets). '* The resulting stable matrix produces a material that contains
the waste in a nonleaehable form, is notidegradable, cost-effective, and does
not render the land it is disposed in unusable for other purposes. A brief
summary of the compatibility and cost data for selected waste solidification/
stabilization systems is presented in Table-s 10.1.6 and 10..1.7.
Cement-basedsystems—These systems utilize type I Portland cement,
water, proprietary additives, possibly fly ash, and waste sludges to form a
20
monolithic, rock-like mass. In an EPA publication, several vendors of
cement-based systems reported problems with organic wastes containing oils,
solvents, and greases not miscible with an aqueous phase. For although the
unreactive organic wastes become encased in the solids matrix, their presence
can retard setting, cause swelling, and reduce final strength. These
systems are most commonly used to treat inorganic wastes such as incinerator
generated wastes and heavy metal sludges from neutralization/precipitation
processes.
10-20
-------
TABLE 10.1.6. COMPATIBILITY OF SELECTED WASTE CATEGORIES WITH DIFFERENT WASTE
SOLIDIFICATION/STABILIZATION TECHNIQUES
1 rr.ilitL'nt Type
Organic Self- Class IliestIon and
U*)|C Cement 1.1 we Thtrnopl antic palyrerr Surface ccnenlIng aynihetlcaliivral
coapoitent bscvd baaed aol IJ11 lest Ion (III) • encapsuI at Ion techniquea formal Ion
Of gsnlcsi
I. Organic H*y impede Many Impede eel- Organic) *»y Kay retard id Hunt first be Fire danger Uiaie*. daconpoaa «l
aolvviiii end selling, «sy -
down, (1 re down of rncipBulai- ai« prcaent abla rcactlona
lug material*
). Sulfjfea Hay retard a* I- Camp at Ible Hay dehydrate Compi title Co«pat Ible Coapat Iblsi Coup at Ibla In stany
ling Anil and reliy.liaie Cases
cau^ecpalllng caualng
unless special apllttlng
cenent la uaed
It. llalldrs Fan My leached Kay retard set. Hay dehydrate CuiipJtlblc Cunpatlble Compatible If Conpatlb|« In nany
/run cemi-Mi. moat are aulfaies casea
nay retard en allyle allied arealso
• ell I I'd present
). Heavy nelals CoQpatlble Compatible Compatible Acid pll Bolu- Compatible Compatible If Co«psttbl« in stany
blllze* aeial sulfaCes cases
hydrAnldes ar« praaetnt
6. Kml IQJCIlvt CunpaiIble CompalIble Comjuu Ibis CunpdiIble CunpaiIble CunpaIIble If CovpalIblc
maicr lit • auHaiea
at* prevent
* lire J-li>rnj|dcliyile i esln.
Source: Rclercncc 22.
-------
TABLE 10.1.7. PRESENT AND PROJECTED ECONOMIC CONSIDERATIONS FOR WASTE SOLIDIFICATION/
STABILIZATION SYSTEMS
f
0
ro
to
Type at treatment
system
Cement-based
Poziolanlc
Thermoplastic
(bl tumen-based)
Organic polymer
(polyester system)
Surface encapsulation
(polyethylene)
Sel f-cenent Ing
Claaalf lea t Ion/mineral
synthea 1 a
Major
materials required
Portland Cemer.t
Lime Flyash
Bitumen
Drums
Polyester
Catalyst
Drums
Polyethylene
Gypsum (from uastc)
Feldspar
lln 1 1
cost of
ran i e r 1 .1 1
50.03/lh
SO.oj/ib
S0.05/lb
*27/cJrum
$0.45/lb
$1 .ll/lb
$!7/drura
Varies
"
50.01/lb
Ajn.iunt uf mj- Cost of m,i-
tvrlal rctjulieil lerl^l requlruil
to lre.il 101) Il.s in (rent 11)0 II";
of r.i-f w:i^te of i iiu w.i-jie Trends lit price
10(1 Ib $ 3. no Sl.ilile
100 Ib $ 3. 00 Slulilc
100 Ib $18.60 Koyed to oil
O.U drum |n icel
41 Ib of $27.70 Keyed to oil
polyester- prices
cjttily^t mix
Varies $ 4.50" K.-ye.l to oil
|»r icei
10 Ib •• Stable
Varies — Sinblc
-
Equipment Energy
Cod 1 ^ ubir
l.ou l.ou
Low l.ou
Very hlgli High
Very high High
Very high High
Moderate Moderate
High Very high
• Based on the full cost of $9l/tiiti.
•• Negligible but energy cost for calcining are appreciable.
Source: Reference 22.
-------
-------
Lime-based (pozzolanic) tect>niqye8~~Pozzolanic concrete is the reaction
product of fine-grained aluminous siliceous (pozzolanic) material, calcium
(lime), and water. The pozzolanic materials are wastes themselves and
typically consist of fly ash, ground blast furnace slag, and cement "kiln
duet. The cementicious product is a bulky and heavy solid waste used
primarily in inorganic waste treatment such as the solidification of heavy
metal and flue gas desulfurization sludger.
Thermoplastic materiel—In a thermoplastic stabilization process, the
waste is dried, heated (260 to 450°F), and dispersed through a heated plastic
matrix. Principal binding media include asphalt, bitumen, polypropylene,
polyethylene, or sulfur. The resultant matrix is relatively resistant to
leaching and biodegradation, and the rates of loss to aqueous contacting
fluids are significantly lower than those of cement or lime-based systems.
However, this process is not suited to wastes that act as solvents for the
thermoplastic material. Also, there ie a risk of fire or secondary air
22
pollution with wastes that thermally decompose at high temperature.
Organic polymers (thermosets)—Thermosets are polymeric materials that
cross-link to form an insoluble mass as a result of chemical reaction between
reagents, with catalysts sometimes used to initiate reaction. Waste
constituents could conceivably enter into the reaction, but most likely will
be merely physically entrapped, within the cross-linked matrix. The
cross-linked polymer or thermoset will not soften when' heated after undergoing
the initial set. Principal binding agents or reactants for stabilisation
include ureas, phenolics, epoxides, and polyesters. Although the
thermosetting polymer process has been used most frequently in the radioactive
waste management industry, there are formulations that may be applicable to
certain precipitation sludges. It is important to note that the concept of
thermoset stabilization, like thermoplastic stabilization, does not require
that chemical reaction take place during the solidification process. The
waste materials are physically trapped in an organic resin matrix that, like
thermoplastics, may biodegrade and release much of the waste as a
"*3
leachate.' It is also an organic material that will thermally decompose if
exposed to a fire.
10-23
-------
-------
Encapsulation is often ysed to describe any s-tdbilization process in
which the waste particles are enclosed in a coating or jacket of inert
material, A number of systems are currently available utilizing
polybutadiene, inorganic polyners (potassium silicates), Portland concrete,
polyethylene, and other resins as macroencapsulation agents for wastes that
have or have not been subjected to prior stabilization processes. Several
different encapsulation schemes have been described in Reference 24. The
resulting products are generally strong encapsulated solids, Quite resistant
25
to'chemical and mechanical stress, and to reaction with water. Wastes
successfully treated by these methods and their costs are summarized in
Tables 10.1.8 and 10.1.9, The technologies could be considered for
stabilizing precipitation sludges, but ere dependent on the compatibility of
the precipitation waste and, the encapsulating material. EPA is now in the
process of developing criteria which stabilized/solidified wastes must meet in
9ft
order to make them acceptable for land disposal.
10.1.2 Hvdrpxide Precipits.iicn
10.1.2.1". Process Description —
Hydroxide precipitation for heavy metals removal from aqueous waste
streams is both an effective and economical treatment technology. The
treatment converts soluble metal ions into insoluble hydroxide compounds. The
netals can then be separated from the liquid through sedimentation and/or
filtration. The most commonly used precipitating agents are line fCaO or
Ca(OH>2J and caustic soda (NaQH),
Hydroxide precipitation has been widely applied in treating industrial
wastewaters. The following is a list of some of the industries which use this
technology;
* nonferrous metal processing,
* ore mining and dressing,
• utility power generation,
« metals plating, and
* battery manufacturing.
10-24
-------
-------
TABLE 10.1,8. ENCAPSULATED WASTE EVALUATED AT THE U.S. ARMY WATERWAYS
EXPERIMENT STATION
Code Mo.
100
SCO
300
LOO
500
600
700
SCO
900
1000
fource of Waste
SO scrubber sludge, line process, eastern
coal
Electroplating sludge
Nickel - cadmium battery production sludge
S0x scrubber sludgy, limestone orocess '
eastern coal
SO scrubber sludge , double alkali process
eastern coal
SO scrubber sludge , IJjnestone crdcess ,
western coal
Plpnent production sludge
Chlorine production brine sludge
Calcium fluoride sludge
SO scrubber sludge, double alkali process,
western coal
fajor Contaminants
r*a efs /ct^.
^& , 5Ujj /^. -
Cu, Cr, 7.n
Ml, Cd
Cu, SOU"/S03=
Na, Ca, S0^*/S03"
P ' Qfv /^n '
UU , iSJ^ / w>W.j
Cr, Fe, CN
»Ta, Cl", Hg
Ca, ?'
Cu, :te, so//so3"
Source: Reference 25.
TABLE 10.1.9'. ESTIMATED COSTS OF ENCAPSULATION
Process CutIon
Esiltated Cost
Resin ruslon:
Unconflned waste
55-Callon drums
Resin spray-en
Plastic Welding
J 110/cry ton
Not determined
! SO, 000 55-gxl druas/year)
Source: Reference 25.
10-25
-------
-------
In the first step, the hydroxide precipitating agent is thoroughly nixed
with the wastewater stream. The reactions which begin in the flash-mix tank •
and which result in formation of the insoluble metal hydroxides are given
below where M is the metal cation removed.
for quicklime:
CaO + H20 = Cat OH) 2
M** + Ca(OH)2 = M(OH)2 + Ca**'
for hydrated lime:
M++ * Ca(OH)2 = M(OH>2 + Ca**
for caustic soda:
M*+ + 2NaOH = M(OH)2- + 2Ka*
Hydroxide precipitation is capable of removing certain metals found in
acid wastewaters. Among the metal ions removed are arsenic, cadaiuai, copper,
27
trivalent chromium, iron, manganese, nickel, lead, and zinc. Table 10.1.10
presents reported residual concentrations to which hydroxide precipitation can
remove these metals. This information is based upon application of hydroxide
precipitation to various industry wastewaters. It is important to note in
Table 10.1,10 that in some cases, e.g., lead, cadmium and zinc, the residual
concentrations reported are Lower than the theoretical solubilities of the
27
pure element in water. Several phenomena influence the effectiveness of
precipitation, e.g., ionic strength, coprecipitation , and adsorption. These
phenomena will ultimately determine the residual concentrations in specific
applications, especially in solutions containing several metal ions.
As stated previously, the most commonly used precipitating agents are
line, hydrated lime, and sodium hydroxide. The following is a brief
description of each reagent type.
Lime slurry — Lime • slurry treatment of metal laden waste streams is one of
the oldest and perhaps most prevalent of all industrial waste treatment
10 _ , . . . .. ..
processes, it is usea extensively as an alkaline reagent in the
10-26
-------
TABLE 10.1.10. HYDROXIDE PRECIPITATION METAL REHOVAL EFFECTIVENESS
Metal
Arsenic
Cadmium
Chromium, trivalant
Copper
Iron
Lead
Manganese
Nickel
Zinc
Inlet concentration Residual concentration
(tn«/L) (mg/L)
0.2 - 0.5 0.03
ND 0.0007
1,300 0.06
204 - 385 0.2 - 2.3 .
10 0,1
0.5 - 25 0.03 - O.I
ND 0.5
5 -0.15
16,1 0.02 - 0.23
Source: Reference 27.
10-27
-------
precipitation of pickling wash waters, plating rinses, acid mine drainage, and
9 28 29
process waters from chemical and explosive plants. ' ' It is used in
many applications as a low-cost alkali due to its pumpable 'form, and
effectiveness in removing Ca salts from the process. However, a major
disadvantage of the process is the formation of a voluminous sludge product.
Since limes are formed by the thermal degradation of limestone
(calcination), they are available in either high cslcitur, (CaO) or dolomitic
29
(CaO-Mg.0) form , These pure, oxidized products are referred to ss
Quicklime. Quicklime varies in physical form and size, but can generally be
obtained in lump (63 to 255 nun), pebble (6.3 to 63 mm), ground (1.45 to
29
2.38 tran), or pulverized (0.84 to 1.49 mm) form. Experimental evidence has
shown an increase in dissolution as the size o£ a lime particle diminishes.
For example, a 100 percent quicklime of 100 mesh (0.149 mm) will dissolve
29
twice as fast as one of 48 mesh (0.35 mm).
Although lime can be fed dry, for optimal efficiency it is slaked
(hydrated) and slurried before use. Slaking is usually carried out at .
temperatures of 82 to 99°C with reaction times varying frca 10 to 30 cinutes.
Following slaking, a wet plastic paste is formed (lime putty) and then
29
slurried with water to a concentration of 10 to 35 percent.
While most line is sold as quicklime, small lime consumers often cannot
economically justify the additional processing step that slaking entails.
Therefore, high calcium-and dolomitic lime are also available in hydrated form
(either Ga(OH) or Ca(OH)2 MgO). This product is made by th'e lime
manufacturer in the form of a fluffy, dry, white powder. It is supplied
either in bulk or in 23 kg (50 Ib) bags. Hydrated lime is suitable for dry
feeding or for slurrying and the resulting purity and uniformity are generally
superior to slaked lime prepared onsite. High calcium hydrate is far more
reactive than dolomitic hydrate, Dolomitic hydrate, which possesses greater
basicity (approximately 1.2 times), is a much slower reactant, although heat
2 g
and agitation can accelerate its inherently slow reactivity.
Both quicklime and hydrated lime deteriorate in the presence of carbon
dioxide and water (air-slaking), Therefore, litne is generally stored within
moisture-proof containers and consumed within a few weeks after manufacture.
The storage characteristics of dry nyarated lime are superior to ouicKlime,
but carbonation may still occur causing physical swelling, marked loss of
chemical activity, and clogging of discharge valves and pipes.
10-28
-------
Dry chemical feed systems consist of either manual addition of 50 Ib bags
or, in large operations where lime is stored in bulk, an automatic mixing and
feeding apparatus. Two types of automatic feed systems are available.
Volumetric feed systems deliver a predetermined volume of lime while
gravimetric systems discharge a predetermined weight. Gravimetric feeders
require more maintenance, are roughly twice as expensive, but can guarantee a
minimum accuracy of 1 percent of set rate versus 30 percent for volumetric
feeders.31
In a typical lime slurry system with storage and slaking equipment,
slurry tank with agitator is used followed by a slurry reeirculation
line. The process flow lines bleed off a portion of the recirculation
slurry to the reactors. The process line is as short as possible (to prevent
caking) and the control valves are located close to the point of application.
Lime neutralization/precipitation operations are typically conducted
32-35
under atmospheric conditions and room temperatures. The precipitation
unit is usually a reinforced tank with acid-proof lining and some sort of
agitation to maintain intimate contact between the metals-containing wastes
and the lime (slurry) solution. Vertical ribs can be built into the perimeter
to keep the contents from swirling instead of mixing.
During operations, adequate venting may have to be provided due to the
possible evolution of heat and noxious gases. Table 10.1.11 presents a
summary of process parameters gathered from various lime slurry precipitation
systems. However, while these provide an indication of typical system design,
testing under actual or simulated conditions is the only sound basis for the
determination of individual waste treatment parameters.
Caustic _spda—Pure anhydrous sodium hydroxide (NaOH) is a white
crystalline solid manufactured primarily through the electrolysis of brine.
Caustic soda is a highly alkaline, sodium hydroxide solution. It is used io
the precipitation of heavy metals and in neutralizing strong acids through the
formation of sodium salts.
Although available in either solid or liouid form, NaOH is almost
"2 £
exclusively used in water solutions of 50 percent or less. The solution
is marketed in either linea 55-gaiion drums or in bulk; i.e., tank car or
truck. As a solution, caustic soda is easier to store, handle and pump,
10-29
-------
TABLE 10,1.11. SUMMARY OF TYPICAL LIMB-SLURRY OPERATING PARAMETERS
Parameter Unitls)
Type of Stone % MgO
Stone Sise mm
Slaking Temperature °C
Slurry Solids 2
Retention Time Min
Sedimentation Time Min
Mineral Acidity Mg/L
Operating range
5 -
0.149 -
B2 -
5 -
5 -
15 -
10,000 -
40
255
99
40
15
60b
- 100,000
Optimum range
5
0.149
Same
a
5
15 - 30
20,080
ais dependent on site specific factors
bHigh calcium Lime will settle in 15 minutes with 1-21 acid wash streams and
30-60 minutes with 3-10 percent acid streams. Dolomite will typically . take
15-60 minutes.
Source: References 10, 28, and 29.
10-30
-------
relative to lime. In comparison to lime alurries, caustic soda will not clog
valves, form insoluble reaction products, or cause density control problems.
However, when sodium hydroxide is stored in locations where the ambient
temperature is likely to fall below 12°C, heated tanks should be provided to
16
prevent reagent freezing.
After lime, sodium hydroxide is the moat widely used alkaline reagent for
precipitation systems. Its chief advantage over lime is that, as a liquid, it
rapidly dissociates into available hydroxyl (OH-) iona. Holdup time is
minimal, resulting in reduced feed system and tankage requirements. Caustic
14 ,
soda's main disadvantage is reagent cost. As a monohydroxide, in
precipitating divalent metals such as nickel, two parts hydroxide are required
per part of metal precipitated. In contrast, dihydroxide bases such as
hydrated lime, only require one part hydroxide per part of divalent metal
precipitated.
This increase in reagent requirements combined with a higher cost/mole
(approximately five times that of hydrated lime), makes caustic soda more
expensive on a precipitation equivalent basis. Generally, in high volume
applications where reagent expenditures constitute the bulk of operating
expenses, lime is the reagent of choice. However, in low volume applications
where low space requirements, ease of handling, and rapid reaction rates are
the deciding factor in reagent selection, caustic soda is clearly superior.
Also, in any system where sludge disposal costs will be high, caustic soda
will compete aore favorably with lime.
The higher solubility of NaOH in water (approximately 100 times that of
lime at 25°C) reduces or eliminates the need for complex slaking, slurryins,
or pumping equipment. In a typical system, caustic is added through an
air-activated valve controlled by a pH sensor. Reagent is demanded as
long as the pH of the waste stream remains below the controller setting
required for precipitation. Agitation is provided by a mechanical mixer to
prevent excessive lag time between the addition of the reagent and the first
observable change in the effluent pH. The precipitated solution is then
10-31
-------
pumped to a large settling tank for liquid/solid separation. Table 10.1,12
provides sodium hydroxide sludge generation factors for seven metallic specie's
commonly encountered in metal-containing wastes,
The precipitation reaction is typically carried out under standard
operating temperatures and pressures. ' The reaction is almost instantaneous
since caustic soda reacts vigorously with water.' At concentrations of
40 percent or greater, the heat generated by dilution can raise the
temperature above the boiling point. Handling precautions are required when
performing dilution or other reagent handling since even moderate
concentrations of NaOH solution are highly corrosive-to skin.
Process configurations for caustic soda treatment are a function of waste
type, volume, and raw waste pH level and variability. For 'example, • the
precipitation'of concentrated acidic roetals-laden waste streams with low dead
times depends on pH as follows; one reactor system for feeds with ,pH ranging
between 4 and 10, a reactor plus a smoothing tank for feeds with pH
fluctuations of 2 and 12, and two reactors plus a sraoothing tank for feeds
with pH less than 2 or greater than 12. Retention times vary with the rate
of reaction and mixing, however, 15 to 20 minutes appears to be optimal for
37
complete neutralization/precipitation in nose systems. The interva-1
between the addition of sodium hydroxide and the first observable change in
effluent pH (dead time) should be less than 5 percent of the reactor residence
time in order to maintain good process control. A summary of typical
operating parameters is provided in Table 10.1.13,
A typical caustic system is designed to add most of the reagent in a
preliminary precipitation stage,1 while a second stage acts as a smoothing and
finishing tank. In this manner, the second reactor is able to compensate for
pH control overshoots or concentrated batch dumps which nay temporarily
overwhelm the primary precipitation system.
Overshoot is due primarily to the lack of sodium hydroxide solution
buffering capacity. For example, Figure 10.1.5 illustrates the titration
curve for the neutralization of a ferric chloride etching solution CpH 0.5)
with a 5 Molar caustic soda solution. The steep slope of the titration curwe
beginning at pH 2.0, combined with a strong demand for alkali prior to that
pc.ir.tj often make over- or under-correct ion unavoidable* Far continuous
precipitation applications of greater than 20 gptn, pH control in Che portion
of tne titration curve wnicn is nearly vertical (between pH 2,0 and 9.0) is
achieved in a second reactor, to prevent excess reagent usage or effluent
discharge violations. 10-32
-------
TABLE 10.1.12, SODIUM HYDROXIDE SLUDGE GENERATION FACTORS
Metal ion
Cr
Ni
Cu
Cd
Fe
Zn
Al
Ib dry solids generated
Ib of metal precipitated
1.98
1.58
1.53
1. 30 •
1.61
1.52
2.89
Source: Reference 14,
TABLE 10.1.13. SODIUM HYDROXIDE NEUTRALIZATION: SUMMARY OF TYPICAL
OPERATING PARAMETERS
Parameter
Unit(s)
Operating range
Ideal range
Sodium hydroxide
concentration
Dead time
Retention time
Batch treatment
throughput
Continuous treatment
throughput-
Suspended solids
Storage temperature
(40-501 NaOH)
% NaOH
% Retention time
Minute
gal/min
gal/min
Weight 1
12 - 50
3-10
5-30
1-20
15
3-10
12 - 20
40 - 50
3-5
15 ~ 20
20
20
10
16 - 20
Source: References 4, 7, and 3a.
10-33
-------
9 r
8 -
6 -
0 TRJAL 1
a TRIAL 2
_u
0 .- 10 20 30 40 50
VOLUME OF 5 MOLAR NoOH (mil
Figure 10.1.5, Neutralization of ferric chloride etchinc
waste by sodium hydroxide.
10-34
-------
10.1.2.2 Process Performance™
Chemic&l precipitation of metal hydroxides through the use of line and
sodium hydroxide is a. classical waste treatment technology used by most
industrial waste treatment systems. The performance of these technologies in
removing metallic pollutants from industrial uaatewaters is well documented in
the literature. Tables 10.1.14 (lime precipitation) and 10.1.15 (NaOH
precipitation) contain general performance indicators which incorporates
effluent concentrations and removal efficiencies developed from plant-specific
full-scale and pilot plant data bases.
In recent years, research has centered around the evaluation of
supplemental chemicals to the already well defined hydroxide precipitation
sedimentation process. Organic and inorganic polyelectrolytes (see
39
Section 10.2), acid, and soda ash (see Section 10,1.4) have all been used
in this capacity. The purpose of these supplementary chemicals is to improve
the efficiency of liquid-solid separation break complexing/chelsting agents,
and take advantage of the lower solubility of carbonate complexes.
Process wastes containing compLexing/chelating agents are often
untreatable with established technologies. The difficulty arises due to the
formation of a highly stable organo-metallie bonds formed between the metal
ion and the complexing/chelating agent. Ammonia is an example of a complexing
agent, with each molecule of ammonia bound to a metal species such as copper
by a single bond. A chelating agent such as EDTA, on the other hand, forms
more than one bond with each metal ion. Complexing and chelating agents are
typically used to keep the metals in solution for plating. During rinsing,
the complexed and chelated metals end up in the processing rinsewater. The
major complexing agents found in metal waste streams are ammonia, cyanide,
fluoborate, and pyrophosphate. The foremost chelating agents are EDTA,
40
Quadrol, citrate, and tartrate.
Established chemical methods for breaking chelator/complexes and removing
metals to low concentrations are starch xanthate, sodium DTC, ferrous sulfate,
waste acids, sulfide ions, sodium hydrosulfite, sodium borohydride, and high
40
pH lime. A typical process used by industry is the combination waste
acid-high pE lime treatment method. For this type of waste treatment process
10-35
-------
TABLE 10.1.14. PERFORMANCE SUMMARY FOR LIME PRECIPITATION OF HEAVY METALS
Metallic species
Arsenic
Cadmium
Chroicium
Copper
Cyanide
Lead
Mercury
Nickel
Selenium
Silver
Thallium
Zinc
Effluent concentration (tng/L)
NO -
ND -
ND -
ND -
ND -
ND -
O.i -
ND -
ND -
ND -
1.1 -
13 -
110
BO
1,800
220
5,500
580
43
5,200
8?
90
20
26,000
Removal efficiency (%)
20
20
4?
33
67
0
69
6
40
99
58
25
_ >99
_ >99
- >99
- >99
- >99
- >99
- >96
- >99
- >99
- >99!
- >75
- >99
Source: Reference 1,
TABLE 10.1.15. PERFORMANCE SUMMARY FOR SODIUM HYDROXIDE
PRECIPITATION OP HEAVY METALS
Metallic species Effluent concentration Cmg/L)
Cadmium
Chromium
Copper
Lead
Nickel
Silver
Zinc
Hexavalent chromium
ND -
18 -
1.0* -
-
ND -
11 -
44 —
ND -
930
3,000
5,900
210
64
560
25 •
Removal efficiency (%)
22
53
36
>99
>99
76
80
73
- >99
->99
- 98
- >99
- >99
*Approximate value.
Source: Reference 1.
10-36
-------
the pH of Che organo-metallie waste is first adjusted to approximately 2 with
dilute mixed waste acid (sulfuric, nitric, hydrochloric, or chromic acid)
and/or virgin hydrochloric acid. After the chelatar/complex breaking step the
pH of the waste solution is raised to approximately 11, resulting in the
39
formation of insoluble metal hydroxides. Table 10.1.16 presents alternate
precipitation technologies for the removal of metals such as copper from
complexed and chelated rinaewatera.
An alternate technology for the precipitation of metal hydroxides which
has shown promise in recent years is magnesium oxide (MgQ). Magnesium oxide
is available in slurry form composed of 55 to 60 percent magnesium hydroxide,
Mg(QrtK. The slurry has & bulk density of 1.5 kg/L and due to its low
41
solubility (0.0009 g/100 mL), must be mildly agitated during storage.
The main advantage of magnesium hydroxide over comparable hydroxide
precipitation technologies is that the precipitate formed is .mote particulate
in nature (due to longer reaction times). The sludge formed has better
handling and dewatering characteristics and sludge volumes are much less.
Table 10.1.17 compares typical physical, chemical, and filtered sludge
properties of magnesium hydroxide to those of caustic soda and hydrated lime.
As can be seen, dewatering characteristics and filtration time for separating
41
solids are considerably enhanced in the case of magnesium hydroxide.
The main disadvantage of magnesium hydroxide is that it costs
approximately three times as much as hydrated lime. In addition, operation of
^
the magnesium hydroxide system is not as straightforward as comparable
hydroxide systems. Reaction times are slower and it will be necessary to make
modifications in waste treatment operating procedures and equipment.
Table 10.1.18 presents the results of a Bureau of Mines research effort
into magnesium oxide precipitation of metals. The researchers found that when
equal pH values are obtained, MgO leaves less dissolved metal and less
suspended metal hydroxide of sedimentation as part of the process. MgO was
able to remove any metal that is precipitated as «• hydroxide. However, a
threefold to fourfold stoichiometric excess was required to reach adequate pH
values (8-9).
10-37
-------
TABLE 10.1.16. CHEMICALS FOR Cu REMOVAL PROM COMPLEXED
AND CHELATED RINSEWATERS
Ammonium
Precipitation Alkaline persulfate Electrole'ss Pyroohosphate Fluoborste
chemical etchants etchants Cu Cu Cu
Insoluble starch
xanthate (1SX)
Sodium dimethyl-
dithiocarbamate (DTC)
Ferrous suifate
Spent pickle liquor
Ferrous sulfide
Sodium hydrosulfite
High-pH line
Sodium borohydride
X
X
X
X
X
X
X
X
X
X X
X X
X
. x
X XX
X
Source: Reference 40.
10-38
-------
TABLE 10.1.17. . COMPARISON OF HYDROXIDE REAGENT PROPERTIES
Property
Molecular weight
Hydroxide content (%)
Heat of solution (Kg-cal/mole)
Solubility (g/100 mL H20)
Reactive pH maximum
Weight equivalency
Freezing point
Solids content of sludge (%}£
Sludge density (lb/ft3)
Filtration time (hr)
Sludge volume (yd3/10,OOQ Ib)
NaOH
40.0
42.5
9.94
42. Oa
14.0
1.37
16. Oc
30.0
80.0
7-3
15.0
Ca(OH)2
74.1
45.9
2.79
0.1853
12.5
1.27
0.0d
35.0
85.0
7-9
12.58
Mg(OH)2
58.3
58.3
0.0
0.0009b
9.0
1.0
o.oe
55.0
100-110
1.5-2.0
6.4
aTemperature, 0°C.
bfemperature, 188C.
C50 percent solution.
^30 percent slurry.
e58 percent slurry.
^Sludge from a plate-and-frame filter press,
Sconsists of metal' hydroxide and gypsum.
Source: Reference 41.
10-39
-------
TABLE 10.1.18. TEST RESULTS FROM TREATING METALS-BEARING WASTEWATER
WITH MgO AND LIME
Chemical analysis, ppra
pH
Fe
Cu
2n
Si
Mn
Co
Cd
Pb
Benef iciat ion orocess «ater-CM
Untreated water
Treated with:
0.1 g/Lb MgO,
0.2 g/L MgO,
0.35 g/L MgO,
filtered
filtered
filtered
5.
8.
9.
9.
L
6
2
4
5.7
0.2
0,2
0.2
Q.63
0. 1
0.1
0. 1
0.55
0. 1
0.1
0.1
NDa
ND
ND
ND
9.9
7. 1
3.4
1.3
ND
ND
ND
ND
ND
ND
ND
ND
ND
ND
MD
ND
Benef iciation proces's waste-BK
Untreated water
Treated with:
0. 16 g/L MgO,
0.21 g/L MgO,
0.31 g/L MgO,
Mine drainage
Untreated water
Treated with 0.5
Prepared solution
Untreated water
Treated with:
0.4 g/L MgO,
0.4 g/L MgO,
0.1 g/L line,
Prepared solution
Untreated water
Treated with:
0.125 g/-L MgO
0.04 g/L lime
Prepared solution
Untreated water
Treated with:
filtered
filtered
filtered
g/L MgO, filtered
No. 1
filtered
settled
c settled
No. 2
, filtered
,c filtered
No. 3
0.2 g/L Mg, settled
0.05 g/L lime
,c settled
6.
8.
8.
8.
2.
8.
4.
8.
8.
9,
5.
8.
8.
4.
9.
9.
4
3
7
9
1
Q
2
9
9
4
4
9
9
0
0
0
0.2
ND
ND
ND
40
0.2
ND
ND
ND
ND
5,0
0,2
0.2
ND
ND
ND
0,1
ND
ND
ND
ND
ND
8.7
0.1
0.5
0. 7
0.21
0.1
0, 1
rro
ND
ND
12.7
0.2
C.2 •
0.2
39
0. 1
ND
ND
ND
ND
2.7
0. 1
0.2
4.2
G. 1
0.8
0.2
ND
ND
ND
N*D
NTS
12.0
0.2
0.2
1.2
ND
ND
ND
ND
ND
ND
17.5
8.3
5.7
1.9
41
15
ND
ND
ND
ND
4.4
0.2
2.2
HD
ND
ND
ND
ND
ND
ND
ND
HD
11,0
.0.2
0.2
1.6
ND
ND
ND •
ND
ND
ND
ND
MD
BD
ND
ND
ND
BD
ND
ND
ND
ND
ND
m
5.2
0.31
1.4
ND
HD
ND
ND
ND
ND
ND
ND
ND
ND
ND
. ND
ND
4.7
O.S
1.6
aThe unit g/L is the grams of MgO used per liter of water treated.
ND - not determined, since initial concentrations were below the analysis limit of
atomic absorption.
cWeight of lime is for CaO, not Ca(OH)2.
Source; Reference 42.
10-4U
-------
The researchers concluded that when influent metals content is low,
increased chemical costs will be balanced by savings from easier sludge
dewatering, compactness, and stability. It is anticipated that as land
disposal costs for metal hydroxide sludges continue to increase, the economics
of this process will become more favorable. In addition, by mixing magnesium
hydroxide with sodium hydroxide in a dual reagent system, sludge reductions of
approximately 45 percent can be realized. Although alkali costs will
increase, savings in sludge conditioning polymers and disposal costs will help
43
to defray the added reagent expense.
10.1.2.3 Process Coats—
The basic equipment train for a>< hydroxide precipitation system consists
of a collection sump, piping system, precipitation reactor, feed system,
fiocculation/ciarification unit, sludge storage tank(s), and plate and frame
filter press* Figure 10,1.6 illustrates the treatment train design used for
the remainder of this section.
The capital and annualized cost information contained in this section was
44
adapted from costing methodology developed by Versar. Versar calculated
the direct costs, indirect costs, and working capital as a percentage of the
purchased equipment and installation (PE&I) costs. The cost elements and
assumptions made by Versar are summarized below:
Assumed value
Cost elements (% PE&I)
Direct costs (DC)
Instrumentation and controls 10
Piping 21
Electrical equipment and materials 13
Buildings 26
Yard improvements 7
Service facilities 41
Total direct cost: 118
10-41
-------
WASTE WATER
EQUALIZATION
TANK
REACTION
TANK
SLURRY OR
CAUSTIC SOLUTION
CLARIFIES
OVERFLOW
UNDERFLOW
FILTER
AQUIOUS
PHASE
TREATED
SLUDGE
pH ADJUSTMENT
EFFLUENT
ENCAPSULATION
ULATED V.ATsSIAL
TO LAND DISPOSAL
jA.iiie, jx*yiuiroxivi6: precipitation.
Source; Reference 44.
-------
Indirect costs (1C)
* Engineering and supervision 29
* Construction expenses 32
• Contractor's fees 7
* Contingency _JE7
Total indirect cost: 95
Fixed capital investment (FCI) PE&I * DC * 1C
Working capital (WC) 47
Total capital investment (TCI) FCI + WC = 3602 PE&I
Annualized costs included variable costs, plant overhead costs, general
and administrative expenses, and fixed costs* The variable costs included
costs for labor, maintenance, materials, chemicals, and hazardous contracted
sludge disposal. The fixed costs include taxes, insurance, and capital
recovery costs.
The chemical requirements for each treatment were based on stoichiometric
requirements. The cost for chemicals were obtained from the Chemical
Marketing Reporter. In deriving annualized costs, a certain set of
14 44-47
assumptions were made. These assumptions are listed below: *
• Plant overhead operating costs are 5.8 percent of the total capital
investment costs.
* Taxes and insurance costs are 1 percent of total capital investment
costs.
* Labor costs are based on 4 hrs/sbift at $20/hr.
• Maintenance costs are at 6 percent of total capital investment costs.
* Power costs are at 2 percent of total capital investment.
• The nonhazardous contracted sludge disposal costs are based on
$200/ton.
• The sludge transportation costs are based on SQ.25/ton-mile and a
transportation distance of 15 miles. It is further assumed that all
hazardous solid wastes generated by the treatment processes uauld be
encapsulated and disposed of as nonhazardous wastes.
10-43
-------
In all cases, capital.recovery was calculated at a 12 percent interest
rate over a period of 10 yearsi The capital recovery factor (CRF) was
estimated as follows:
(1 + i)n- 1
= 0.177
where:
i = interest rate
n = number of years,
Costs for items which were not in the size range of available information
46
extrapolations were made using the following equation:
_ . _ Capacity A
Cost A = cose B —-—. J ,
Capacity B
The exponent, "x", was determined with available information and is presented
where necessary in the footnotes to the cost summary tables.
The capital costs for the base equipment in hydroxide treatment train
have been adapted from cost figures and tables contained in "Reducing Water
Pollution Costs in the Electroplating Industry." The base system discussed in
this section is designed to handle an aqueous waste stream containing 200 mg/L
of heavy metal ions. The flow rates for this system were developed for three
different sizes: 1,000, 10,000, and 100,000 gal/hr. Operation of this system
is assumed to be 24 hrs/day, 300 days/year.1 Mixed reactor construction costs
for the first stage flow/concentration equalisation tank are presented in
Figure 10.1.7. The equalization tank has been sized for 1-hour of retention
time and was fabricated from reinforced concrete.
The precipitation reaction tank consists of a continuous neutralization/
precipitation vessel equipped with pH control, reagent storage, and reagent
feed systems. The reagent feed and storage system is sized for a 1-week
supply and uses hydrated lime as a precipitation agent. Sulfuric acid
capability is also included in case of pH overshoot. Figure 10.1.8 presents
the mixed reactor construction and installation costs. The reactor has been
sized for a minimum of 30 minutes of retention time to ensure a complete
precipitation reaction.
-------
o
-p-
Ul
O
O
O
o
- 10
r~ CO
Ufll
c/)
Z
O
o
40
30
20
10
0
J I I I I I I
I
5 10 20
VOLUME , M3
_L_LJ_LLjJ
5O JOO
Figure 10.1.7. Construction costs for reinforced concrete equalization reactor.
Source: Reference 7.
-------
L
c
o
9,
w *a R
** ^^
•" t.1
§5
I MINIMUM UNIT
SiZE
21 —
i
£0 40 SC 60
FLOW RA7E,jsl/nin
ICO
120
Figure 10.1.8. Investment cost for continuous single-stage
precipitation.
10-4b
-------
Assuming complete reaction between the heavy metal ions in the waste
stream and the hydroxide ions in the precipitation reactor, some sort of
separation will be necessary to remove the metal hydroxides and other
insoluble pollutants from the reactor effluent. Figure 10.1.9 shows the
hardware and installed costs for a flocculation/clarificat.ion unit used to
enhance the settling characteristics of the suspended solids. The unit is
assumed to have a separate flocculation tank, a polymer feed system, a
"lamella" or slant-tube separator, and a zone in which the sludge collects
before being discharged. The costs like those of the equalization and
precipitation reactors are a function of flow rate. The solids concentration
of the underflow is assumed to be 2 percent, while the overflow is assumed to
be solids-free.
Typically, the underflow from the flocculation/clarification unit must be
stored onsite in sludge holding tanks before the sludge is shipped to a
disposal site or transferred to another dewatering stage. The investment cost
for sludge tanks is presented in Figure 10.1.10. The tanks are of carbon
steel construction and the cost is a function of tank volume. The sludge
holding tanks for the hydroxide treatment base case have been sized for
14
10 hours of clarifier underflow.
In many cases, further concentration of thickened sludge through the use
of mechanical dewatering equipment is desirable in reducing sludge disposal
costs. Figure 10.1.11 presents the unit costs for a recessed plate filter
press as a function of the feed volume capacity (filter cake volume is also
given). The feed solids concentration in this case is assumed to be
2 percent, the cake solids concentration is 20 percent, with an 8-hour
14
press-cycle. Items not included, but will contribute to the cost of
installation include:
» High pressure feed pumps;
* Filtrate return lines (to clarifier); and
* Cake solids handling equipment.
Table 10.1.19 details the costs developed for the continuous h«drsted
lime precipitation system previously shown in Figure 10.1.10. The high cost
of the 1,000 gph, continuous system relative to the ether two pretreatasent
10-47
-------
30
10 l-
Noses;
insisiled COjt « 1 .25 x hs/
Cost incigdes Diate-woe
cost ,
r wsih flo=cuUung
SO 100
I
150
FLOW RATE i
Figure 10.1.9.. Investment cost for flocculation/clarification units,
10-48
-------
I
S »<»
B*jeQ on ciftxjn tttel constru
Cestf nciude fi&«r-iretfWorce<3
Figure 10.1*10* Investment cost for sludge storage/thickening units,
Source: Reference 14.
2
O
> 200
sea
in
20 30 40
EQUIPMENT COS! ($t,QGO!
Ccsi mclodticartaai sieel Ifpme. ooiy-
progylane ptsi*s tfta &Mer CiQins,
Cat
cat
Fee
fiaiien casts n&i inciMCee.
volume bijdd on \ V," ihiek i
a
votumt) ciasciiy aaseo an*. I*
i * 2%; eahesoiidf * ?0*'«; e
* I haun.
Figure 10,1,11. Hardware cost for recessed plate filter presses,
Source: Reference 14.
10-49
-------
TABLE 10.1.19. CONTINUOUS HtBRATED LIME PRECIPITATION COSTS3
Purchased equipment ' and installation (PE&I)
Equilazation tank
Precipitation reactor
Flocculator/ciarif ier
Sludge molding tank(s)
Filter Press
Total capital investment (360% PEI)
Operating costs ($)
Operating labor ($2Q/hr)
Maintenance (6% TCI)
General plant overhead (5.8% TCI)
Utilities (2% TCI)
Taxes and insurance (1% TCI)
Chemical costs (£4Q/ton)
Sludge transportation ($0,25/ton-mile)
Sludge disposal ($200/ton)
Annualized capital (CFR = 0.177)
Total cost/year
Cost/1, 000 gallon
L.OOO
(t)
17,000
24,000:
18,000
3,000'
_IO_, OOP
72,000
259,000
72,000
15,500
15,000
5,200
2,600
500
200
12,000
45,800
168,800
23
Flow rate (gph)
10,000
29,000
40,000
50,000
6,000
25jOOO
150,000
540,000
72,00,0
32,500
31,400
10,800
5,400
5,300
2,300
120,000
95,800
375,500
5
100,000
50,000
69,000
140,000
48,000
100,000
407,000
1,465,200
72,000
87,900
85,000
29,000
14,700
53,000
22,500
1,200,000
259,000
1,823,100
2.5
Source: Reference 44.
a1987 Dollars.
10-50
-------
processes illustrates why precipitation systems under 50,000 gpd (2,000 gph)
are typically batch in nature to reduce equipment costs. In addition, the
large costs attributed to sludge disposal in every system demonstrates the
main drawback to hydroxide precipitation. As land disposal costs increase,
treatment processes such as hydroxide precipitation, which generate large
quantities of hazardous sludge will lose their cost advantage over the more
expensive recovery technologies.
10,1.2.4 Process Status—
Hydroxide precipitation is a widely used and well developed technology
for reducing metals effluent concentrations to acceptable levels. The process
operates at ambient temperature and pressure and is well suited to automatic
control. Its ability to treat a wide variety of industrial waste streams has
been well demonstrated in bench! pilot, and full-scale systems. Environmental
impacts can result from emissions during the precipitation process and the
48
production of large volume of potentially hazardous sludge. Exit gases
can be scrubbed by using a control system, however, sludge reduction methods
48 49
(seeding, dilution, vacuum filtration, etc.), " have only partially
offset the problems associated with sludge generation. Therefore, new methods
of sludge disposal and reduction and recycle/reuse options (such as
agricultural liming) should be considered. The advantages and disadvantages
of lime-precipitation and caustic soda precipitation are summarized in
Tables 10.1.20 and 10.1.21, respectively.
10.1.3 Sulfide Precipitation
10.1.3.1 Process Description—
The basic principle of sulfide precipitation is similar to that of
hydroxide treatment, in that the precipitation process converts soluble metal
ions into insolubte (sulfide) compounds. Some advantages over hydroxide
precipitation are that with sulfides, heavy metals can be removed to extremely
low concentrations at a single pH. In addition, the use of sulfides allow
precipitation of contaminants, even in the .presence of chelating agents.
Sulfide precipitatiar. has beer, lisiced :c relatively few applications,
however, due to the toxicity and odor of hydrogen sulfide (H.S) evolved from
.... 1 l
tne precipitation process.
10-51
-------
TABLE 10.1.20. ADVANTAGES AND DISADVANTAGES OF LIME PRECIPITATION
Advantage^
- Proven technology with documented neutralization efficiencies,
No temperature adjustments normally necessary.
- Modular design for plant expansion.
Can be used in different configurations.
- Able to coprecipitate a mixture of metal ions to achieve residual
metal solubilities- lower than that achieved by precipitating each
metal at its optimum pH.
- Reagent is easy to handle, and has treatment effectiveness for wide
range of dissolved materials,
Disadvantages
- The theoretical minimum solubilities for different metals occur at
different pH-values. For a mixture, of metal ions, fct must be
determined whether a single pH can produce sufficiently low
solubilities for the metal ions present in the wastewaters.
Hydroxide precipitates tend to resalubilize if the solution pH is
increased or decreased from the oinitnuin solubility point; thus
maximum removal efficiency will not be achieved unless the pH is
controlled within a narrow range.
- The presence of complexing ions, such as phosphates, tartrates,
ethylenediaminetetraacetic acid (EDTA), and ammonia may have adverse
effects on metal removal efficiencies when hydroxide precipitation
is used.
- Hydroxide precipitation usually makes recovery of the precipitated
metals difficult because of the heterogeneous nature of most
hydroxide sludges.
10-52
-------
TABLE 10.1.21. ADVANTAGES AND DISADVANTAGES OF'CAUSTIC SODA PRECIPITATION
Advantages
- Proven technology with documented neutralization efficiencies
— Strong alkali with rapid reaction rate
- Smaller tanks and retention times than comparitive reagents
- Inventory and storage handling procedures are less complicated due
to liquid form
- Storage does not require continuous agitation to maintain homogeneity
- Does not require complex slaking or slurrying equipment
- Produces more soluble by products in low pH applications
Disadvantages
Chemical costs are significantly higher ($205/ton vs. $46/ton for
hydrated lime)
- Does not impart any buffering capacity to industrial waste streams
Close attention must be given to the design of che pH control
- Caustic soda precipitation will result in a fluffy gelatinous floe
increasing"the size of the clarification chambers and sludge
dewatering equipment.
Cannot effectively precipitate sulfate waste streams due to
solubility of sodium sulfate.
Source:- Reference 3.
1U-53
-------
Sulfide precipitation is used to remove lead, copper, silver, cadmium,
z.inc, mercury, nickel, thallium, arsenic, antimony, and vanadium frore
wastewaters. Typically, the precipitation reaction is conducted under near
neutral conditions IpH 7.0 to 9.0). Exceptions to this rule are arsenic and
antimony which require a pH below 7 for optimum precipitation. As with
hydroxide treatment, cyanides are usually oxidized prior to precipitation.
The first step in the sulfide precipitation process is the preparation of
a sodium sulfite solution. The solution is then added to the reaction tank
44
(30 minutes retention time) in excess to precipitate the pollutant metal
as illustrated in the following reaction:
NiSQ4 = NiS + Na2S04 CD
Sodium Nickel Nickel Sodium
Sulfide Sulfate Sulfide Sulfate
The process is controlled by means of a feedback control loop employing
ion-selective electrodes. Physical separation of the metal sulfide takes
place in thickeners or clarifiers, with reducing conditions maintained by
excess sulfide ions. The final step is usually oxidation of the-excess•
44
sulfide ions through aeration or hydrogen peroxide addition. Currently,
two methods of delivering sulfide ions to the process reactor are available.
The first method utilizes soluble-sulfides such as a sodium sulfide (Na^S)
or sodium hydrosulfide (NaHS). A second method (Sulfex process) uses a
sparingly soluble metal sulfide such as ferrous sulfide (FeS) as a source of
sulfide ions. Each process will be discussed individually in the following
subsections.
Soluble sulfides—Pure sodium sulfide (sodium-sulfuret) is a white,
crystalline solid (mp 1180°C, sp gr 1.856). Commercial material is white
to light yellow or pink. it crystallizes from aqueous solutions as the
nortahydrate, Na S.9KLO. In air, sodiuffi sulfide slowly converts to sodium
carbonate and sodium thiosulfate and is deliquescent. .Reactions with strong
oxidizing agents give elemental sulfur.
10-54
-------
Pure sodium hydrosulfi.de (sodium sulfbydrate, sodium hydrogen sulfide,
sodium bisulfide) is & white, crystalline solid (mp 350°C, sp gr 1.79). It is
highly soluble in water, alcohol, or ether. The commercial product occurs in
different shades of yellow and is highly deliquescent. Exposure to air
converts it to sodium thiosulfate and sodium carbonate. In the presence of
organic matter, combustion can occur. Heating releases hydrogen sulfide,
which is a toxic gas.
Sodium sulfide is marketed as 30 to 34 weight percent fused crystals and
60 to 62 weight percent flakes. Each container has a corrosive label and
a product label stating that the product causes severe burns to eyea or skin,
and that contact with acid liberates poisonous hydrogen sulfide gas. The
material is nonflammable, noncombustible, and nonexploaive. Sodium
hydrosulfide is marketed as 70 to 72 weight percent flakes and 44 to 60 weight
percent liquor in the high purity grades, and as 10 to 40 weight percent
liquor from recovered caustic wash in the oil-refining desulfurization
processes. Shipment labeling is the same as for sodium sulfide. The product
is shipped either as flake in drums or as solutions in tank cars or tank
trucks.
The lower freezing points o£ solutions of sodium hydrosulfide provide an
advantage over those of sodium sulfide in shipping by tank truck and tank
car. Recently, systems have been designed to enable customers to make their
own sodium sulfide solutions by reaction of NaHS and NaOH.
The high solubilities of sodium sulfide and sodium hydrosulfide eliminate
the need for slaking and slurrying apparatus. Reagent is added either from
storage in the case of liquid reagents or from rapid-mix tanks when using
2 52
solid reagents. * Reagent demand is determined through a specific-ion
sulfide reference electrode pair, which is set to a preselected
potential. Normally, sulfide reagent demand depends on the total metal
concentration contained in the effluent waste stream. For continuous
processes where metals concentrations are constant, electrode set points can
be set at the potential which corresponds to the maximum electrical
pocential-sulfide concentration gradient and where the wastewater solution has
2.50
the least detectable odor. For batch processes, a simple jar test prior
ior. car. accurately determine optimal sulfide dosages.
10-55
-------
Since sodium sulfide and sodium hydroeulfide have such high solubilities,
dissolved sulfide concentrations are correspondingly high. This high
concentration of dissolved sulfide causes a rapid precipitation of the -metals
dissolved in the water asitnetal sulfides. However, it often results in the
generation of small particle fines and hydrated colloidal particles. The
rapid precipitation reaction tends more discrete particle precipitation than
toward nucleation precipitation (the precipitation of a particle from solution
onto an already existing particle). The resulting pool—settling or
poor-filtering floe is difficult to separate from the wastewater discharges,
This problem has been solved by the effective use, separately or combined, of
coagulants and flocculants to aid in the' formation of large, fast-settling
. . ,. 53
particle floes.
One major disadvantage of the soluble sulfide precipitation method is the
formation of hydrogen sulfide (H S3 from dissolved sulfide ions.
Figure 10.1.12 is a graph developed by Centec Corporation for determining the
percentage of the dissolved sulfide in the forn of H.S as a function of the
pH of the solution. According to Cetitec, the relationship shows that at a pH
of 9, H^S accounts for only 1 percent of the free sulfide in solution.
The rate of evolution of H?S from a sulfide solution per unit of water/air
interface will depend on the temperature of the solution (which determines the
H^S solubility), the dissolved sulfide concentration, and the pH. In
practice, considering typical response lags of instruments and incremental
reagent addition, control of the level of dissolved sulfide and pH would
require fine tuning and rigorous maintenance to prevent an H_S odor problem
52
in the work area. In currently operating treatment systems, the H^S
odor problem is eliminated by enclosing and vacuum evacuating the process
vessels.
Insoluble-sulfides~~Tbe insoluble—su-lfide (Sulfex) process precipitates
dissolved metals by mixing the wastewater with an FeS slurry in a solid/liquid
contact chamber. The FeS dissolves to maintain the sulfide ion concentration
2
at a level of 2 mg/L. Due to its instability, the ferrous sulfide has to
be generated onsite from sodium sulfide and ferrous sulfate. The sulfide is
released from ferrous sulfide only when other heavy metals with lower
equilibrium constants for their sulfide form are present in solution (see
Table 10.1.22).
10-56
-------
100
10
2
o
£0
W
S
0.1
0.01
pH frt jjK = 7)
8 9
I
{pH - r>K)
Nste.—pK (logarithmic practical ionizalion cotisiantj is used to measure the degree of
Specific eltciriesl conductance
of soluiton at 77" F
of pK
50s F 68" F 1 CM" F
0'
100
1.000 ...
50,000" .
7.24 7.10 6.82
7.22 7.08 6.80
7.18 7.04 6.76
7.09 6.95 6,67
"Oisiilled H2O.
Figure 10.1.12. Percent cf dissolved sulfide in the H2S forta.
Source: Reference 52.
10-57
-------
TABLE 10.1.22. SOLUBILITIES OF SULFIDES
Metal
sulfide
Manganous sulfide
Ferrous
sulfide
Zinc sulfide
N Lckel
sulfide
Stannous sulfide
Cobalt
sulfide
Lead sulfide
Cadmium
Silver
Bismuth
Copper
sulfide
sulfide
sulfide
sulfide
Mercuric sulfide
(64°
1.
3.
1.
1,
1.
3,
3.
3.
1.
1.
8.
2.
to
4 x
7 x
2 x
4 x
0 x
0 x
4 x
6 x
6 x
0 x
S V
J A
0 x
5p
77
10
10
10
10
10
10
10
10
°F)a
-15
-19
-23
-24
-25
-26
-28
-29
10-49
10
-97
10-45
10-49
Sulfide
concent rat ion
(tnol/L)
3
6
3
1
3
1
1
6
3
4
9
4
.7 x.
.1 x
.5 x
.2 x
.2 x
.7 x
.8 x
.0 x
.4 x
.8 x
.2 x
.5 x
10
10
10
10
10
10
10
10
10
10
10
-8
-10
-12
-12
-13
-13
-14
-15
-17
-20
-23
10-25
aSolubility product of a metal sulfide, Ksp, equals
the product of the molar concentrations of the metal
and sulfide.
Source: References 54 and 55.
10-58
-------
When the pH is maintained between 8.5 and 9, the liberated iron will form
a hydroxide and precipitate as well* The unreacted ferrous sulfide is
filtered or settled out with the metal sulfide precipitate, while the effluent
is practically sulfide free. Anionic polymers aid settling of metal sulfide
precipitates. The sludge is easily dewatered by conventional techniques. In
chelated systems, a 4-molar excess of ferrous sulfide is required to obtain
maximum heavy-octal removal (see Table 10.1.23 for operating parameters).
The following reactions occur when FeS is introduced into a solution
containing dissolved metals and metal hydroxide:
FeS = Fe*2 + S~2 (2)
M+2 -*- S"2 - MS (3)
M(OH)2 - M*2 + 2(OH)~ (4)
Fe*2 * 2(OH)~ = Fe(OH)2 (5)
The addition of ferrous ions to the wastewater and their precipitation as
ferrous hydroxide (FetOHKJ results in a considerably larger quantity of
solid waste from this process than from a conventional hydroxide precipitation
process.-
When the Sulfex process was compared to hydroxide precipitation in a
series of jar test studies and pilot plant demonstration tests, the following
2
conclusions were reported.
* When treating the same influent, the Sulfex process obtains lower
residuals of copper, cadmium, nickel, and zinc than can be obtained
with the hydroxide process.
* Satisfactory effluent quality is usually obtained with the Sulfex
process witnin the 8,5 to 9.0 pH range which is within the 6.0 to
9.5 pH range permitted by EPA for discharge.
* The removal of a particular heavy metal is more effective when it is
in a solution containing other heavy metals than when it is the only
metal in solution.
10-59
-------
TABLE 10.1,23, SUMMARY OP TYPICAL INSOLUBLE-SULFIDE PRECIPITATION
OPERATING PARAMETERS
Parameter
Reaction temperature
Reaction pH
Reagent excess
Influent metal
concentration
Retention time
Sedimentation rate
Operating Optimum
Unit range range
°C Eoora ' . -
S.D. 6,0 - 9.0 8.5 - 9.0
% 0 - 400 100 - 300
mg/L 1 - 500 20 - 50
Minute 30 - 60 30
fcpm/ft2 0-2 28
aTut»e settler.
Source: Reference
10-60
-------
The Sulfex process can be applied in precipitators (and similar
devices) at surface rates up to 2.0 gpm/fc* when tube settlers are
used.
The required dosage of ferrous sulfide reactant is dependent upon
the type of waste being created. It should normally vary from about
1.5 times theoretical requirement for wastes with no coaplexing
agents to three or more times theoretical for wastes containing
completing ageots.
The concentration of settleable ferrous sulfide solids in the nixing
zone, the pH of the process, and use of certain polyelectrolytes are
important to obtaining satisfactory results in the Sulfex process.
It nay be more economically desirable to pretreat wastes containing
high concentrations of dissolved heavy metals (i.e., a total heavy
tcecal concentration greater than 50 mg/L) by hydroxide before
polishing wich Sulfex.
10,1.3.2 Process Performance—
While not as prevalent as hydroxide treatment, sulfide precipitation has
seen increasing usage in recent years due to improvements in both reagent
dosage and hydrogen sulfide emission controls. The following are
illustrations of soluble, insoluble, and calcium sulfide precipitation
technologies.
Soluble sulfide precipitation—At a 37 gallon/minute (gpm), industrial
pretreatnent facility a full-scale demonstration of the soluble sulfide
precipitation process for the pretreatment of a metal finishing wasEewater was
performed. Soluble sulfide precipitation was selected because the lower
solubility of metal sulfides was expected to result in better metal removal
efficiency Chan conventional hydroxide precipitation.
Three segregated wastes were treated 'separately. Cyanide-containing
wastes were treated in a two-stage alkaline chlorination process for complete
cyanide oxidation. Chromium-containing wastes were acidified to pH 2.5 and
treated with sodium metabisulfite to reduce hexavalent chromium to the less
soluble trivalent form. Following separate treatment, these wastes were
combined with the acid/alkali and metals contaminating wastes for treatment by
soluble sulfide precipitation.
10-61
-------
This treatment system consists of pH adjustment with caustic soda,
addition of ferrous eulfate and anionic polymer as coagulants, addition of
sodium sulfide to precipitate metals, flocculation, parallel plate
clarification, gravity sand filtration, and peroxide destruction of residual
sulfide. Sludge processing consists of gravity thickening and dewatering in a
plate and frame filter press. Table 10.1.24 compares limitations for a
6 month period.
Insoluble sulfide precipitation—In 1980,'three plants using the Sulfex
process to remove heavy metals from wastewater discharge were investigated to
assess system performance. Two of the plants (Plants A and B)- use the Sulfex
process singularly, while the third (Plant C) uses the process as a polishing
52 '
step after hydroxide precipitation and clarification.
Plant A uses both electroless and electrolytic plating processes to plate
plastic components. The heavy metals in the wastewater (copper, nickel, and
chromium) are complexed/chelated with a variety of proprietary agents.
Plant 1 manufactures carburetors for the automotive industry. Wastewater from
the metal finishing portion of the process contains' chromium, zinc, iron,
phosphates, organic chelating agents, and assorted plating chemicals. Plant C
treats wastewater from a barrel-dip, zinc-phosphating line.
Table 10.1.25 presents the chemical consumption and sludge generation
rates for Plants A, B, and C. While Plants B and C were successful in
lowering metallic contaminants to effluent discharge requirements, Plant A was
unable to treat both bexavalent and total chromium. The poor performance in
chromium removal was primarily due to an increase in the level of hexavalent
chromium in the mixer/clarifier without a commensurate increase in the FeS
feed to compensate for the increased demand. Consequently, the level of
unreacted FeS in the sludge blanket was gradually depleted and eventually,
insufficient FeS was present in the blanket to achieve the normal high level
of removal. The FeS stored in the sludge blanket prior to reagent depletion
was able to maintain the high removal efficiency.
10-62
-------
TABLE 10.1.24, TOBYHANNA ARMY DEPOT WASTE AND TREATED
EFFLUENT ANALYSIS (mg/L)
Parameter
Cadmium
Chromium
Copper
Lead
Nickel
Silver
Zinc
Cyanide
Aluminum
Tin
Suspended solids
Oil and grease
Waste
average
1
1
2
0
1
3
1
6
0
.34
.14
.35
.43
.61
-
.4
.08
.67
.003
-
-
Effluent
Average
0
0
0
0
0
0
0
0
4
0
18
12
.09
,31
.07'
.19
.08
.01
.37
.04 -
.3
.01
.8
.8
Maximum
0
1
0
0
0
0
2
0
18
1
.25
.15
.47
.4
.35
.02
,69a
.12
.0
Daily maximum
Design
1.2
7
4.5
0.6
4.1
-
4.2
0.8
1
2.5
152s
22
-
Permit
0.69
2.77
3'. 38
0.69
3.98
0.43
2.61
1.2
-
-
60
52
Daily average
Design
0.5
2.5
1.8
0.3
1.8
-
1.8
0.23
0.5
1
_
-
Permit
0
1
2
0
2
0
1
0
31
26
.26
.71
.07
.43
.38
.24
.48
.65
-
-
aExceeds permit limit.
- Indicates data not available or no standard specified.
Source: Reference 56.
10-63
-------
TABLE 10.1.25.' WASTEWATER TREATMENT PROCESS CHARACTERISTICS
- FOR FOUNTS A, B, AND CB
Characteristic
Wa s tewat er
Average flow rate (gal /rain)
pH:
Feed
Effluent
Average feed concentration (ppm):
Nickel
Copper
Hexavalent chromium
Total chromium
Zinc
Iron
Phosphorus
Average effluent concentration (ppm):
Nickel
Copper
Hexavalent chromium
Total chromium
2 171 C
Iron
Phosphorous
Treatment ch&micals - •
Lime-b
Ib/h
Calcium chloride (for phosphate removal):
Ib/h
Cationic polymer.
Ib/h
Anionic polvraer:^
Ib/h
Ferrous sulfide:
Ib/h
Sludge generation factors
Dry goods generation:
Ib/h
First stage
Second stage
lb/1,000 gal wastewater
Underflow volume (gal/h at 0.75% solids)
Filter cake volume (gal/h at 30% solids)
Plant A
39
2.0 - 4,0
9.0 - 10.0
31
28
76
88
WA
NA
NA
0.54
0.03
0.10
0.20
_
-
-
8.8
NA
0.1
NA
12/SC
' 23.7
NA
NA
10.1
380
7.9
Value
Plant B
21
4.5 - 6.0
8.5 - 9.5
NA
•NA
.27
39
ia
1.4
NA
-
-
0.005
0.13
0.02
0.05
-
2.0
NA
0.17
NA
4/5d
7.2
NA
NA
5.7
114
2.4
Plant C
16
2.5 - 3.0
7.5 - 8.5
NA
NA
0.07
8 •
24
127
289
-
-
0.02
0.10
0.12
0.60
0.3
8.1
17.0
0.02
0-01
0.30b
16.4
16
0.4
17
262
5.3
aAll three plants use an ISP process to remove metals from wastewater,
bat Plant C uses ISP as a polishing system.
Observed rates.
Based on three tiroes the stoichionietric requirement.
Based or. four times the staichiometric requirement.
Source; Reference 52.
-------
Calcium sulfide—Many- of the problems associated with soluble and
insoluble sulfide precipitation (i.e., excess reagent requirements and H. S
evolution) can be minimiEed with calcium sulfide as the sulfide
source. ' Solid CaS can be added to the wastewater as a slurry. The
addition of'CaS as a slurry produces easily settleable precipitates: calcium
sulfide particles act as nuclei for production of metal sulfide particles, and,
the dissolved calcium ion functions as a coagulant. Since calcium, which is
added as CaS, is mostly dissolved in the wastewater after reaction, the
increase in the sludge volume is minimal. For the same reason, unlike FeS,
the CaS requirement is near stoichiometric.
Calcium sulfide is stable only in dry solid form. In aqueous solution,
it reacts with water to produce Ca(HS}. and Ca{OH)_.
2CaS + 2H20 - Ca(HS)2 * Ca(OH)2
or: S~ - HS~ + OH~
The1 main reactions involved in the precipitation of metal sulfides after
adding the CaS solutions are:
M4"1" + HS~ = MS + H4"
H+ + OH- =
M++ + S~ = MS
Research conducted by the General Electric Company, Scbenectedy,
New York, investigated the effectiveness of calcium sulfide as a precipitation
agent. The investigation involved batch treatment of the wastewater by
suLfide precipitation with the addition of lime until the pH reached 7 and
next, 0.1 M CaS solution to a desired pH value, normally 9.0. The
precipitates were flocculated with 2 mg/L Nalco 7763 polyelectrolyte.
Vigorous mixing (600 rpm) of the solution for 2 minutes followed by moderate
mixing (30 rptn) for 1 minute was sufficient for effective flocculation. The
floes were settled for 30 minutes before sampling supernatant liquid for
analysis. The solution was further filtered with 0.2 um Acropor filter to
remove any suspended solids.
10-65
-------
Both actual and simulated metal finishing wastewaters were treated using
hydroxide and ealciuta sulfide (a mixture of Ca(OH)„ and Ca(HS)3)
solutions. In addition, both treatment methods were evaluated on wastewaters
containing chelating agents to test their effects. The calcium sulfide
preparation system produced the CaS solution from H,S and Ca(OH)_. The
ratio of Ca and S was controlled by measuring the pH of the solution. Vessels
for pH adjustment, sulfide precipitation, and flocculation were included in
the system!
the results of these experiments, showed that sulfide precipitation is
very effective for the removal of heavy metals such as Cd, Cu, Pb, Ag, and
Zn. The method works in the presence of chelating agents and removes metals
to extremely low concentrations. The calcium sulfide slurries, prepared by
reacting lime with hydrogen sulfide or sodium hydrosulfide, are effective
sulfide sources. The addition of calcium sulfide can be controlled simply, in
most cases, by measuring the pH. The processes employing such techniques have
been demonstrated in bench-scale experiments using uastewater samples. The
two-stage process may be employed when the wastewater contains a large amount
of iron and nontoxic suspended solids.
10.1.3.3 Process Costs-
Table 10.1.26 presents the cost data developed for a 'continuous soluble
sulfide precipitation system. The purchased equipment and installation costs
are based on the treatment process shown in Figure 10.1.6 and the assumptions
made in Section 10.1.2. An additional aeration vessel consisting of a
reinforced concrete reactor, 4-b acid resistant spargers ($82/sparger), and
30 feet of 6 inch pipe ($2.40/ft) has been included in the treatment train to
reduce the'Quantity of H«S fumes evolved.
Operating labor requirements have been increased from 4 to 6 hours per
shift due to the greater need for process control (to prevent excess sulfide
O C *) CO
dosing) associated with this process. ' * Maintenance, overhead,
utilities, taxes and insurance have all remained constant; however, reagent
chemical costs have increased dramatically. Since sodium sulfide flske costs
$410/ton vs. $40/ton for hydrated lime, an equivalent influents metsls
concentration would result in.a greater tnan 10 fold increase in reagent cost
when comparing the two systems.
-------
TABLE 10.1.26. CONTINUOUS SOLUBLE SULFIDE PRECIPITATION COSTS1
Purchased Equipment and Installation (PE&I)
Equalization Tank
Precipitation Reactor
Flocculator/Clarifier
Aeration Vessel
Sludge Holding Tank(s)
Filter Press
Total Capital Investment (360% PE&I)
Operating Costs ($)
Operating Labor ($20/hr)
Maintenance (62 TCI)
General Plant Overhead (5,8% TCI)
Utilities (25; TCI)
Taxes and Insurance ( ll TCI)
Chemical Costs Na2S ($410/ton)
FeS04 ($145/ton)
Sludge Transportation (S0.25/ ton-mile)
Sludge Disposal ($200/ton)
Annualized Capital (CFR = 0.177)
Total Cost/Yr
Cost/1,000 gallon
1,000
(*)
1 7 , 000
24,000
18,000
17,400
3,000
10,000
89,400
321,800
108,000
19,300
18,700
6,400
3,200
4,050
350
300
13,800
57,000
231,100
32
Flow Rate
10,000
29,000
40,000
50,000
29,500
6,000
25,000
179,500
646,200
108,000
38,800
37,500
12,900
6,500
40,500
3,600
2,600
138,100
114,400
502,900
7.0
(gph)
100,000
50,000
69,000
140,000
50,500
48,000
100,000
457,500
1,647,000
108,000
98,800
95,500
32,900
16,500
404,800
35,700
25,600
1,380,000
291,500
2,489,300
3.5
4987 Dollars,
10-67
-------
Sludge disposal costs, are roughly equivalent for sulfide and hydroxide
precipitation, while annualized capital costs are slightly greater due to the
requirement of an additional reaction vessel. Overall costs for sulfide
precipitation, based on the assumptions presented in this section, are
approximately 40 percent greater than those associated with hydrated lime.
However, sulfide precipitation economics become more favorable when compared
to the more expensive hydroxide reagents such as sodium hydroxide C$175/ton •
for 30 percent solution) or magnesium hydroxide C$0.78/lb). If lower
dissolved heavy metal concentrations were desired, the most economically
efficient use of sulfide precipitation would be as a final polishing step to,
or in conjunction with (co-precipitation), hydroxide precipitation.
10.1.3.4 Status of Technology—
Sulfide precipitation has been demonstrated to be an effective
alternative to hydroxide precipitation for removing various heavy metals from
o CO C fi... £. 1
industrial wastewaters, * ' The major advantage of the sulfide
precipitation process.is that because of the extremely low solubility of metal
sulfides, very high metal removal efficiencies can be achieved. The major
limitations of the sulfide precipitation process are the evolution of toxic
hydrogen sulfide fumes and the discharge of treated wastewaters containing
52 57
residual levels of sulfide ' (see Table 10.1.27 for summary of advantages
and disadvantages of sulfide precipitation).
The use of ferrous sulfide (insoluble sulfide process) as a source of
sulfide reduces or virtually eliminates the problem of hydrogen sulfide
evolution. The use of ferrous sulfide, however, requires reagent
consumption considerably higher than stoichiometric and significantly higher
sludge generation, than either the hydroxide or soluble sulfide treatment
processes.
The use of calcium sulfide as a source of sulfides reduces the problems
(H.S evolution and excess reagent requirements) associated with the previous
57
two technologies. However, as with ferrous sulfide, calcium sulfide
precipitation results in high solids generation. These solids must be removed
in a subsequent treatment step, such as sedimentation or filtration. Sulfide
sludges are less subject to leaching than hydroxide sludges, hany landfills
now require post-treatment of the residuals through.such technologies as
stabilization or encapsulation be performed prior to land disposal,
10-68 • :
-------
TABLE 10.1.27. ADVANTAGES AND DISADVANTAGES OF SULFIDE PRECIPITATION
Advantages
- The sulfide process has the ability to remove chromates and dichromates
without preliminary reduction of the chromium to the trivalent state.
The high reactivity of sulfidea with heavy metal ions and* the
insolubility of metal sulfides over a broad pH range are attractive
features compared with the hydroxide precipitation process.
Sulfide precipitation, unlike hydroxide precipitation, is relatively
insensitive to the presence of most chelating agents arid eliminates the
need to treat these wastes separately.
Disadvantages
Sulfide reagent will produce hydrogen sulfide fumes when it comes in
contact uith acidic wastes. This can be prevented by maintaining the pH
of the solution between 8 and 9.5 and may require ventilation of the
treatment tanks.
- As with hydroxide precipitation, excess sulfide ion roust be present to
drive the precipitation reaction to completion. Since the sulfide ion
itself is toxic, sulfide addition must be carefully controlled to
maximize heavy metals precipitation with a minimum of excess sulfide to.
avoid the necessity of pos t- treatment . Where excess sulfide is present,
aeration of the effluent stream would be necessary to oxidize residual
sulfide to the less harmful sodium sulfate (
The cost of sulfide precipitants is high in comparison with hydroxide
precipitant, and disposal of metallic sulfide sludges may pose problems.
Source: References 2, 52, and 57-61.
1U-69
-------
10.1.4 Carbonate Precipitation
10.1.4.1 Process Description—
Carbonate precipitation may be used to remove metals either by direct
precipitation using a carbonate reagent such as calcium carbonate (limestone)
or sodium carbonate or by converting hydroxides into carbonates using carbon
dioxide. The solubility of most metal carbonates is intermediate between
hydroxide and sulfide solubilities; in addition, carbonates form easily
filtered precipitates.
Calcium Carbotvate--Limestone is available in either high calcium
(CaCO.) or dolomitic (CaCO- MgCO-) form. Both types of limestone
are available as either a powder or crushed stone. Crushed stone diameters
are typically 0.074 mm (200 mesh) or less since both the reactivity and
completeness of the reaction increase proportionately to the available surface
area. High calcium is most commonly used because of its greater reaction
rate and its more widespread availability. Dolomitic limestone reactivity
will increase if it is finely ground, and sludge production will be minimal
due to the formation of soluble magnesium sulfate. -However, its reactivity is
generally too slow even with grinding, and hence not suitable for most
30
applications.
The inherent problem with calcium carbonate precipitation is that it is
only effective for reducing metallic species such as trivaient chromium and
67
iron in its operational pH range (5.0 to 7.0). In addition^ the
inhibition of the stone particles in the presence of high quantities of
sulfate and/or metallic ions make it less attractive than other reagents.
Limestone is a solid-based reagent that liberates CaO for precipitation
through surface dissolution. The inhibition of the particle surface through
calcium sulfate precipitation increases retention times, reagent purchases,
equipment sizing, and lowers waste throughput. ' ' Improved reaction
kinetics can be achieved by increasing the available solid surface area
through greater limestone loading. However, both reagent purchase and
sludge disposal costs will increase proportionately with the excess limestone
applied.
10-70
-------
The primary advantage-of limestone neutralization is that limestone is a
low cost and widely available reagenc. However, limestone is limited in its
ability to neutralize over pH 6.0 or treat acid concentrations greater than
5,000 mg/L. ' There have been attempts to use limestone in combination
with lime in a. dilute, dual alkali mode. The limestone is used as a
pretreattnent to raise the pH to approximately 3,0 or 6.0 with lime completing
the process of precipitation. The limes tone/lime process is usually more
complicated than a simple lime slurry process, resulting in higher projected
costs and limited application. However, in high volume applications the
savings in reagent (when used in pebble form) may offset any increase in
capital expenditures.
Sodium Carbonate-~Sodiuro carbonate (Na-G0») is a highly reactive
soluble alkali that -is marketed most often as an anhydrous powder. Wet
crystal bulk storage typically facilitates solution feeding. In dry form it
is also easily fed from hoppers. Positive provision for dissolution is
desirable for dry feed applications. Suitable materials for handling the
compound or its solutions include plastic, iron, rubber, and steel. Shipment
is made in bags, barrels, or in bulk with transfer usually performed by
pneumatic conveyor.
In the chemical trade the terms "ash," "soda ash," "soda," and "calcined
soda" are used for the anhydrous salt, although soda ash is the most common
name in English-speaking countries. Sodium carbonate is moderately soluble in
cold water and soluble to approximately 30 percent of solution weight in hot
water the solution is strongly alkaline. (Melting point, 851°C; heat
capacity at 25°C, 1043.01 J/(kg-K) [249.3 cal/(kg-K)]; heat of fusion,
315.9 kJ/kg (75.5 kcal/kg); density at 20°C, 2533 kg/ra . Bulk densities of
various commercial grades range from 576 to 1072 kg/m (36-67 Ib/ft 5.
Ordinary light soda ash produced by calcining crude bicarbonate is
satisfactory for many uses. Dense soda ash is most often manufactured by
hydration of light ash to produce larger sodium carbonate monohydrate crystals
followed by dehydration. Hydration may be accomplished by either feeding
light ash and water to mixers or blenders or by adding light ash to a saturated
solution of soda ash containing a slurry of monohydrate crystals. The
monohydrate crystals are fed to a continuous dryer. The dehydrated product
from the dryers needs only screening before packing and shipping. Most dense
ash is shipped in bulk to large industries.
10-71
-------
Sodium carbonate is an alternative to sodium alkali.for acidic-metals
wastestreams lacking buffering capacity Such as deionized acid-bath
rinsewaters. The use of sodium carbonate (a weak base) with strong acids,
such as sulfuric, will impart a buffer to the .wastewater stream, thereby
facilitating pH control and precipitation within the neutral range. These
buffering reagents will produce a smaller change in pH per unit addition than
comparable unbuffered, strong bases such as high calcium lime or caustic
37 . '
soda. This phenomena can be seen in Figure 10.1.13, which illustrates the
neutralization of a 1 percent sulfuric acid solution with caustic soda and-
soda ash. A small incremental addition of caustic soda caused the pH to
change from 2 to 11 standard units- Alternatively, approximately three times
the quantity of soda ash resulted in a modest pH change from 6 to 9 units.
Due to its carbonate-based reaction mechanism, the neutralization/
precipitation metals-containing acidic rinsewaters with soda ash (as with
limestone) proceeds at a much slower pace than comparable hydroxide-based
reagent systems such as lime or caustic soda. Accordingly, continuous flow
reactors must be sized to provide a minimum of 45 minutes hydraulic retention
in each stage. In addition, soda ash is commercially available only in a
dry form. Consequently, onsite batch mixing and solution preparation
facilities, siteilar to those of hydrated lime, are mandatory when using this
chemical as a neutralizing agent. The solubility of soda ash also limits its
use since a chemical solution feed strength of only 20 percent by weight can
be maintained at ambient temperatures without salt recrystallization.
Continuous mixing of the prepared solution is recommended to maintain
homogeneity -
An advantage of soda ash is lower sludge generation since sodium-based
end products are more soluble than calcium-based products. However,
sodium-based sludges do not filter as readily or to as high solids content as
calciuEr-based sludges. In addition, the clarified liquid effluent may not be
as low in metals content or total dissolved solids as insoluble end product
systems such as lime. "All these factors must be carefully weighed before
selecting sodium carbonate or any other alkeline reagent as a precipitating
agent.
10-72
-------
o
i
12.0
II.0
10.0
9.0
0.0
7.0
6.0
5 O
1.0
3.0
Z. O
I. 0
NoOII 50%
1.0
1
2.0
3.0 4.0 5.0 6.0
GI1AMS OF HEAGrNT ADDED
7.0
0.0
9.0
No CO
Z 3
I
10.0
Figure 10.1.13, Titration curve for the neutralization of a 1% I^SO, solution
with sodium hydroxide and sodium carbonate.
Source: Reference 37.
-------
Carbon Dioxide--Carbon dioxide is a relatively old but, as of yet,
undeveloped technology for treating metals-containing wasteetreatns.
Typically, carbonic acid is generated directly in the
neutralization/precipitation' chamber by injecting carbon dioxide into the
wastewater solution. Upon hydration, the carbon dioxide will form carbonic
acid and react with available hydroxides to form less soluble carbonates.
Carbonic
acid
Ni(OH)2
Nickel
hydroxide
NiCOj +
Nickel
carbonate
2H20
(1)
Ca(OH)2
CaCOj
2H20
(2)
Carbonic
acid
Hydrated
lime
Calcium
carbonate
pH 9.4
{ saturated }
Compressed (liquid] carbon dioxide is stored and transported at ambient
temperatures in cylinders containing up to 22.7 kilograms. Larger quantities
are stored in refrigerated, insulated tanks maintained at -18°C and
20 atmospheres, Transportation is by insulated tank truck and rail car.
The standard method of applying compressed carbon dioxide for
precipitation is to vaporize carbon dioxide in a heat exchanger or across a
flash valve. The pressurized gas is forced through porous diffuser tubes
placed along the bottom of a batch treatment tank. Carbon dioxide gaa is
released from the diffusers as fine bubbles (15 microns) which are
preferentially absorbed by the surrounding wastewater. This type of treatment
requires a slow-moving effluent stream with a treatment tank of sufficient
depth to ensure that the carbon dioxide is fully absorbed before reaching the
surface. Since hydration of carbon dioxide forms carbonic acid, it is
recommended that the diffuser assembly be constructed of a corrosion-proof
material.
10-74
-------
The primary advantages of compressed carbon dioxide are minimal capital
requirements, uncomplicated piping, and the inability to over-acidify the
wastewater. Its primary disadvantage* are a low dissolved oxygen content
(4,5 percent) at the point of injection, and a high reagent cogt on a
neutralization equivalent basis (approximately $200 to 3>300/ton. However, for
large volume users of 200 tons or more per year, che unit cost per ton of
CO f. Q
compressed carbon dioxide drops to $90 to tlOQ/ton. '
10.1.4.2 Process Performance—
Carbonate precipitation technology is sometimes preferred ovgr hydroxide
precipitation because in some instances it provides superior precipitation
properties; i.e., with cadmium it produces cadmium carbonate which is
preferred cadmium hydroxide for recovery purposest Also, nickel and lead
precipitation with carbonate gives lower final levels than precipitation with
hydroxide.
Treatment of cadmium with sodium carbonate (soda ash) will give good
levels of removal at a slightly lower pH than'hydroxide, typically in the
range of 9.5-10. Due to the value of cadmium, it is often desirable to send
the precipitated sludge to & reprocessor for recovery of the cadmium, or to
reuse it. Whether the cadmium is in the hydroxide or carbonate form may be
important to the reprocessor plant operator.
Bench—scale tests conducted by Nassau Smelting and Refining Co.
studied lead precipitation by caustic, lime and caustic soda/soda ash. It was
found that both lime and caustic soda/soda ash gave good results. The optimum
pH was 9.0 to 9.5. Influent lead was 5 mg/L and final lead was 0.01 to
0.04 mg/L.
Figure 10.1.14 shows solubility levels of lead with different alkali
agents. As can be seen, the soda ash/caustic soda systems produced slightly
better results than the straight-line system. Separaa A? 30 was used as 8
coagulant aid.
Investigators at the Illinois Institute of Technology performed a series
72
of solubility experiments. In this series, precipitation experiments were
performed over 24-hour periods at constant pH and C (total carbonate). Two
levels or carbonate were evaluated from each metal: a low background
1U-7J
-------
o>
E
•o
o
a>
0.30
0.20
0.10
0
Soda ash and
Caustic so.do
8.0 • 8.5
9.0 9.5
PH
1.0.0
10.5
Figure 10,1.14. Lead solubility in three alkalies*
Source: Reference 71.
10-76
-------
—3 ft
carbonate level of less than 10 ' H( 2 nsg/L inorganic carbon) and a
— 3 2
carbonate level of approximately 10 " MC7.6 mg/L inorganic carbon).
Values of pH ranging from pH 6 to pH 13 were tested. In addition, hydroxide
experiments were performed under the same conditions for cadmium, copper,
lead, and zinc.
Cadmium Solubility—Minimum cadmium solubility of 0.08'mg/L was obtained
at pH 10-10.5. In the pH range of 6,5 to 8.5, the carbonate system yielded a
soluble cadmium concentration range of 81.0 mg/L to 0.66 rag/L. The hydroxide
system, over the same pH range of 6.5 to 8.5, yielded a much greater soluble
4
cadmium concentration range of 8 x 10 tng/L to 129 mg/L. Lower soluble
cadmium concentrations are observed in the test system with carbonate
present. At higher pH values of 9-10, the soluble metal concentrations are
comparable for both systems. This suggests that both systems are controlled
by hydroxide solubility at pH 9-10 rather than pH near 10. At pH above 10.0
there is a significant difference in soluble cadmium concentrations. Soluble
cadtcium' concentrations for the carbonate system were much lower. This appears
to be due to a slight increase in carbonate concentration at pH 9.5, which
would decrease soluble cadmium concentrations.
Copper Solubility—The minimum soluble copper concentration attained was
0.005 mg/L at pH 8.9 to 9.3. Over the pH range 6.7 to 7,9, soluble copper
concentrations were reduced from 3.5 mg/L to 0,016 rng/L.
Minimum solubility of 0.015 tng/L to 0.018 mg/L was obtained in the
carbonate test system, in the pH range of 8.6-10.4. From pH 7.5 to 9.5 the
soluble copper concentration in the hydroxide test system ranged from 0.021 to
0.005 mg/L. With an increase in carbonate concentrations to C_ =
-3.2 •
10 * M, the soluble copper concentration in the carbonate system ranged
from 0.061 to 0.016 mg/L.
Lead Solubility—It is apparent that the carbonate induces lower soluble
lead concentrations than occur in the hydroxide system, at pH below 8. At
pH 7.0 to 7.5, the carbonate system yielded a minimum soluble lead
concentration of G.G25 mg/L while che hydroxide system produced a lead
concentration of 0.131 mg/L. Above pH 8, carbonate functions as a ligand co
increase lead solubility.
10-77
-------
Zinc Solubility—Solubility patterns for both systems are similar, with
data points for the carbonate system generally below those of the hydroxide
system, but higher than the theoretical carbonate solubility curve. There is
some evidence that the zinc carbonate precipitation system approaches
equilibrium extremely slowly, perhaps requiring more than 10 days to near
equilibrium solubility. This has been postulated to result from the more
rapid kinetics of zinc hydroxide precipitate formation, even in a system
thermodynamically stable for zinc carbonate. The subsequent kinetics of zinc
solubility then would be limited by the slow transformation of solid phase
zinc hydroxide to zinc carbonate,
10.1.4.3 Process Costs—
Table 10.1.28 details the cost data developed for a continuous sodium
carbonate precipitation system. The purchased equipment and installation
costs are equivalent to those of the hydrated lime precipitation system except
that a retention time of 1 hour instead of 30 minutes has been used to size
the precipitation reactor (due to the slower reactivity of sodium carbonate).
In addition, chemical reagent costs and useages are significantly higher for
sodium car.bonate when compared to hydrated lime. For example, approximately
2.9 Ibs of sodium carbonate (at $120/ton) are required to precipitate 1 Ib of
heavy metal, while only 2.2 Ibs of hydrated lime (at $40/ton) are required
per Ib of metal. However, due to the higher solubility of the sodium cation
in the sodium carbonate complex sludge generation is only 7 percent higher Con
a dry weight basis).
Overall costs for sodium carbonate, based on the cost data presented in
this section, are only 1 to 18 percent greater than those presented for
hydrated lime. The viability of this technology as an alternative to'either
hydroxide or sulfide precipitation is enhanced due to the lower pH
requirements (usually 8-9) for carbonate precipitation. The lower pH
requirement will result in lower alkali demand for neutralization and
consequently less sludge generation. Therefore in any consideration of
alternate precipitation technologies, influent pH should also be examined.
10-78
-------
TABLE 10.1,28. CONTINUOUS SODIUM CARBONATE PRECIPITATION COSTSa
Purchased Equipment and Installation (PE&I)
Equalization Tank
Precipitation Reactor
Flocculator/Clarif ier
Sludge Holding Tank(s)
Filter Press
Total Capital Investment (360% PE&I)
Operating Costs (S)
Operating Labor ($2Q/hr)
Maintenance (61 TCI)
General Plant Overhead (5.8% TCI)
Utilities (2% TCI)
Taxes and Insurance (1% TCI)
Chemical Costs (S120/ton)
Sludge Transportation ($0.25 /ton-mile)
Sludge Disposal (S200/ton)
Annual ize'd Capital (CFR = 0.177)
Total Cosc/Yr
Cost/1,000 gallon
1,000
(S)
17,000
24,000
18,000
3,000
10,000
72,000
259,000
72,000
15,500
15,000
5,200
2,600
2,100
200
12,800
45,800
171,200
24
Flow Rate
10,000
29,000
60,000
50,000
6,000
25,000
170,000
612,000
72,000
36,700
35,500
12,200
6,100
20,600
2,100
128,400
108,300
421,900
6
Cgph)
100,000
50,000
150,000
140,000
48,000
100,000
488,000
1,756,800
72,000
105,400
101,900
35,100
17,600
206,400
21,400
1,284,000
311,000
2,154,800
3
al987 Dollars.
10-79
-------
10.1.4,4 Status of Technology—
Carbonate precipitation has been demonstrated to be a viable alternative
to either hydroxide or sulfide precipitation for removing various heavy metals
from industrial wastewaters. The solubility of rnoet metal carbonates is
intermediate between hydroxide and sulfide solubilities. In addition, the
reagent cost is also intermediate. The main advantages of carbonate
technology are buffering capability, superior handling characteristics
(i.e., little dust, good flow, and no arching in the feeder), and widespread
availability. Main disadvantages are slow reaction time (typically a minimum
of 45 minutes retention) and low solubility (20 percent by weight). Since
carbonates are not particularly corrosive and soda ash generates less sludge
than comparable calcium-based technologies, environmental impacts are few.
See Table LO.1.29 for summary of advantages and disadvantages of carbonate
precipitation.
-------
TABLE 10,1,29. ADVANTAGES AND DISADVANTAGES OF CARBONATE PRECIPITATION
Advantages
- Carbonate reagents have a relative ease of handling and can be
obtained in bulk by railcar or truck or in 100 Ib bags
Calcium carbonate forms easily filtered precipitates
Sodium carbonate imparts buffering capacity and generates less sludge
Disadvantages
- Retention times are longer due to slower reacting carbonate-based
chemistry
- Carbonates do not mix easily into solution ana have the potential for
evolving carbon dioxide which, without aerationj will slow reaction
times further
Calcium carbonate particles have the potential to become deactivated
if calcium sulfate precipitates on particle surface
Sodium carbonate sludges do not filter as readily or to as high
solids content as calcium-based sludges
Calcium carbonate is only able to achieve an operational pH-range
of 5-7
Source: References 62, 66, 70, and 72.
10-81
-------
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Filtration, Proceedings of the 36th. Industrial Waste Conference, Purdue
University. 1981,
59. Resta, J.E. et al. Soluble-Sulfide Precipitation Treatment of Metal
Finishing Wastewater. Proceedings of the 16th. Mid-Atlantic Industrial
Waste Conference, 1984,
60. Talbot, R,S. Co-Precipitation of Heavy Metals with Soluble Sulfides
Using Statistics for Process Control. Proceedings of the 16th.
Mid-Atlantic Industrial Waste Conference. 1984.
61. Peters, R.E., et al. The Effect of Chelating Agents on the Removal of
Heavy Metals by Sulfide Precipitation. Proceedings of the 16th.
Mid-Atlantic Industrial Waste Conference. 1984.
10-85
-------
62. Gloves, H.G. The Control of Acid Mine Drainage Pollution by Biochemical
Oxidation and Linestone Neutralization Treatment, 22nd. Industrial Waste
Conference, Purdue University. 1967.
63. -Genm, H.U. Neutralization of Acid Hmetewaters with Upflow Expanded
Limestone Bed. Sewage Works Journal 16:104-120. 1944,
64. Tuily, T.J, Waste Acid Neutralization. Sewage and Industrial Was'tes
30:1385. 1985.
65. Volpicelli, G., et al. Development of a Process for Neutralizing Acid
Wastewaters by Powdered Limestone. Environmental Technology Letters,
Vol. 3, pp. 97-102. 1982.
66. Kirk-Othmer Encyclopedia of Chemical Technology. Vol. 4, 3rd. Edition,
John Wiley & Sons, New York, NY. pp. 725-741. 1981.
67. Griffith, M.J. et al. Carbon Dioxide Neutralization of an Alkaline
Effluent. Industrial Waste. March 1980.
68. Ponzevik, D. Liquid Air Products. Telephone conversation with Stephen
Palmer, Alliance Technologies Corporation. September 6, 1986.
69. Berbick, D., Cardox Corporation. Telephone conversation with Stephen
Palmer, Alliance Technologies Corporation. September 25, 1986,
70. Lanoutte, K..H. Heavy Metals Removal. Chemical Engineering. October 17,
1977.
71. Day, R.V., Lee, E.T., and E.S. Hochuli, Bell System's hetals Recovery
Plant. Industrial Waste. July-August 1974.
72. Patterson, J.W. Effect of Carbonate Ion on Precipitation Treatment of
Cadmium, Copper, Lead, and 2ine. 36th. Industrial Waste Conference.
Purdue University. 1981.
10-86
-------
10.2 COAGULATION AKD FLDCCULATION
Chemical coagulation and flocculation are two terms often used
interchangeably to describe a process whereby a chemical addition is made to
enhance sedimentation (removal of solid particles from suspension by
gravitational settling) operation. Coagulation and floeculation are often
used to remove the insoluble and colloidal heavy metal complexes formed by
precipitation. In this text, chemical coagulation is defined as particle
agglomeration brought about by the reduction of electrostatic particle surface
charges. Flocculation is a time-dependent physical process of aggregation o£
fine particles into solids large enough to be separated.
The coagulation process involves the destabilization of the suspension by
neutralizing or decreasing the repulsive forces on the particles, so that the
123
particles will approach each other and agglomerate, ' '
The charge on organic, inorganic and biocolloids is typically negative
when suspended in water. The, negative charge attracts positive ions due to
electrostatic forces which are distributed as shown in Figure 10.2.1. The
inner layer (termed the stern layer) contains adsorbed ions and is typically
about the thickness of a hydrated ion. The second diffuse layer contains a
shear plane within which ions move with the particle. Outside the shear plane
ions move independently of the particle as dictated by fluid and thermal
2
motions. The electrical potential difference between the shear plane and
the bulk solution is termed the Zeta potential. Zeta potential is a
measurable quantity and can be used qualitatively to predict the potential for
Lor
2
I 2
coagulation. ' As the Zeta potential approaches zero, coagulation should
increase.
However, the overall success of the coagulation/flocculation process is
ultimately dependent upon the flocculating- and settling characteristics of the
particles. The rate at which coagulated particles coalesce is primarily
related to the frequency of the collisions between the particles. Collisions
occur as a result of heavier faster particles overtaking lighter slower
particles. The collision frequency is proportional to the concentration of
particles and the difference in settling velocities. Since the total
number of collisions increases with time, the degree of flocculstion also
generally increases with residence time in the reaction chamber.
-------
STERN
! iYFB *• *•
LAYER
SHE
•+
:AR
PLANE
f
4^.
„, DIFFUSED ^.
LAYER
Figure 10.2.1. Double layer charge distribution.
Source; Reference '4.
1U-88
-------
'The rate of flocculation cannot be predicted from collision frequency
alone. The coalescence of particles depends upon raany factors, such as the
nature of the surface, the presence of charges, shape, and density. At
present, there is no adequate theoretical model to predict the rate o£
flocculation in a suspension.
As with precipitation, most coagulation/flocculation processes operate
under the same fundamental chemical principles and utilize similar types of
equipment and process configurations. For example,
coagulation/flocculation processes typically entails the following three steps
1. Addition of the coagulating/flocculating agent'to the treated
wastewater,
2. Rapid mixing to disperse the coagulating agent throughout the liquid,
3, Slow and gentle mixing to allow for contact between small particles
and agglomeration into larger particles.
Coagulant Addition—Probably the most important parameters to be defined
in the design of a coagulation/flocculation system are the type of and dosage
of the coagulant, the pH and the mixing-characteristics. The most common
method of determining these parameters is through a jar test (described in
references I and 6). The jar test is a laboratory scale test where the
wastewater to be treated can be subjected to variable conditions of pH,
coagulant type, dosage and mixing, flocculating and settling times. The
effect of various coagulant aids can also be investigated in this test. The
4
results of the test provide the following'types of information:
* Optimum pH value for efficient coagulation with different coagulants,
* Optimum coagulant dosages for effective flocculation.,
• Effectiveness of coagulant aids.
• Most effective order of chemical addition.
• Correct mixing times.
* Flow settling characteristics.
10-89
-------
Quantity of sludge requiring disposal.
Quality of clarified water to "be expected from a particular
treatment.
From these data the chemical requirements and unit sizes for coagulation,
flocculation and sedimentation can be determined.
Rapid Mix—Rapid (flash) mixing residence times have been reported as
7 7 1 $
30 seconds to 5 minutes, -2 to 5 minutes, and 10 to 30 seconds. '
Mixing characteristics are determined by the velocity gradient in the mixer (a
measure of the shear intensity). Insufficient mixing will affect the
performance of subsequent steps and overmixing can break up previously formed
£loc or the incoming wastewater solids. Static -mixers can also be employed
although the mixing characteristics are a function of the flow which cannot be
controlled. The velocity gradient, G, is usually chosen at about
300 ft/sec/ft.
Slow and Gentle Mixing—The slow and gentle mixing stage is usually
carried out in a flocculator/clarifier. Clarification is defined as a
quiescent flow condition with a hydraulic flow velocity sufficiently low to
allow particles with some minimum settling velocitv to separate from the waste
overflow. The solids collect in the base of the chamber where a rake or
suction device is used to remove the collected solids,.
Sometimes, sludge recycle is practiced to gently mix the treated
wastewater with a slurry of previously settled ^sludge solids. The recycle
solids present a dense concentration of nucleation sites to promote particle
growth,
Flocculator/clari£iers used for the removal of heavy metal contaminants
, . 9-14
come in three basic types:
» Basic settling chambers, where the feed is distributed at one end
and overflows at the other. This type of unit often requires a
mixing zone to flocculate the particles before clarification. Units
are available in rectangular or circular shapes with either flat or
conical bottoms.
-------
* Mixer-clarifiers where the incoming feed is nixed with the sludge
maintained in the unit. This unit basically combines a flocculating
chamber with sludge back-mixing and a settling chamber,
* Plate settlers (Lamella) where inclined plates reduce the distance
particles must fall to be removed. These units are often well
suited to application where space to house the equipment is limited
since the units are mostly vertical rather than horizontal.
Mixers commonly used in floceuLation/clarification units are typically
either oscillating or rotary types. The oscillating types are most applicable
to flocculation processes where very gentle flocculation is required. The
rotary types consist of the paddle wheel and turbine designs. Typical design
values for mixing and flocculation are shown in Table 10,2,1.
TABLE 10.2,1. MIXING AND FLOCCULATLON DESIGN CRITERIA
Detention time Velocity gradient (G)
(minutes) (m/s/m or see"*)
Mixing 0.2 -2 300 - 1,500
Floccuiation 5-30 10 - 100
Fragile floes (e.g.,
(e.g., biological floes) - 10-30
Medium strength floes
(e.g., floes as encountered
in turbidity removal) - 20-50
High strength chemical
floes (e.g., floes encountered
in precipitation processes) - 40 - 100
Source." Reference 15. l»
The major design parameters of a flocculator are:
» Residence time, t seconds;
* Velocity gradient, G, ft/sec/ft or sec"1; and
• Ratio of floe to total volume of suspension, C, dimensionless,
10-91
-------
Residence time is determined from the total flow rate and total flocculator
volume. The velocity gradient for mechanically stirred units can be
determined from:
where ' P • power requirement 3ft—Lb/sec
*)
^ = fluid viscosity, Ib^-sec/ft
V = flocculator volume, cu ft
4
Power requirements are determined from:
(2)
z
o
where: A = paddle area, ft"1
p = fluid density
v = relative velocity of paddles in fluid, fps, usually about
0.7 to 0.8 of paddle tip speed
14
The velocity gradient can also be expressed as:
G = (power/viscosity x volume)1''2 = I (Nm/s )/ (NsAn2)m3]1/?2 (3)
where: N = force CN)
m = distance (m)
s = time (s)
Values of G from 20 to 90 s are typical for flocculation units.
Tapered flocculation employs high entrance values of G and lower values as the
flow progresses to the exit. Values of G x t (where t = seconds of residence
time) ranging from 30,000 to 150,000 are comroonly employed for flocculation in
domestic water treatment. Flocculator retention times o£ 5 to 30 minutes are
typical. Experimentally derived values of G and G x t are advocated for
14
industrial waste applications.
10-92
-------
10.2.1 Process Description
The coagulant/flocculants currently in commercial use are conveniently
classified as inorganic, synthetic organic, and naturally occurring organic
polymers. The following subsections are organized according to these three
categories. Each subsection will highlight the unique aspects and typical
uses of each coagulant/flocculant type,
10.2.1.1 Inorganic Flocculants—
Inorganic coagulants are used primarily for waste streams having dilute
concentrations of constituents that become insoluble during neutralization/
precipitation treatment. A major disadvantage of this technology is that it
adds to the quantity of sludge generated by the precipitation process.
Many soluble salts can function as indifferent electrolytes, typically
following the Schulz-Hardy rule for coagulation effectiveness,
i.e., coagulation of sols is caused by the ions with charges opposite in sign
to the charges on the sol particles; the flocculating power of bivalent ions
is about 20 to 80 tines greater than that of univalent ions, and the
flocculating power of trivalent ions is many times greater than that of
bivalent ions (see Table 10.2.2).15>16
Generally, inorganic coagulants destabilize colloidal particles in the
following manner:
* Repression of the double layer.
• » Charge neutralization.
* Entrapment by sweeping floe.
Repression of the double layer involves increasing the ionic strength of
the solution. As ionic strength increases, the thickness of the layer is
reduced, thus allowing particles to come in closer proximity to each other at
which point VanderWaal forces may cause coagulation. Repulsive forces can
be reduced by charge neutralization. Destabilizing chemicals are added to the
colloid within the stern layer so that the effective charge outside the shear
Q
layer is reduced. In this case, overdosing can cause a charge reversal and
restabilizstion. Entra-poent rsauires the use of large doses of coagulants
10-93
-------
TABLE 10.2.2. THE RATIO OF THE FLOCCULAT10N POWER OF SALTS
WITH Me+, Me-M- AND Me-t-++ IN SOLUTIONS
Sol
Salts
Ratio of the
flocculation power
Ag NaCl2, La(N03)3
1:60:10000
3 NaCl, BaCl2>
1:70:625
ASS
23
1:80:625
Au
NaCl, Bad,
1:60:6660
Me = metal
Source: Reference 15.
IU-94
-------
which form gelatinous hydrolysis produces. These products can effectively
mesh the suspended .matter. -Because massive amounts of coagulants are used in
this procedure, the volume of sludge created is greatly increased,
The three main classifications of inorganic coagulants are:
1 > Aluminum derivatives.
.2. Iron derivatives.
3. Line,
Aluminum Derivatives — In the literature of coagulants the term alum
refers to a commercial aluminum sulfate hydrate, A1?(SQ ),.7H.O. It
also is called paperraakes1 alum or filter alum, and is available either in the
dry form or in solution. Dry alum is available in several grades, with a
minimum aluminum content of 17 percent expressed as Al.O,. Liquid
alum is about a 49 wt percent solution of AL_(-SO, )_• lAHjO, or about
8.3 wt percent aluminum as Al«0-. It can be stored indefinitely without
deterioration. Alum is the most widely used inorganic flocculant.
Although alum may be considered as Al for calculating the composition of
the pure salt that ion does not exist in water environments. It forms
complexes with water to give a compound such as Al(H_0)fi and then
loses protons by hydrolysis to assume a range of either positive or. negative
, 17
charges .
The best range for alum coagulation is pH 5.5 to 8.0, however, actual
removal efficiency depends to a large extent on competing ion and chelant
concentrations. However, if the coagulation rate is too low, increasing the
particle concentration through the use of synthetic organic polyeleccroly tes
18
can improve system performance.
An alternative to aluminum sulfate is sodium a laminate which is
commercially available either in dry form or in solution, with an excess of
base present. It provides a strongly alkaline source of water-soluble
aluminum, particularly useful when addition of sulfate ions is undesirable.
Sometimes it is used in conjunction with alum for pH control. Another
aluminum derivative is polyaluminum chloride (PAC) which is a partially
hydrolyzed aluminum chloride solution with an aluminum content of
10-95
-------
10 we percent expressed as Al.O^* It is reported to provide faster and
stronger floes- than alum in some applications but has yet to achieve
widespread use.
Iron Derivatives — Compared with aluminum, the bydrated ferric ion is more
acidic, it forms stronger complexes with simple anions, and its amorphous
hydroxide is less acidic but the two show a gross similarity in hydrolysis
reactions. Aging characteristics of the polynuclear products of the ferric
ion are more dependent on the anions. Mininmrn solubility of ferric
hydroxide occurs in the pH range 6.8 to 8.4, where concentration of soluble
Fe(lII) species is about 10 " M, but eauilibration with polynuclear
species in solution -may be slow. Ferrous ions form analogous tnonormclear
species but comparable data on tendency to form polymers are not available.
Minimum solubility of ferrous hydroxide occurs near pH 10.7, where
— ft c
concentration of soluble (Fe(II) species would be about 10 * M, but a
tendency for air oxidation to the ferric species complicates the system,
Because of the color of iron compounds, they tend to be used in waste streams
rather than in water supplies.
Liquid ferric chloride, a dark broun oily-appearing solution which is 35
to 45 wt percent FeCl-, is the customary form for flocculant use. Ferric
chloride also is available in solid form. Ferric sulfate is marketed as dry
granules, Fe?(SO, )„. ?H?0. Ferrous sulfate, also known commercially as
copperas, is generally available in dry form with the nominal composition
Lime Derivatives — While lime is used primarily for pH control or chemical
precipitation, it is also used as a co-f locculant . For a summary of
properties j see hydroxide precipitation.
In general, inorganic coagulants are used sparingly in industrial waste
treatment applications. Primary usage is in the precipitation/coagulation of
soluble phosphates and trace metals at municipal POTW's. See Table 10.2.3 for
summary of manufacturers of inorganic flocculants.
10-96
-------
TABLE LO.2.3. MAIN PRODUCERS OF INORGANIC FLOCCULANTS IN THE UNITED STATES
Company Products8
Allied Chemical Corp. a ' f
American Cyanatnid Company A
Associated Metals & Minerals Co. ' • c
Burris Chemical, Inc. a
Cacco, Inc. c
Cities Service Company, Inc. a d
Conservation Chemical Co. c
The Cosmin Corp. e
Diamond Shamrock Corp. C
The Dow Chemical Company c
E.I. du?ont de Nemours & Co., Inc. a c
Essex Chemical Corp. a
Filtrol Corp. a
Philip A. Hunt Chemical Corp. c
Imperial West Chemical Co. a c
Balco Chemical Co. b
NL Industries, Inc. . e
01 in Corp. a
Pennwalt Corp. c
Pfizer, Inc. e
Philadelphia Quartz Co. • f
Quality Chemicals, Ltd. ' e
Reynolds Metals Co. b
Southern California Chemical Co. c
Stauffer Chemical Co. . a
K. A. Steel Chemicals, Inc. ' c
Vinnings Chemical Co. b
aa = alum; b = sodium aluminate; c = ferric chloride; d = ferric sulfate
e = ferrous sulfate; f = sodium silicate.
Source: Reference 19.
10-97
-------
10.2.1,2 Synthetic Organic Floccuiants—
Synthetic organic polymers arc used almost exclusively in the coagulation/
flocculation of industrial heavy metal precipitates.. Typically, synthetic
organic coagulants/flocculsnts are watei—soluble polymeric substances with
average molecular weights ranging from about 10 to greater than 5 x 10 .
If sooae subunits of the polymer molecule are charged, it is termed a
' 19
polyelectrolyte. When the charge on the subunits is positive, the polymer
is termed catioriie; when the charge is negative, it is termed anionic.
Polyelectrolytes containing both positive and negative charges in the same
molecule are termed polyampholytes. Some water-soluble polymers contain
little or no charged subunits (less than 1 percent). These are termed
, . , 19
nomonic polymers,
Polyelectroly tes operate through the mechanism of chemical bridging and
physical enmeshment. The polymer is usually a long organic chain which
contains many active sites with which particles can interact and adsorb.
Bridging occurs where the polyelectrolyte acts as a bridge, joining colloidal
•particles together to form a larger particle. Destablization occurs by
, . , . , . 15,16
slowing down particle motion.
The coagulant/flocculant most generally used in the agglomeration of
metals-containing wastewaters is an anionic organic polyelectrolyte. This is
because metallic precipitates and metal hydroxides in particular, possess a
slight electrostatic positive charge resulting fron charge density separation.
The negatively charged reaction sites on the anionic polyelectrolyte attract
. . 20
and adsorb the slightly positive charged precipitate. However,-
21
studies have been conducted that show that anionic polyeleetrolytes adsorb
onto electronegative suspended particles as well. It is hypothesized that the
attractive adsorption energies between the anionic polyelectrolytes and the
electronegative particles are stronger than the repulsive electrical energies.
Synthetic organic polyelectrolytes are commercially available in the form
of dry power, granules, beads, aqueous solutions, aqueous -gels, and
oil-in-water emulsions, High (M.W. 1-5 x 10 ) and very high (M.W.
5 x 10 ) molecular weight polymers such as anionic polyelectrolytes tend to
be sold as dry products or as oil-in-water emulsions due to increases in
viscosity. Gsr.sTslly "^i^u-^d ~"X--o i s^trcl^'te «a\"3»"p«"« &*ra r\r a £.*»•»-»-<* A h&f*s^e:&
l-U-98
-------
they require less floorspace, reduce the labor involved, and reduce 'the
potential for side reactions because the concentrate can be diluted as used in
20
automatic dispensing systems.
Dosages for treatment of metals-containing wastes generally fall in the
22
range of 0.5 to 2 mg/L with a mg/L being the most common. The
polyelectrolyte and wastewater are initially combined in the rapid-mix section
of the clarifier. Usually, rapid-mix chambers, provide a reactor volume equal
to 5 to 30 seconds of the design flow rate. Excess polyelectrolyte dosing at
this point could be detrimental in that it may. waste chemicals and result in
restabilization of the metallic precipitates.
The most commonly used commercial anionic polyelectrolytes are
19
poly(acrylic aeid-co-acrylaraide) and hydrolized polyacrylamide.
Polyacrylamides are infinitely soluble in water but are limited in practical
applications by viscosity. As a polyelectrolyte, polyacrylamide exhibits
sensitivity to salts and variations in pH. For example, excess salt will
cause an exponential decrease in viscosity with increasing salt
19
concentration. In addition, at alkaline and intermediate pH's, the
flocculating power will increase but at pH's lower than 4.5 flocculating power
is decreased.. Bridging theory suggests that by increasing the number of NH~
groups hydrolyzed to OH groups, the effectiveness of the polymer should
21
increase because the polymer coil becomes more extended. Experimental
evidence (see Figure 10.2.2) shows that at pH 6.0, the chain is fully extended
(20 percent hydrolysis) while at pH 4.5, the OH groups are unionized and hence
increasing the hydrolysis level does not extend the length of the chain, thus
decreasing effectiveness.
Table 10.2.4 lists typical properties of two low anionic charge (1 to
22
10 percent anionicity) polyacrylamides. The liquid polyelectrolyte is an
oil-in-water solution which is diluted to 3 percent concentration upon use.
High shear action (above 475 rpm) is not recommended during make-up since it
can cause degradation of the high molecular weight flocculant. . Fifteen to
30 minutes of mix time in a flocculant make-up tank is recommended to insure
complete dissolution and partially hydrolyze the polymer, prior to
22
introduction into the rapid-mix tank. The dry polyacrylatoide is usually
•dissolve— ^ o 3 0s f-Q ^er^^t co^*j^i.ri ^ith cct^^^ets ^•5«s«!^O'->*~'«'"*T-» *-»*~o •!-»-?-'r>c?
-------
Z v.w r SO!,COO pN ' £.0
I «„ • 601,000 pn « 4 5
£1 S. « 111,COO pH • S.D
* 5, " in.COO pm « S
10 I! 20
PERCENT KTDRCLTS1S
Figure 10.2.2.
Settling rate ratio versus hydrolysis
for linear polyacrylamide- '
LO-100
-------
TABLE 10.2,4. PROPERTIES OF LIQUID 'AND DRY ANIONIC POLYACRYLAMIDES*
Typical Properties3
Appearance
Specific gravity at 25°C
Typical effective viscosity
as is*, at 25°C, cps
Typical viscosity,**
as X solution, cps
Freezing point
Flash point,
Tag Closed Cup
Shelf life
Environmental Properties
BOD, mg/L
COD, mg/L
870
7,060
Opaque white liquid
1.0 + .02
0.5
1.0
2.0
300-600
0"C
670
1,295
5,300 '
20°C
585
1,130
4,530
40°C
400
790
3,945
0°F (-15°C)
>200°F (93"C)
9 months
Typical Properties
Appearance
Degree of anionic charge
Bulk density
pH of 0.5% solution @
Viscosity***, cps
% solution
0.1
0.5
1.0
Environmental Propert
BOD
COD
25SC (77°F)
0°C 20°C_
50 40
300 250
1,400 1,200
ies
approximately 0
9,800 mg/L
White powder
Low
43-45 lb/ft3
(688-720 kg/ni3)
5.4
40°C
30
200
1,000
*Viscosity at infinite shear speed (approximates the pumping situation)
**Brookfield.
***Brookfield,
shagni£Ioc 1620A, American Cyanimid,
bMagnifloc 834A, American Cyanimid.
Note: Based on a 1 percent solution.
Source: Reference 22.
10-101
-------
after 30 minutes. For best results, it is recommenced that the solution be
further diluted (100:1) with clean water prior to feeding into the rapid-mis
22
tank. Stock solutions are usually stable for at least 2 weeks.
10.2.1.3 Natural Organic Polymers—
Current tlocculants derived from natural products include starch, starch
derivatives, 'plant gums, seaweed extracts, cellulose derivatives, proteins,
and tannins. Starch is the most widely used of thes'e products, 'followed by
guar gum. Although price/kilogram for natural products tends to be low
relative to synthetic flocculants, dosage requirements tend to be high. In
addition, the composition of natural products tends to fluctuate, and they are
more susceptible to microbiological attack which creates storage
23
problems. In recent years, the most promising of natural organic polymer
flocculation technologies is a process utilizing insoluble starch xanthate
(ISX).24
The ISX process was originally developed at the U.S. Department of
Agriculture under a grant from the EPA. Production of ISX involved xanthating
a relatively inexpensive chemically cross-linked, insoluble natural starch
compound to form an..anionic polymer capable of coagulating/flocculating heavy
metals.
The ISX process has been demonstrated to be capable of producing an
effluent with very low residual metal concentrations Csee Table 10.2.5).
The resulting ISX-metal sludge is said to dewater to 50 to 90 percent solids
because it is nongelatinous. In addition, claims indicate that metal can be
recovered from the ISX-metal sludge by acidification or incineration of the
-, j 26
sludge.
Two methods of ISX treatment have been applied on a commercial scale.
The first method used in conjunction with commercial treatment, involves
mixing an ISX slurry with neutralized/precipitated wastewaters in a reaction
tank. The treatment is effective over a wide pH range, but for optimum
coagulation/flocculation performance, this technology is typically operated at'
pH 9 in conjunction with a cationic polymer. In the coagulation reaction, ISX
acts as an ion exchange liquid, bonding with heavy metal ions in exchange for
ISX and nickel:
10-102
-------
TABLE 10.2.5. HEAVY METAL REMOVAL EFFICIENCIES USING STARCH XANTHATE
AS DETERMINED BY USDA
Metal
Copper
Nickel
Cadmium
Lead
Trivalent Chromium
Silver
Zinc
Iron
Manganese
Mercury
Concentra
Influent
31.8
29.4
56.2
103.6
26.0
53.9
32.7
27.9
27.5
100.0
tion, mg/L
Effluent
0.007
0.019
0,009
0.025
0.003
0.245
0.046
not detectable
1.630
0.004
Source: Reference 25,
10-103
-------
s s
2ROCSNa + Ni*2 =
The second method involves Busing ISX as a filter precoat to polish
effluent. In a typical operation, a sludge slurry would be pumped from a
holding tank or clarifier to a precoated filter for dewatering of the sludge
and removal of the metal ions remaining in solution. Table 10.2.6 contains
removal efficiencies for a facility using ISX as a filter precoat.
Since ISX is susceptible to biological attack, it is typically shipped
and stored under refrigeration (40°F)^ Shelf life, is approximately 6 months
and typical costs are $1.95 lb- for a 25 Ib container and 4l.70/lb for a 250 Ib
27
container.
. Daily preparation of the ISX slurry would involve mixing predetermined
amounts of ISX powder and water in a chemical feed tank. The slurry should be
prepared at the ratio of approximately 2 pounds of ISX per gallon of water,
ISX dosage is determined from- laboratory testing. Calibration of the
metering system involves monitoring the flow rate and adjusting, the control
system to deliver slurry in the required amount. The average capacity of ISX
is in the range of 1«1 to 1.5 milliequivalents of metal ion per gratn ISX.
Thus, for a divalent nickel ion, one pram of ISX would remove 32 to 43 mg of
25
nxckel ions from solution. Maintenance of this systera involves periodic
flushing of the lines to prevent build-up of ISX and restriction of the lines,
and periodic checks of the metering system calibration.
In addition to storage and handling difficulties, disposal of process
residuals or sludges is a major problem associated with the sterch xanthate
process. Laboratory test results indicate that heavy metal removal capacity
28
is approximately 0.0011 moles per gram of starch. Consequent ly ,
relatively large sludge volumes will be produced for the quantity of heavy
metals removed. Conventional land disposal does not appear to be an
environmentally acceptable alternative because the organic structure of the
starch xantbate-metal sludge can decompose rapidly and release the metal to
the environment. Incineration is being considered for possible metal recovery
but off-gas scrubbing facilities would be necessary to insure that heavy
-------
TABLE 10.2.6. METAL REMOVAL RESULTS USING ISX AS A FILTER PRECOAT
Metal
Initial
Concentration3
tmg/L)
ISX Treated
Concentration
(aiR/L)
Cr
0.8
0.02
7.0
0.02
2.5
0.10
aBefore ISX treatment.
Source: Reference 28.
10-105
-------
these stack gas control facilities may be prohibitive. Also, the heavy metals
collected in the scrubber liquor would again have to be removed before the
liquor could be reused or discharged to a receiving stream,
10.2,1.4 Pretreatment and Post-Treatment Requirements—
Coagulation/flocculation is a well established technology, and in
general, is very reliable. It is used primarily to treat aqueous metals-laden
waste screams. The properties of the waste being treated which can affect
performance include:
• Flow variations;
» Solids concentration variations;
• pH variations;
* Temperature variations;
* Cyanide content;
• Hexavalent chromium concentration; and
• Oil and grease concentration.
«
The effect of flow variations appears mainly in the sedimentation step.
Temperature variations can also cause upsets in sedimentation by creating
undesirable thermal currents. Changes in solids concentration and pH can
affect the performance of the coagulation and flocculation process in systems
where the agglomeration rate is a function of these parameters. Also,
compounds in the wastewater that interfere with coagulation (such SB sulfides
and mercaptides) can result in reduced agglomeration effectiveness. To
29
minimize these effects, equalization basins are generally recommended. In
addition to creating an influent of more consistent quality, sulfides or
mercaptides can be oxidized to a less reactive or inert state. Also, in
systems where pR influences the agglomeration rate, pH adjustment may be
required. For a discussion of cyanide destruction, chromium reduction, and
oil removal technologies refer to Sections 13.0, 10.3 and 10.1, respectively.
10-106
-------
10.2.2 Process Performance
As previously indicated, most heavy metal coagulation/flocculation
applications involve the use of an inorganic or polymeric reagent. The many
disadvantages of naturally occurring organic polymers has currently limited
their use to a feu select applications. For example, insoluble starch
xantbate which was once used at over 100 facilities is now utilized at less
27
than 50 as a result of storage, application, and disposal problems.
Process performance and costs for heavy metal coagulation/flocculation
systems are very sensitive to coagulant dosage, type, and flow rate.
Figures 10.2;3 and 10.2.4 illustrate the effect of iron dose and clarifier
o A
over flow rate on arsenic and selenium removal efficiencies. An anionic
polyelectrolyte was introduced into the feed line to the clarifier to assist
in floceulation. These results show that arsenate removals exceeded
90 percent at clarifier overflow rates up to 1,200 gpd/ft . Selenium
removal {56 to 89 percent) was limited by the fraction of selenate (which is
not adsorbed by Fe(OH)-j) in the waste . stream. Minimum iron and polymer
doses for good performances were 14 mg/L and 0.15 mg/L (pH range 6.2 to 6.5),
respectively.
Tables 10.2.7 through 10.2.9 demonstrate the various treatment options
available for effectively removing such heavy metals as lead, zinc, cadmium,
manganese, copper, and nickel. In Table 10.2.7, lime (for precipitation
and coagulation) was combined with either of two polymers, Magnifloc 1561/1820
or Pe'rcol 728 for flocculation. Both types of polyelectrolyte worked equally
as well as a flocculant in removing lead, zitic, and cadmium. The optimum
dosage for Magnifloc 1561/1820 was determined to be 1.5/0.5 mg/L,
respectively, while 1.0 mg/L of Percol 728 was sufficient for greater than
99 percent removal. Table 10.2.8' shows the removal of cadmium, copper, 'iron,
lead, manganese, nickel, and zinc using alum (35 mg/L) as a coagulant and an
anionic polyelectrolyte (1 mg/L) &s a flocculant. Sodium hydroxide is used to
adjust the influent pH (7,3 to 8,9) to the 8.4 to 9.25 range for precipitation.
Table 10.2.9 presents performance data for a system using ferrous sulfate
(Fe:Ni ratio = 0.7) as a coagulant to remove nickel from an aqueous waste
32
H* " " "
10-107
-------
100
1 m3/m2 per day = 24.5 qpd/ft2
0 1C- 20 30 id su
CLARIFIER OVERFLOW RATE, rr,3/m2 per day
Figure 10.2,3. Effect of iron dose and clarifier1overflow rate
on arsenic removal efficiency,
Source: Reference 30.
10-108
-------
100
90
80
70
o
2
u
cr
£Q
30
*j n
1 mj/rn per day = 2£.5 gpd/sq h
10 20 30 10
OVERFLOW HATE, m3/m2 per day
50
Figure 10.2.4. effect of iron dose and clarifier overflow rate
Source: Reference 30.
10-109
-------
TABLE 10.2.7. METALS REMOVAL USING LIME AND AN10NIC POLYELECTROLYTES
Clarifier
Parameter
pH, units
Suspended solids, mg/L
Calcium, mg/L
Lead , mg/L
Zinc, tng/L
Dissolved solids, mg/L
Cadmium^ mg/L
production
Influent
5.65-6.78
640
60
59
72
535
0.45
operation
Effluentb
10.00
35
60
0.77
1.26
450
0.01
Dual media
filter3
Effluent
7.0
7.0
-
0.25
0.^1
-
™
aTypic£i filter effluent cosl/sand.
'"'Typical clarified effluent lime/polymer treatment.
Source: Reference 31.
10-110
-------
TABLE 10.2.8. METALS REMOVAL USING ALUM AND AN AWIONIC POLYELECTROLYTE
Before new After newa
Parameter (mg/L) treatment plant treatment plant
Cadmium
Copper
Iron
Lead
Manganese
Nickel
Total suspended solids
Zinc
0.036
0.084
12.0
1.5
2.2
LT 0.13
16 (avg.)
8.2
0.02-9
0.02
0.25
0.09
0.25
0.06
6.0
0.29
*35 mg/L A12(S04)3
1 mg/L anionic polyelectrolyte
?H = 8.4-9.25
Flow rate = 10 mgd
Source: Reference 31.
TABLE 10.2.9. RESIDUAL NICKEL CONCENTRATIONS FOR VARIOUS POLYMER
ADDITIONS: Fe:Ni = 0.7, CT = 0 mg/L
Anionic polymer Cationic polymer No polymer
concentration (mg/L) concentration (mg/L) (control) •
pH = 9
Soluble
Total
pH = 10
Soluble
Total
0.
5.
5.
0.
0.
1
1
6
22
40
0.
1.
1.
0.
0.
5 '
25
30
12
15
1.
0.
0.
0.
0.
0
60
70
10
20
0.
0.
0.
0.
0.
1
73
91
12
31
0.
0.
0.
0.
0.
5
70
80
15
32
I.
0.
0.
0.
0.
0
70
85
12
30
1
1
0
0
.6
.8
.12
.50
Source: Reference 32.
10-111
-------
It was'Concluded that the addition of cationic and anionic polymers slightly
enhanced settleability at both pH 9 and 10. Lime was used as the
precipitation and neutralization reagent.
10.2.3 Process Costs
Table 10.2.10 contains the purchased equipment and installation costs,
and annualized operating costs for a continuous coagulation/flocculation
treatment system. The system consists of a continuous flocculation/
clarification unit, sludge holding tank(s), and a filter press. The
floeculation/clarification unit size is a function of the volumetric flow
rate. The influent to the unit is assumed to contain 200 tng/L of heavy metals
which have been previously precipitated with sodium hydroxide to form
approximately 400 mg/L of suspended solids. The overflow from the
clarification section is assumed to be solids-free, while the underflow is
assumed to contain 6 percent solids. The coagulant alum is added (150 mg/L)
along with the flocculant, Magnifloc 1820A (1 mg/L) in the flocculation tank
prior to the clarifier. The sludge holding tanks (10 hours retention) and the
filter press (8-hour cycle) have been sized to handle the solids content in
the underflow. Capital and amualized operating costs are based on
assumptions previously presented in Section 10.1.2,
A large percentage of total annual costs for the continuous coagulation/
flocculation system developed for this section are a result of sludge disposal
costs. Sludge production is increased roughly 20 percent by the addition of
alum with an equivalent increase in sludge transportation and disposal costs.
A 30 percent savings in reagent costs can be realized by using TeSQ,
l$14S/ton) instead of A1?(SO,), [£205/ton), but sludge generation will
be equivalent, if all the aluminum is precipitated in the clarification
section as an hydroxide.
Labor costs for this treatment technology are also a large percentage of
the overall annual operating costs. This is due to the high operator skills
required in making the coagulant/floceulanc reagent additions. In addition,
since the dosage requirements for coagulants such as alum, ferric chloride,
stnd f^xroysr sulf],t@ sire r:orj£tC!*cbi,O!i!£tiT*LC fire^'J0"*" ^sr. "tssts srs ~£~eess~" *"£
prevent underdosing or overdosing.
10-112
-------
TABLE 10.2.10. CONTINUOUS COAGULANT/FLOCCULANT COST DATAa
Purchased equipment and installation (PE&I)
Flocculator/clarifier
Sludge molding tankls)
Filter Press
Total capital investment (3601 PEI)
Annual operating costs ($)
Operating labor C$20/hr)
Maintenance (62; TCI)
General plant overhead (5.81 TCI)
Utilities (2% TCI)
Taxes and insurance (1% TCI)
Chemical costs: A^CSO^ ($2Q5/ton)
1820 A ($1.29/lb)
Sludge transportation ($Q.25/ton-mile)
Sludge disposal ($200/ton)
Annualized capital (CFR = 0.177)
Total annual costs
Cost/1,000 gallon
1,000
$
18,000
3,000
11,000
32,000
115,200
72,000
6,900
6,700
2,300
1,200
900
100
200
11,300
20,400
122,000
17
Flow rate (gph)
10,000
50,000
3,000
15,000
68,800
244,800
72,000
14,700
14,200
4,900
2,500
9,200
800
2,100
112,600
43,300
276,300
4
100,000
140,000
20,000
60,000
220,000
792,000
72,000
47,500
45,900
15,800
7,900
92,300
7,700
21,000
1,126,100
140,200
1,576,400
2
al987 Dollars.
10-113
-------
10.2.4 Overall Process Status
Coagulation/fioceulation is a well-developed process widely used for many
industrial wastewaters containing suspended and colloidal solids. The
equipment used is relatively simple, readily available, and can often be skid'
mounted in a modular design. In many cases, coagulation/flocculation can be
added to existing process trains with only minor mocif icatior.s. For high
volume applications, the cost of this technology drops dramatically improving
economic viability. In addition, the process is often improved by high ionic
strength and is applicable to high influent metal loadings.
Disadvantages and primary environmental considerations result from a
metals laden high-water-content sludge which must be treated (i.e.,
solidification, encapsulation, etc.) and then disposed. In addition, the
process is also not readily applied to small intermittent flows and many of
the coagulants used (Al ISO K, Fed,, etc.) form corrosive
solutions. Finally, process efficiency is highly sensitive to initial
contaminant concentration and the surface area, of the primary floe formed in
the rapid-mix chamber.
10-114
-------
REFERENCES
1. Process Design Manual for Suspended Solids Removal, U.S. Environmental
Protection Agency, EPA-625/l-75-QQ3a. January 1S75,
2. Sundstram, D.W., and H.E. Kiel. Wastewater Treatment. Prentice-Hall,
Inc., Englewood Cliffs, NJ. 1979.
3, Clark, J.W. et al. Water Supply and Pollution Control. 3rd. Edition,
Harper & Row. 1977.
4» Chillingworth, M.A. et al., Alliance Technologies Corporation.
Industrial Waste Management Alternatives Assessment for the State of
Illinois. February 1981.
5. Arthur D. Little, Inc. Physical, Chemical, and Biological Treatment
Techniques for Industrial Wastes. U.S. EPA SW-14S. November 1976.
6. Manual on Disposal of Refinery Wastes, Volume on Liquid Wastes,
Chapter 9—Filtration, Flocculation and Flotation. American Petroleum
Institute, Washington, DC. 1969.
7. Beychok, M.R. Aqueous Wastes from Petroleum and Petrochemical Plants.
John Wiley & Sons, New York. 1967.
8. U.S. EPA. Treatability Manual. Volume III. Washington, DC.
EPA-600/8-80-042c,d. July 1980,
9. 'Monk, R.D., and Willis, J.F. Designing Water Treatment Facilities.
American Water Works Association Journal. February 1978.
10, Janson, C.F. et al. Alternatives for the Treatment of Heavy Metals.
AICKE Paper No. 149c. November 1981.
11. Lanouette, K.H. Heavy Metals Removed. Chemical Engineering.
October 17, 1977.
12. U.S. EPA. Reducing Water Pollution Costs in the Electroplating
Industry. EPA-625/5-85-016. September 1985.
13. Cushnie, G.C., Centec Corporation. Navy Electroplating Pollution Control
Technology Assessment Manual. Naval Civil Engineering Laboratory
CR84.019. February 1984.
14. Cushnie, G.C. Removal of Metals from Wastewater: Neutralization and
Precipitation. Pollution Technology Review, So. 107, Noyes Publications,
Park Ridge, NJ. 1984.
15. Jurgensons, B,, and M.E. Straumaris. Colloid Chemistry. 2nd. Edition,
HacMi.1 Ian Company, Hew "ork, »T.
10-115
-------
16. Mysels, K.J. Introduction to Colloid Chemistry. Interscience
Publishers, Inc., New"York, NY. 1959.
. 17. Kirk-Othtner Encyclopedia of Chemical Technology. Vol. 24, 3rd. Edition,
John Wiley 4. Sons, New York, NY. 1981.
18. Clifford, C., Subramonian, S., and T.J. Sorg. Removing Dissolved
Inorganic Contaminants front Water. Environmental Science & Technology,
Vol. 20, No. 11. 1986.
19. Kirk-Othmer Encyclopedia of Chemical Technology. Vol. 18, 3rd. Edition,
John Wiley & Sons, New York, NY. 1986.
20. Capaccio, R.S., and R.J. Sarnelli, Flocculation and Clarification.
Plating and Surface Finishing. October 1986.
21. Caskey, J.A., and E.J. Primus. The Effect of Anionic Polyacrylamide
Molecular Conformation and Configuration on Flocculation Effectiveness.
Environmental Progress. May 1986.
22, American Cyanamid Technical Brochure. Magnifloc 1839 A. 1986.
23. Kirk-Othtner Encyclopedia of Chemical Technology. Vol. 10, 3rd. Edition,
John Wiley & Sons, New York, NY. 1981.
24. Wing, R.E. Dissolved Heavy Metal Removal by Insoluble Starch Xanthate
(ISX). Environmental Progress. November 1983,
25. R.E. Wing. "Processes for Heavy Metal Removal from Plating
Wastewaters." First Annual Conferenceon Advanced Pollution Control for
theMetal Finishing Industry, U.S. Environmental Protection Agency,
Washington, DC. EPA-600/8-78-010. May 1980.
26. R.E. Wing, et al. Removal of Heavy Metals from Industrial Wastewaters
Using__InsoluDle Starch Xanthate. U.S. Environmental Protection Agency,
Washington, DC. EPA-6QC/2-78-085. May 1978.
27. W. Stout. Stout's Supply telephone conversation with Stephen Palmer,
Alliance Technologies Corporation. February 10, 1987.
28. U.S. EPA. Sources and Treatment of Wastewater in the Nonferrous Metals
Industry. EPA-600/2-60-074. April 1980.
29. MITRE Corporation. Manuel of Practice for Wastewater Neutralization and
Precipitation. EPA-60G/2-81-148. August 1981.
30. D.T, Merrill, et al. Field Evaluation of Arsenic and Selenium Removal of
Iron Coprecipitation. Journal of the Hater Pollution Control
Federation. January 1986.
10-116
-------
31. Osantowski, R., and J, Ruppersberger. Upgrading Foundry Wastewater
Treatment. 39th. Industrial Waste Conference, Purdue University. 1982,
32. McFadden, F., Benefield, L., and Reed, R.B. Nickel Removal from Nickel
Plating Waste«ater Using Iron, Carbonate, and Polymers for Precipitation
and Coprecipitation. 40th, Industrial Waste Conference, Purdue
University, 1983.
10-117
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10.3 CHEMICAL REDUCTION
Chemical reduction is a reaction in which one or more electrons are
transferred to the chemical being reduced (reductant) froir, the chemical
initiating the transfer (the reducing SRent). Chemical reduction can also
be defined as a change in oxidation states where the oxidant (reducing agent)
2
is an electron donor such as zinc in the reaction:
Zn - Zn*+ + 2e
The reductant is the substance which accepts electrons:
Cu+* + 2e * Cu
The overall reaction is called a reduction-oxidation (redox) reaction;
2n + Cu++ = Cu + Zn**
Redox processes are very common in aqueous systems since most organic and
3
many inorganic reactions involve oxidation and reduction. It! reactions
involving covalent bonds, the gain or loss of electrons by an element may not
be clearly defined. The assignment of electrons to an atom is thus carried
out according to rules. If two atoms share electrons in a covalent compound,
the electrons are arbitrarily assigned to the atom that is more
electronegative. If an electron pair is shared by two atoms of the .same
electronegativity, the electrons are split between them.- After this division
of charges has been made, the charge remaining on the atom is known as its
oxidation number or state. The suet of oxidation numbers is equal to zero for
2
molecules and is equal to the formal charge for ions.
In principle, the equilibrium composition of a- redox system can be
determined front a thertnodynatiie analysis as in'the case of acid-base
reactions. Many inorganic redox reactions have fast reaction rates and
chemical equilibrium is approached within typical process times. Redox
reactions involving organic compounds, however, are often slow at ambient
. . . 2 •
conditions.
10-118
-------
10.3,1 Process Description
Chemical reduction as a waste treatment process is an established and
wel1-developed technology- The reduction of bexavalent chromium's valence
state to decrease toxicity and encourage precipitation is presently used as a
treatment technology in numerous electroplating facilities. Major advantages
of chemical reduction uhen used to reduce hexavalent chromium is operation at
ambient conditions, automatic controls, high reliability, and modular process
equipment. Process equipment typically requires a tank for pH adjustment
and reduction, metering equipment, ORP (oxidation-reduction potential) and pH
controls and instrumentation, mechanical agitation, adequate venting, and
.4
separate tanks for subsequent precipitation and sedimentati-on. The
retention'time in tne reduction tank is pH dependent but should be at least
four times the theoretical time for complete reduction.
A number of chemicals are used as reducing agents. The most common
include; sulfur dioxide, sodium metabisulfite, sodium bisulfite, ferrous
sulfate, and sodium borohydride. Other reducing agents which can potentially
be used for wastewater treatment are dithiocarbonate, hydrazine, aluminum,
zinc, and fortnaldehyde. The prevalent reducing agents are discussed in the
following subsections.
Sulfur Dioxide —
For waste streams which contain chromates, gaseous sulfur dioxide is a
widely used reducing agent. The net reaction involves chromic acid and
sulfurous acid (produced through the reaction of sulfur dioxide and water) as
follows:
3H2S03 + 2H2CrC>4 = C^CSO^)} + 5H20
Because the' reaction proceeds rapidly at low pH,.an acid (typically
ric) is added to maintain the pH between 2 and 3. To prevent th
release of sulfur dioxide during treatment, a pH of approximately 3 is
recommended. ' A.t
(see Figure 10.3.1).
sulfuric) is added to maintain the pH between 2 and 3. To prevent the
ulf
recommended. ' A.t pH levels above 5, the reaction rate slows drastically
10-119
-------
50 r
0.5 1 2
RETENTION, minutes
10 20 50
Figure 10.3.1. Effect of pH on chromium reduction rate.
Source: Reference 7.
10-120
-------
Figure 10,3,2 shows a typical wastewater treatment process schematic for
the reduction of chromates. The ORP control set point for this process varies
by approximately 150 millivolts per change in pH unit, with S0_
ft ?
automatically metered to maintain ORP in the 250 to 300 range * (see
Table 10.3.1). Consumption of SO will normally average 50 to 100 percent
of stoiehiome trie requirements. Dissolved oxygen or reducible organics will
consume a significant portion of the reducing agent if the reaction vessel is
open to air.
Sulfur dioxide as with all reduction processes can be employed either as
a batch treatment or as a continuous process. Retention time is typically 30
to 45 minutes at a pH of 3, and reactor vessels should be sized accordingly.
Theoretical chemical requirements per pound of chromium reduced are 2 Ibs of
S0_ plus 35 mg for each liter of water being treated. These
relationships, however, should be confirmed by field tests (see Table 10.3,2
for summary of treatment levels).
Sodium Metabisulf ite and Sodium Bisulfite —
Sodium Metabisulfite (Na^S^Oj) and Sodium Bisulfite (NaHSQ.,) are
soluble sulfite salts used as alternatives to gaseous SO* for the reduction
of hexavalent chromium. These salts (see Section 10.1.3) for a description of
physical properties) are available either as a dry powder flake (70 to
8
72 weight percent) or solution (44 to 60 weight percent). The product is
shipped either as flake in drums or as solutions in tank cars or tank trucks.
Reagent is added either from storage in the case of liquid reagents or from
rapid-mix tanks when using flakes. The reaction when using sodium bisulfite
as a reducing agent is:
3NaHS03 + 3H2S04 + 2H2Cr04 = Cr2( 804)3 + 3NaHS04 * 5H2°
Sulfuric acid is added to depress the pH of the wastewaters to the
optimum pH range of 2-3 (see Figure 10.3.1) as well as provide the required
hydrogen for reaction completion. Table 10.3.3 lists a summary of treatment
levels obtained by this technology. '
In this systsn, chro'ni.tic besri~g wastes are separated fros the ether
metals waste streams and collected in a flow equalization chamber where flow
and pE deviations are averaged. The equalization chamber is equipped witn a
10-121
-------
FLOW
EQUALIZATION
REDUCTION
PRECIPITATION/
CLARIFICATION
CHROMATE
WASTE
'REDUCING
AGENT
ACID
ALKALI
SLUDGE
LINE
SLUDGE HOLDING TANK
__ CHROMATE
FREE WASTE
SLUDGE
DISPOSAL
Figure 10.3.2. Continuous chromium reduction!precipitation system.
10-122
-------
TABLE 10,3,1, RELATIONSHIP BETWEEN OR? AND HEXAVALENT CHROMIUM CONCENTRATION
ORP
590
570
540
330
300
Cr
40
10
5
1
0
+6
ppm
ppm
ppm
ppm
Source: Reference 7.
TABLE 10,3.2, SUMMARY OF TREATMENT LEVELS REPORTED FOR HEXAVALENT
CHROMIUM WASTES
Reduction
agent
Sulfur dioxide
Chromium4"" Concentration (mg/L)
Initial Final
0.3-1.3
1,300 1.0
——. n
0.01
0.05
0.23-1.5 . 0.1
Source: Reference 6.
10-123
-------
TABLE 10.3.3. SUMMARY OF TREATMENT LEVELS REPORTED FOR HEXAVALENT CHROMIUM,
Chromium"1"" Concentration (rag/L)
Reduction Agent Initial Final
Bisulfite
Bisulfite
Bisulfite plus
bydrazine
Metabisulfite
Metabisulfite
Metabisulfite
Metabisulfite
Source: Reference 7.
140
8-20.5
70
0.7-1,0
0.05-0,1'
0.1
0.5
0.025-0.05
0-. 1
0.001-0^.4
10-124
-------
level controlled pump that delivers the wastewater to the reduction unit.
Retention time is typically 30 to 60 minutes. Acid, usually sulfuric, is
added at a point just prior to the reduction tank. Bisulfites are added
directly to the reduction chamber by means of a metered feed system with pH
and ORP controls. Retention times for the reaction tank are typically
10-60 minutes with theoretical reagent requirements of 1.5 Ibs of NaHSO« and
1 lb of H?SO, per pound of Cr reduced. The trivalent chromium is
removed by precipitation. Usually lime or caustic is added to increase the pH
between 7.5 and 8.5 for minimum solubility of chromium hydroxide. Theoretical
reagent requirements for precipitation would be 2.2 Ibs Ca(OH)9 or 2.5 lb
7 ^
NaOH or 3 lb Na?CO. (see Section 10.1 for retention times and equipment
specifications).
Treated wastewater is discharged to a mixer/clarifier where a flocculant
may be added to improve hydroxide precipitate settling characteristics. The
overflow from the clarifier is then discharged to the sewer system, while the
solids in the underflow are collected in a holding tank for subsequent
dewatering (see Section 10.2).
While this type of system is prevalent, many plants experience excess
consumption of reducing agents. The major cause of excess sulfite consumption
is hypothesized to be the dissolved oxygen present in the chromium
•wastewaters. For example, based on stoichiometry, one mole of oxygen will
consume two moles of sulfite ion:
+ 02 = 2
Oxygen molecules from the gas phase are transferred to the liquid phase
in proportion to the difference between the existing concentration and the
equilibrium concentration of gas in solution. Since chromium reduction
reaction vessels are usually open and the reaction is not instantaneous,
oxygen diffusion into the chromium waste solution will continuously consume
reducing agent. Therefore, to prevent excess reagent consumption due to
dissolved oxygen and eliminate hydrogen sulfide odor problems, it is
recommended that process vessels be enclosed and adequately vented.
vei »s suur Dioxide sys~
terns) is an excess consumption of acid and bases. Since historically chromium
reduction has consisted of first a pH depression to reduce chromium to a
10-125
-------
trivalent state followed by a pH elevation to precipitate the chromium ions as
hydroxides, acid and base reagent consumption adds significantly to the
operation and maintenance of a chromium reduction system. This problem is
compounded by the sodiun sulfite salts which often fora sodium hydroxide as a
reaction byproduct, thus requiring an even greater excess of acid during the
pH depression step. Therefore, chromium— bearing waste streams are typically
segregated and treated separately to reduce reagent consunption. The reduced
chromium-bearing stream can then be either precipitated/clarified separately
or combined with other metal-bearing streams for further treatment.
Ferrous Sulfate —
Ferrous sulfate heptahydrate solids (FeSO, .7H_0) are water soluble,
blue-green crystals having a density of 1.898 g/cm and a melting point of
64CC. Most ferrous sulfate ie waste product derived from the pickling of
steel surfaces in the steel industry. Supply exceeds the demand, and the
major portion of the waste presents a serious disposal problem. Ferrous
sulfate is available either in flakes or solution form. In moist air the
flakes oxidize to- basic iron (III) sulfate (Fe«(SO, ),, ) . Aqueous
solutions are also subject to oxidation and are very sensitive to alkalis,
temperature, and light.
In waste treatment applications, ferrous sulfate has been used in a
variety of ways. Three methods reported in literature are acid reduction,
alkaline reduction, and ferrite coprecipitation.
Acid reduction of hexavalent chromium with ferrous sulfate consists of
adding ferrous sulfate heptahydrate to an acidic hexavalent chromium solution
+2
(pH 2-3). The ferrous ion (Fe ) will react with the hexavalent chromium,
reducing the chromium and oxidizing the ferrous ion to basic iron (III)
sulfate. The reaction occurs as follows:
6FeS04 + 7H20 + 6H2S04 = Cr2(S04)3
3Fe2(S04)3 «- 15H20
In terms of reaction rate retention times, pH and chemical metering
controls, acid ferrous sulfate reduction is similar to other sulfur-based
reduction systems such as sodium netabisulf ite and sodium hydrosulf ite. The
10-126
-------
main advantage of this process is an abundant and inexpensive supply of
ferrous sulfate. Disadvantages include excess acid and base requirements to
adjust the wastewater pH to 2 for reduction and then back to 8.5 to 9 for
precipitation. Another disadvantage is that three moles of ferrous ions
are required per mole of hexavalent chromium reduced. In addition the
precipitation of the ferric ion (Fe ) as a hydroxide contributes greatly
to the amount of sludge generated. One study found that the use of ferrous
sulfate rather than a soluble sulfite such as sodium hydrosulfite, for the
reduction of h«xavalent chromium results in a sludge product 31 times as great
12
as the volume of sludge produced by the bisulfite process.
Alkaline reduction of hexavalent chromium with ferrous sulfate was a
process evaluated under a grant by Arizona State University. " ' The
main advantages to this process are a rapid reduction of chromate at pH levels
between 8 to 10 (eliminating the acid depression step) and a reduction of
process equipment since the process can be accomplished in the same reactor as
the neutralization/precipitation process. Disadvantages include sludge
generation and a lack of control in chemical metering.
This process, like acid ferrous sulfate reduction is capable of reducing
chromate concentrations to 0.05 mg/L, The process produces considerably more
sludge and is consequently more expensive than the conventional process of pH
reduction and the use of SO.. However, for hexavalent chromium
concentrations of 10 mg/L or less, ferrous sulfide reduction economics may be
., . 15
worth considering.
Ferrite coprecipitation is a process similar to acid ferrous sulfate
reduction for the conversion of soluble metal ions to insoluble metal
hydroxides or ferrites. The process, which was developed in Japan, involves
the mixing of ferrous sulfate heptahydrate with a heavy metal-bearing
wastewater. The ferrous ion will coexist with the heavy metal ions in
solution. Alkali is added to neutralize the acidic solution and a dark green
hydroxide is formed as follows:
XM++ + Fe^O-X) + 6(OH)~ = MxFe(3_x) (OH)6
X = 1,2,3
10-127
-------
In a variation on the traditional ferrous sulfate reduction process,
oxidation with air is performed during which dissolution and complex formation
occur yielding a black ferrite as follows:
MxFe(3_x) (OH)6 + 1/202 = HxFe(3_x)04 + 3H20
Tables 10.3.4 and 10.3.5 contain data from a test facility which uses a
batch process to treat vastewater (4.5 gpm) and an installation which treat
off-gas scrubber licmor from a municipal refuse incinerator (20 gpm). These
data show that ferrite coprecipitation is an effective process for the removal
of heavy metals. However, little published data exist on the success o£
tnis process in the United States. It is reported to be labor intensive and
like all iron precipitation technologies generates a voluminous sludge
16
product .
Sodium Borohydride —
Sodium bo.rohydri.de (NaBH, ) is a mildly alkaline reducing agent
available either as a 9? percent free-flowing powder or as a stabilized water
solution of 12 percent sodium borohydride and 40 percent sodium hydroxide.
The basic reduction reaction involves the donation of 8 electrons/molecule of
SBH to an electron deficient metal cation. The net reaction is:
NaBH4 + 4M"1"* * 2H20 = 4M° + NaBO2 + 8H*
Since one mole of sodium borohydride (SBH) can reduce four moles of
divalent metal ion (or eight moles of monovalent) , relatively low amounts of
reagent usage can result in a substantial reduction of metallic
i J i Q
contaminants. " Table 10.3.6 illustrates • theoretical usage levels and
the overall quantities of metals recovered. In practices the metal/SBH ratio
is lower since other reducible compounds (aldehydes, ketones, etc.) may react
with borohydride, 'increasing reagent consumption. Typically, SBH requirements
1°
are 1.5 to 2 times the theoretical use level.
Figure 10.3.3 illustrates a sodium borohydride treatment system,
^n triis
system, tne pH is maintained between 6 and ?, although SBH can reduce under pH
10-128
-------
TABLE 10,3,4. PERFORMANCE DATA FROM A FERRITE COPREC1PITATION
TEST FACILITY (CONCENTRATION, tnR/L)
Metal
Mercury
Cadmium
Copper
Zinc
Chromium
Nickel
Manganese
Iron
Bismuth
Lead
Influent
7.4
240
10
18
10
1,000
12
600
240
475
Effluent
0.001
0.008
0.010
0.016
0.010
0.200
0.007
0.06
0.100
0.010
Source: Reference 16,
TABLE 10,3.5, PERFORMANCE OF FERRITE COPSECIPITATION IN
OSAKA UNIT (CONCENTRATION, rag/L)
Metal
Mercury
Arsenic
Trivalent Chromium
Hexavalent Chromium
Lead
Cadmium
Iron
Zinc
Copper
Manganese
Influent
6
0.7
25
0.5
480
15
3,500
650
23
60
Effluent
0.005
0.01
0.01
not detectable
0.05
O.OL
0.04
0,5
0.08
0.5
Source; Reference 16.
10-129
-------
Gold
Silver
Cadmium
Mercury
TABLE 10.3.6. THEORETICAL SODIUM BQROHYDRIDE USE LEVELS AND
QUANTITIES OF METALS RECOVERED
Metal
Oxidation .
State
Sodium Eorohydride
Theoretical Use Levels
Metal
Recovery
Powder SWS
(g SBH/kg metal) (tnL SWS/kg metal)
(Ib metal/lb SBH
?
Copper Cu
Lead Pb2*
Nickel Ni
143
46
16?
850
270
1000,
7
22
6
Au
3+
Cd
24
16
72
62
48
430
260
370
280
14
23
12
21
Palladium PtH
Platinum Pt^*
Cobalt Co^4"
Rhodium Rh^
Iridium Ir
91
100
16?
143
100
540
600
1000
850
600
11 •
10
6
7
10
Treatment levels shown are for 97% active SBH powder and SWS", a stabilized
water solution of 12% SBH and 40% NaOH (by weight).
10-130
-------
Sodium Bisullile
20% solulion (by weighl)
VenMel solulion
Waste
Slreani
( Incoming
Waste metol
solution)
-------
conditions as low as 4.5 and as high as ,11. Sodium bisulfite is added prior
to sodium borohydride to lower the oxidation states of competing species, but
does not totally reduce the metal cations present.
In the second stage SBH solution is added. The stabilized water solution
can be handled in a similar fashion'to 50 percent sodium hydroxide (see
Section 10,1.2), It is suitable for ORP control and can be metered from a
storage tank or directly from a 55-gallon drum. Some users further dilute
the SBH solution 10:1 with deionized water prior to addition. Dilution allows
for faster mixing and helps to prevent over-dosing which may impede downstream
flocculation and settling. The solution can be used in processes with flow
rates ranging from 5 gal/min up to 1,500 gal/min and metal concentrations from
1 Q
as lo« as 2 mg/L to as high as 20,000 tng/L.
The sodium borohydride added in the second stage reacts with any residual
bisulfite from stage one to form sodium dithionate;
The sodium dithionate further reduces any oxidizing agents left in the waste
stream, partially reduces metal cations, and regenerates bisulfite which
provides a mildly reducing environment. Contact time between the SBH
solution and wastewater can be as low as 5 minutes or as long as 60 depending
on the metals concentration. The precipitated metal must be removed from the
treated wastewater quickly (in less than 1-hour) because redissolution of the
metals can occur. The sludge produced is high in metal content, finely
divided and of lower quantities than comparable technologies (68 percent less
12
than lime on a dry weight basis).
In this system a binary flocculation system is used to agglomerate the
finely divided metallic fines. A cationic polymer (a polyamine) is added in
stage 2 and an anionic polymer (bydrolyzed polyacrylamide) is added to a flash
mix tank just prior to the clarifier. This treatment method is capable of
producing a high quality effluent without filtration, and when filtered with
industrial filter media is capable of reducing many metallic contaminants to
below detection limits.
The two main disadvantages of this process are high reagent cost and the
introduction of boron to the effluent flow stream. The high cost of sodium
borohydride solution (£2.40/lb vs. S.023/lb for hydrated lime) has limited
10-132
-------
this treatment technology -to applications either low in competing reducible
species or situations which require extremely low metallic effluent
concentrations. In addition, SBH ia unable to break cyanide complexes and it
is necessary to first destroy the complex (via hypochlorite or chlorine
oxidation) prior to treatment. Sodium borate, a by-product of the SBH
reaction, introduces boron at 3 to 10 percent of the level of metals removed.
Further treatment such as filtration or ion exchange may be necessary before
discharge.
Pretreatment and Post-Treatment Requirements—
A pretreatment in itself, chemical reduction is typically applied to
chromium bearing aqueous waste streams segregated from other process
flowstreams. An exception to this is sodium borohydride reduction which is
used to reduce a wide variety o£ metallic contaminants, although on a limited
scale. The properties of the waste being treated which can affect performance
include:
* Flow variations;
* pH variations;
* Presence of chelator/complexants;
• Competing nonpriority reducible species;
* Cyanide content; and
* Oil and grease concentration.
In facilities which experience a wide variation in flow rates, pH values,
or pollutant concentrations of the wastewater, flow equalization as
pretreatment is o£ten used. A variety of process options exist (see
Section 10.1), but all systems basically provide some sort of flow resistance,
stream segregation, or influent concentration averaging to prevent waste
treatment system overloading. In all methods of flow equalization, care must
be exercised during the wastewater analysis to completely characterize any
peak flews =r ccncestretior.s. In addition, flexibility ir. system deslgr,
should be provided for any future expansion, change in location, or deviation
. c, • 20
ir, flow rstes.
10-133
-------
Oil and grease, cyani-des, ehelator/cotaplexants; and nonpriority
reducibles, are all factors which will increase reagent consumption and
impede, if not prohibit, chemical reduction operations. Oil and grease
removal is typically the first process seep in any waste treatment train. A
wide variety of treatment equipment and chemicals currently exists in both the
literature and in industry. Cyanides also effect the feasibility of chemical
reduction technologies by forming strong cyano-eomplexes or evolving toxic
hydrogen cyanide gas at the acidic conditions required for many of the
reduction technologies. Treatment of cyanide waste streams typically consists
of segregation followed by oxidation'(see Section 14). Chelator/complexants
and noopriority reducibles present a difficult problem when chemically
reducing metallic contaminants such as hexavalent chromium. Since these
compounds are often an integral part of the chromate waste stream, waste
stream segregation is difficult if not impossible. Two established methods of
pretreatment for the removal of chelator/eomplexants and nonpriority
reducibles are pH depression and binary reduction systems. In the pH
depression methods, the pH of the waste stream is lowered to approximately 2.0
through the use of acid. The low pH helps to break complexes and since it is
already a part of the overall chromium reduction process eliminates the need
for additional equipment. The second technology, binary reduction, uses a
less expensive reductant such as hydrazine, dithioearbonate, or sodiura met a—
bisulfite to "prereduce" waste streams containing excess chelators/complexes
or oxidized compounds. The prereducer acts as a scavanger while the primary
reductant works to reduce the metallic contaminants of concern.
' Chemical Reduction in itself ,does not produce any res.iduals. However, to
completely remove metallic species from the waste stream, chemical reduction
is usually followed by precipitation, coagulation/flocculation, sedimentation,
and sludge consolidation. The resulting toxic sludge must often then be
treated (i.e., encapsulation) and land disposed. For a discyssion of
post-treatment techniques see Section 10.1.
10.3.2 Process Performance
Sulfur-based chromium reduction technologies have gained wide acceptance
in industry for reducing waste stream hexavalent chromium concentrations. The
most prevalent reduction reagents are sodium sulfite salts, sulfur dioxide,
10-134
-------
and sulfuric acid. However, iron salts such as ferrous sulfate have also
shown potential for" reducing chromium wastes. As demonstrated in Table 10.3.7
sulfur-based chromium reduction technologies operate under a wide range of
influent conditions. * * . Variable flowrates (5 to 140 gpm), pH
conditions (2 to 10 standard units), and hexavalent chromium concentrations
(2.23 - 136 mg/L) were all successfully treated by the reduction technologies
examined. The most prevalent method of treatment cited among, the numerous
examples in the literature is sulfuric acid adjustment to pH 2.0 followed by
sodium bisulfite reduction and hydroxide precipitation.
The chromium reduction processes examined are very efficient in nature
with complete reduction typically achieved in less than 1 hour. In addition,
these technologies are able to successfully treat a wide range of chromate
wastes as demonstrated in Table 10.3.8.
Since most of the reduction technologies examined utilize the same
process equipment and are capable of reducing hexavalent chromium
concentrations to less than 0.01 mg/L, system selection is usually based on
economic considerations. Criteria such as reagent cost/lb Cr reduced,
excess reagent requirements, sludge generation, and pH adjustment costs will
all influence the overall economics of the system selected.
Sodium borohydride (NaBR,) has also shown promise in chromium and other
17 9 &—> R'
heavy metals reduction. ' Table 10.3.9 presents sodium borohydride
performance data for a wide variety of waste streams and metallic
contaminants. The success of sodium borohydride reduction is highly dependent
on mixing,, residence time, pH, nonpriority reducible concentrations, and
reaction kinetics. Sodium borohydride treatment at facilities A-F removed to
acceptable levels all metallic constituents of concern, enabling the facilities
to meet discharge standards. The design of the systems, as with sulfur-based
technologies, is based on standard, automatically controlled.(ORP) equipment,.
traditionally used in industrial wastewater treatment.
For facilities whose waste streams contained chelators and/or complements
(Facilities A, D, and P) pH adjustments and sodium bisulfite or ferric
chloride were required to improve SBH reduction efficiencies. All metals
except for nickel (50 percent or less removal efficiency) were removed
effectively. All facilities reported improved sludge characteristics with the
purity of the recovered metal limited only by the presence of other reducible
species.
10-135
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TABLE 10.3.7. PERFORMANCE DATA FOR SULFUR BASED REDUCTION SYSTEMS
Facility
Parameter
Wastestrean
Influent pH
Influent
flowrate (gpm)
Type of scid
Reduction pH
Reduction Reagent
Retention
Tiros (min)
Precipitation
Reagent
Precipitation pH
Influent Cr+6
A«
Nickel/
chromium
rinse
7.1
12
Sulfuric
2.5
Sodium
bisulfite
60
Caustic
soda
9.5-10.0
2.23
Bb
Simulated
chromium
rinse
6.0
35-140
Sulfuric
2.0
Sod iutn
bisulfite
10
Hydrated
lime
B.7
6.60
Ce
Chromium
rinse
2.5
10
__
6.5-7,0
Ferric
chloride/
Sodium
Eulfide
2^0
—
—
136
Dd
Chromium
rinse
2.2-3,0
90
Sul furic
2.3
Sociuni
bisulfite
Hydrazine
NA
Sode
ash
7.0-8.5
8-21
£fi
Simulated
chromium
rinse
7.0
5
__
7.0-10.0
Ferrous
suf late/
Sodium
Sulfide
In-line
mixing
CauSCic
soda
7. 0-10. D
5-50
F*
Chroniiytn
rir.se
. N*
70-87
Sulfuric
2.0
Sulfur
d ioxide
NA
Lime
NA
m
Concentration (ug/L)
Effluent Cr*6
Concentration (mg/L)
Effluent pH
NA = not available
^Source: Reference 21.
"Source: Reference 10,
C5ource." Reference 22.
^Source; Reference 23.
eSource: Reference 15.
'•Source: Reference 24.
O.D1
0.01
5.7
0.01
6.5-7.0
7.1
?. 0-9.0
0.05
6.0-8.5
10-136
-------
TABLE 10.3.8. HAZARDOUS WASTES TREATED BY CHROMATE REDUCTION
Waste code
Soluble • U032
DOO?
F019
Insoluble K002
K003
K004
K005
K006
K008
K086
F006
Description
Calcium chromate
Chromium (hexavalent with
chromic acid)
Sludges from chemical conversion coatings
Production sludges from:
chromium yellow
raolybdate orange
zinc yellow
chromium green oxide
chromium green oxide
Oven residues from chromium greenoxide
Pigments and inks
Insoluble chromates
Source: Reference 25.
10-137
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TABLE 10.3.9. SODIUM BOROHYDRIDE PERFORMANCE DATA
Facility
Parameter AS C D £
Wastes cream Printed3 Mercury1* Tetraalkyl^ Lithographic*1 P;rintedc
Circuit Cell Lead Film Circuit
Boards Electrolytic Manufacture - Boards
Influent NA 10-20 900-1,500 50-210 23
Flowrate (gpm)
1st Stape " J
pH 5.5 -- — 11.0
Retention •
Time (min) 20-nO — — 15
Reagent Sodiura ; — — Ferric —
Bisulfite chloride
2nd Stage
pH 8.0 HA 9.0 1 1.0 8.0-11
Retention
Tine (nin) 20-30 15-30 15-20 30 - 30
Reagent Sodiun SBH SEH SBB SEN
Sorohvdride
CSBH)'
Influent Metals
Concentrac iors (mg/L)
Copper 20.0 — — -- 786.0
Lesd — — 5.35 — 0.5?
Nickel — -- — -- 0.06
hercury — 10-50
Silver ' — ~ — 10-120 —
Cadmium -- — — * 5-60 —
Effluent Metels
Concent.rst ion Irag/L)
Copper 1.0 — — — 1.4?
Leid — — 0.1 — 0.10
Nickel ~ — — — 0.03
Zinc — — — — 0,03
Hc*"£U'**y ' 0« l~Dt 8 . : — ~~ *~~"
Silver — — — 0.05
CadEiur, — — — D.OS
aRefercnce 1?,
^Reference 26.
^Reference 27.
df;e-»T»n/.« IS
T
Gcuinerc i al
Wattetreitmenc
Plant
B'ateh
(7,900 ga
5.5
15
Sodium
Bisulfite
a.o
30-45
sen
237.0
0.32
0.96
5.10
24-0
0.12
o.u
O.i'l
0.08
0.01
u
10-138
-------
As with sulfur-based .chemical reduction technologies, sodium borohydride
implementation is influenced by overall process economics. Reagent costs are
high, and several facilities (A, B, D, and F) reported impaired economic
performance due to required pretreatments and/or excess reagent usage. All of
these factors must be considered when evaluating sodium borohydride as either
a primary or secondary treatment system for control of heavy or precious metal
waste streams.
10.3.3 Prgcess_Co_st_a
Figure LO.3.4 illustrates the basic process train developed for a
continuous ehromate reduction and precipitation system, the equipment for
.this system includes a flow equalization tank, a continuous chromate reduction
tank, a precipitation reactor, a lamella type £locculator/clarifier, sludge
holding tanks, and a plate and frame filter press. The continuous ebrooate
reduction tank includes high level alarms, portable pH and oxidation-reduction
potential CORP) meters,'a portable mixer, and storage tanks and feed pumps to
add sodium metabisulfice and sulfuric acid. Reactor, retention time is
30 minutes. The capital equipment cost data for the chromate reduction system
29
is based on Figure 10.3.5. The equipment specifications and cost
assumptions for the remaining operations are based on assumptions previously
30
presented in Section 10.1.
The influent stream to the chromium waste treatment system is assumed to
contain 200 mg/L of hexavalent chromium at pH of 6.0. Approximately 7.1 tng/L
of sulfuric acid is required to depress the chromate reduction influent stream
to a pH of 2.0, Sodium metabisulfite is added to the waste stream of a
stoichiometric rate of 15:1 and a complete reaction is assumed to occur.
Approximately 400 mg/L of hydrated lime is required to raise the pH and
precipitate the trivalent chromium as an hydroxide. This reagent addition
will result in the formation of 400 og/L of chromium hydroxide sludge on a dry
weight basis and a small quantity (1.4 mg/L) of calcium sulfate precipitate.
In reality the quantity of sludge produced will be very much a function of the
caicium added as hydrated lime and the quantity of sulfates present in the
waste stream. In this hypothetical model the only sulfates present are those
that were introduced by the sulfuric acid.
10-139
-------
WASTE WATER-
EQUALIZATION
TANK
CONTROL
REACTION
TANK
NEUTRALIZATION AND
CHROMIC HYDROXIDE
PRECIPITATION
•SODIUM METABISULFITE
SULFUFiie ACID
•HYDR4TED LIME
CLARIFIER
UNDERFLOW
FJLTEn
TREATE:
SLUDGE
S.QUEDUE
PHASE
EFPLUEK'T
ENCAPSULA7
ION
SOLIDS TO
LAND DISPOSE
Figure 10.3.4. Chromate reduction system.
LO-140
-------
50 f~
20 30 SO
FLOW RATE SgaUmin)
50
60
leganti:
«p*^^— Total insialiedC
" —- Hardware cost
Notes:
Bslch units:
!r.s;a!:ei car. = 2 K hardware cos;.
Unit consists of TWO ^-hoyr rgaclion lanks
wsin necsssary auxiliaries.
RLgure 10.3.5, Investment cost for chromium reduction units,
Source: Reference 29.
10-141
-------
Table 10.3.10 presents the annual costs for the continuous chrornate
reduction system illustrated in Figure 10.3,4. Flow rates are 1,000, 10,000,
and 100,000 gal/hour. As with all waste treatment systems which rely on
chemical precipitation as the primary method of contaminant removal, sludge
disposal costs constitute a large percentage of the total annual costs. In
addition the high cost of treatment chemicals may prohibit the use of this
technology at high influent metals concentrations. As land disposal becomes
increasingly expensive in anticipation of land disposal restrictions, waste
treatment options such as chromium reduction which generate large quantities
of potentially hazardous sludge will become less viable from both an economic
and liability standpoint.
10.3.4 Overall Process Status
Chemical reduction of hexavalent chromium through sulfur-based reagents
is a well established and fully-developed technology. Environmental
considerations result primarily froa residuals generated in the
precipitation-sedimentation process following chromium reduction. In
processes which use ferrous sulfate as the reducing-agent, sludge generation
can be significant. In. addition, a potential hazard in reagent storage and
handling is present for those facilities using gaseous sulfur dioxide.
Table 10.3.11 contains a summary of the advantages and disadvantages of
hexavalent chromium reduction.
Sodium borohydride, which has been applied on a limited basis as an
alternative chemical reduction process in some chloralk.ali and metal finishing
facilities, has a potential as a viable waste treatment 'option. Sludge
production is less than comparable technologies and with the exception of
nickel, metal removal efficiencies are sufficient to meet effluent limitation
guidelines. The main limitations to this technology are high reagent costs,
the in'troduction of boron into the effluent waste stream, and the evolution of
hydrogen gas as part of the reduction process. None-the-less, as land
disposal costs continue to increase, sodium borohydride's ability to produce a
-compact, high density, pure sludge product will enhance its selection as an
alternate metals reduction process.
10-142
-------
TABLE 10,3.10. ANNUAL COSTS FOR A CONTINUOUS CHROMATE REDUCTION SYSTEM8
Purchased Equipment and Installation (PE&I)
Equalization Tank
Reduction Tank
Precipitation Reactor
Flocculator/Clarif ier
Sludge Holding Tankts)
Filter Press
Total Capital Investment (360? PE&I)
Annual Operating Costs ($/Yr.)
Operating Labor ($2Q/hr.)
Maintenance' (61 TCI)
General Plant Overhead (5.8% TCI)
Utilities (2% TCI)
Taxes and Insurance ( 1% TCI)
Chemical Costs:
Lime ($40/ton)
Sulfuric Acid ($72/ton)
Sodium Metsbisulfite ($32/1001b)
Sludge Transportation (SO. 25 / ton-mile)
Sludge Disposal ($200/ton)
AnnualiEed Capital (CFR-0.177)
Total Cost/year
Cost/1000 gallon
1000
17,000
15,000
24,000
18,000
3,000
10,000
87,00
313,200
72,000
18,800
18,200
6,300
3,100
200
2,200
1,500
200
9,600
55,400
187,500
26
Flow rate Cgph)
10,000
29,000
66,000
40,000
50,000
6,000
25,000
216,000
777,600
72,000
46,700
45,100
15,600
7,800
1,800
21,600
15,200
1,800
96,400
137,600
461,600
6
100,000
50,000
114,000
69,000
140,000
48,000
100,000
521,000
1,875,600
72,000
112,500
108,800
37,500
18,800
18,500
216,100
152,200
1,800
963,800
332,000
2,034,000
3
S1987 Dollars.
10-143
-------
TABLE 10.3.11. ADVANTAGES AND DISADVANTAGES OF CHEMICAL REDUCTION
OF HEXAVALENT CHROMIUM
Advantages
- Well proven technology with documented reduction
efficiencies.
- Operates at ambient temperature and pressure lowering
energy requirements.
Process equipment is modular and widely available
from a variety of manufacturers and suppliers.
Is applicable to a wide range of chromium wastewaters
from numerous industrial sources.
Disadvantages
For .high concentrations of influent chromium, the high
cost of treatment chemicals may be prohibitive.
Chemical interference by oxidizing agents in nixed
waste streams may add substantially to reagent
requirements.
- Sludge production is relatively high and in the case
of ferrous sulfite can be significant.
Storage and handling of gaseous sulfur dioxide is
somewhat hazardous.
10-144
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REFERENCES
1. U.S. EPA. Treatability Manual. Vol. III. Washington, DC.
EPA-600/8-80-Q42c,d. July i960.
2. Sundstrom, D.W,, and H.E. Kiel. Wastewater Treatment. Prentice-Hall,
Inc., Englewood Cliffs, NJ. 1979.
3. Stumm, W., and J.J, Morgan. Aquatic Chemistry, John Wiley & Sons,
New York, NY. 1970.
4. M.A. Chillingworth, et al., Alliance Technologies Corporation.
Industrial Waste Management Alternatives Assessment for the State of
Illinois, February 1981.
5. R.E, Wing. Complexed and Chelated Copper-Containing Rinsewaters.
Plating and Surface Finishing. -July 1986.
6. G.C. Cushnie, Centec Corporation. Navy Electroplating Pollution Control
Technology Assessment Manual. Naval Civil Engineering Laboratory
CR 84.019. February 1984.
7. K.H. Lanouette. -Heavy hetals Removed, Chemical Engineering.
October 17, ,1977.
8. Kirk-Othmer Encyclopedia of Chemical Technology, Vol. 22, 3rd, Edition,
John Wiley & Sons, New York, NY. 1981,
9. U.S. EPA. Proposed Development Document for Metal Finishing.
E?A-44Q/l-82-091b. 1982.
10. Taylor, C.R., and S.R. Quasim. More Economical Treatment of Chrotnium-
Bearing Wastewaters. Proceedings from the 37th, Industrial Waste
Conference, Purdue University. 1982.
11. Kirk-Othmer Encyclopedia of Chemical Technology, Vol. 10, 3rd, Edition,
John Wiley & Sons, New York, NY. 1981.
12. J.R. Aldrich, et al. Hazardous Sludge Reduction. 70th. AES Annual
Technical Conference proceedings, Indianapolis, IN. June 1983.
13. Higgins, I.E., and V.E. Sater. Combined Removal of Cr, Cd, and Ni from
Wastes. Environmental Progress. March 1984.
14. Higgins, I.E., and S.G. TerMaath. Treatment of Toxic Metal Wastewaters
by Alkaline Ferrous Sulfate and Sodium Sulfide for Chromium Reductions,
Precipitation, and Coagulation. 36th. Industrial Waste Conference,
Purdue University. 1982.
15. Higgins, I.E., and B.R. Marshall. Combined Treatment of Hexavalent
Cnromium with Other Heavy Metals at Alkaline pH, 17th. Mid-Atlantic
Industrial Waste Conference. June 1985.
10-145
-------
16. U.S. EPA, Sources and Treatment of Wastewater in the Nonferrous Metals
Industry. EPA-6QO/2--8Q-074. April 1980.
17, Lindsay, M.J., and M.E. Hack-man, Morton Thiokol, Inc. Sodium Borohydride
"Reduces Hazardous Waste. Purdue Research Foundation, West Lafayette,
IN. 1985.
18. J.A. Ulman, Control of Heavy Metal Discharge in the Printed Circuit
•'Board Industry with Sodium Borohydride. AES SUR/FIN Annual Technical
Conference and Exhibit. 1984.
19. Ven Met Brochure, Morton Thiokol, Inc., Ventron Division. 1984.
20. MITRE Corporation. Manual of Practice for Wastewater Neutralization-and
Precipitation. EPA-600/2-81-148. August 1981.
21. Rabosky, J.G., and T. Altares. Wastewater Treatment for a Small Chromium
Plating Shop: A Case History. 38th. Industrial Waste Conference, Purdue
University. 1983.
22. R.S. Talbot. Co-Precipitation of Heavy Metals with Soluble Sulfides
Using Statistics for Process Control. 16th. Mid-Atlantic Industrial
Waste Conference. 1984.
23.- J.J. Martin. Chemical Treatment of Plating Waste for Removal of Heavy
Metals. U.S. EPA-R2-73-044. May 1973.
24. Anonymous, Cleaning Up an Industrial Discharge. Environmental Science
and Technology. August 1973.
25. D.W. Grosse. A Review of Alternate Treatment Processes for Metal-Bearing
Hazardous Waste Streams. Journal of the Air Pollution Control
Association. May 1986.
26. M.M. Cook, et al., Morthon Thiokol, Inc. Case Histories: Reviewing the
Use of Sodium Borohydride for Control of Heavy Metal Discharge in
Industrial Wastewaters, 34th. Industrial Waste Conference, Purdue
University. 1979.
27. Palmer, S-, and T.J. Nunno. Case Studies of Existing Treatment Applied
to Hazardous Waste Banned from. Landfill: Facility B Draft Final Report.
Contract No. 68-03-3243.
28. Palmer, S., and T.J. Nunno. Case Studies of Existing Treatment Applied
to Hazardous Waste Banned from Landfill: Facility A Draft Final Report.
Contract No. 68-03-3243.
30. Versar Inc. Technical Assessment of Treatment Alternatives for Wastes-
Containing Metals and/or Cyanides. Contract No. 68-03-3149, U.S.
EPA/OSW. October 1984.
10-146
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10,4 FLOTATION
Low density suspended material may often be separated from a liquid
matrix by flotation. In this operation, fine air bubbles introduced into the
solution result in the attachment of the bubbles to the particles. The
attached bubbles cause the particles to rise to the liquid surface, where they
are removed by skimming. This process is referred to as dissolved or
4 5
dispersed-air/foam flotation. *
For materials that are dissolved and not suspended, other steps are
needed to precipitate the contaminant prior to flotation. For example, a
surfactant such as carboxylic acid, can be added. This process is referred to
1^6
as ion flotation. ! ' The "collector" reacts with the dissolved material
to form an insoluble product and facilitates the attachment of bubbles to the
particle surface. Improved cost-effectiveness can be achieved if the ion is
precipitated first, and then floated with a subsequently smaller quantity of
collector required. This is called precipitate flotation and is ideally
carried out in a flotation column. Another technology, adsorbing colloid
flotation, removes dissolved materials by adsorbing them onto colloidal
particles which are then removed by flotation.
Since dissolved air flotation by itself is incapable of removing
dissolved metallic, contaminants, recent research efforts have centered on the
technologies of ion, precipitate, and adsorbing colloid flotation. Therefore,
the remainder of this section will be devoted to these technologies.
10.4.1 Process Description
The principal components of an air flotation system are a pressurizing
pump, air supply, retention tank, and flotation unit, as shown in
Figure 10.4.1. The system may also be operated with recycle as shown in
Figure 10.4,2. In the recycle system, a portion of the clarified effluent is
2
contacted with the dissolved air and pumped to the retention tank. The
aerated recycle stream is then mixed with fresh sludge at the entrance of the
flotation unit. A recycle system avoids the shearing action of the
pressurising pump on the influent waste which impairs performance due to floe
break-up. Table 10.4.1 sutnmarizes typical operating parameters for this type
of system.
10-147
-------
Effluent
liquid
Halation
unit
Figure 10.4.1. Schematic diagram of dissolved air flotation syster
without recycle.
Figure 10.4.2, Schematic diagram of dissolved air flotation system
with recycle.
oouree: Keierence £.
10-148
-------
TABLE 10.4.1. TYPICAL OPERATING PARAMETERS FOR DISSOLVED AIR FLOTATION UNITS
Parameter
Unit
Range
Air Pressure
Air-to-Solids Ratio
Recycle Ratio
Overflow Rate
Solids Loading
Detection Time
ps la
mass/mass
o/o
gpd/ft2
Ib/day ft2
45 - 95
1:10
20 - 150
700 - 2,500
25 - 100
20 - 60
Source: Reference 1, 2.
10-149
-------
Central to any flotation technology, regardless of the separation
mechanism used (i.e., ionic bonding, precipitation, adsorbtion), is the
addition of a surfactant. Table 10,4,2 lists several commonly used flotation
surfactants.
There are currently two physical models for the description of Che
attachment of contaminant particles to air-water interfaces in the presence of
7 g
surfactant. ' In the columbie model, ionic surfactant is adsorbed on the
air-water interface resulting in a surface charge- density on the bubbles
.(usually negative, because of the use of an anionic surfactant). The
dissolved ion, precipitate, or absorbing floe is then given a surface charge
by adjusting the pH or the concentration of other potential-determining ions.
The adsorption mechanism involved is the electrostatic attraction between the
ionic surfactant (collector) and the metallic contaminant of opposite charge.
In the contact angle model, surfactant ions adsorb to the primary layer
of the metallic particle, presenting the ionic head of the surfactant to the
solid and the long hydrophilic, hydrocarbon tail to the solution. The
interfacial free energies (surface tensions) are not; such that the contact
angle of the air-water interface on the metallic particle is- different from
zero which permits the attachment of the particle to a bubble.
o
Some basic conclusions common to both models are:
» Increasing ionic strength tends to decrease flotation efficiency.
• Increasing the length of the surfactant hydrocarbon1tail decreases
the bulk liquid concentration (moles/liter) of surfactant required
to produce flotation.
» Increasing particle size increases flotation efficiency.
» Increasing temperature increases the concentration of surfactant
required.
Recent advances in flotation equipment design have involved substituting
vertical columns for the open, agitated vessels commonly used in conventional
snt d«
8-12
flotation systems. Figure 10.4,3 illustrates a pilot plant developed by
Thackston et. al. at Vanderbilt University, Nashville, TN.
In this system, simulated wastewater is pumped from the storage tank
through the pilot plant at the desired flow rate. The addition of the
coagulant and NaOE, for pH control, occurs upstream of the nain pump. The
10-150
-------
TABLE 10.4,2. TYPICAL FLOTATION SURFACTANTS
Type
Formula*
Charge on the soluble ion
Sulfhydryl collectors
xantbate
dithiophosphate
atone chiocarbamate
thiol Cmercaptan)
dixanthogen
thiocarbanilide.
Colloidal electrolytesc
fatty acids and their soaps
alkyl or aryl alkyl sulfonates
alkyl sulfate
primary amine salt
secondary araine salt
quaternary ammonium salt
ROCSSNa
RHNCSOR
RSR
(ROCSS)2
(C6H5NH)2CS
RCOOE, RCOONa
ROSOjSa
RNH3C1
R2NH2C1
atiionic
anionic
aniontc
anionic
anoinic
cat ionic
cat ionic
cationic
aR = CH3tCH2)n
bFor sulfides, R » C2 - €5.
cGenerally straight chain Cj_2 to Cj_g, or a benzene or naphthalene ring
may be incorporated into the R group.
Source: Reference 3.
10-151
-------
WASTE
TANK
NaOH
TANK
I
19 STA6E
BAFfLED
COLUMN
!
•
CLARIFIER
BRO
KEN
FOAM
CONTAINER
EFFLUENT
:Flgure 10^4.3. Schematic diagram of adsorbing foam flotation pilot plant.
10-152
-------
-------
coagulant is fed at the required rate by a variable feed.pump. The NaOH
solution flows by gravity through a solenoid valve connected to a pH
controller set to produce the desired pH in the first mixing-flocculating
chamber. After passing through the main pump, the wastewater enters a series
of three mixing-flocculating chambers, after which NLS (sodium lauryl sulfate,
an .anionic surfactant) is injected at the required rate. The «aste is then
sent through the top of the flotation column to a spray nozzle. The
uastewater then flows downward through the rising foam over an arrangement of
19 baffles installed to prevent foam overturning. The sir is supplied through
a fitted glass disk in the bottom of the column. The treated effluent leaves
the column through the bottom and the foam is piped out of the top of the
column to a rotating disk foam breaker. The effluent pH is monitored
continuously.
Table 10.4.3 contains operating parameters for three colloidal adsorption
$ 11
foam flotation pilot plants. ' All three studies were continuous flow
applications which focused on lead removal from synthetic and industrial waste
streams. Subsequent case studies have focused on remcvir.g Cu, Cd, hg, Zn, Cr,
and arsenate at both the bench and pilot scale.
Pretreatment requirements reported in the literature for chemical
flotation include:
• Flow equalization;
• pH adjustment;
* Coagulation (adsorbing colloid foam flotation); and
• Precipitation (precipitate flotation}.
Some method of flow equalisation (see Section 10.1) should be provided to
average waste stream influent concentrations to prevent system overloading and
maintain optimum performance characteristics. Other pretreatonents such as pH
adjustment (see Section 10.1), coagulation (see Section 10.3), and hydroxide
precipitation (see Section 10.1) are all technology specific and have been
extensively reported in the literature. Typical requirements would be a
mixer-reactor £or pn adjustment (equipped wicn appropriate reagent tanks and
controls) followed by a precipitation or coagulation reactor depending on
which'flotation technology is utilized.
10-153
-------
-------
TABLE 10.4.3. OPTIMUM OPERATING PARAMETERS
Parameter, units From Hanson From Miller From Thackston
pH 5,5-6.5 6,0-7,0 6.9-7.1
FeClII), mg/L 150 100-150 90
NLS, mg/L 35-40 34-40 25
Hydraulic loading
gal/min/gq ft
Air supply
cu ft^min/sq ft
Effluent Lead
Concentrations (mg/L)
118-176
2-3
0.2-0.3
0.66-0.98
0.1
148-178
2.5-3.0
0.4-0.5
1.31-1.64
0.4
326
5.5
0,2
0.7
0.1
Source: Adapted from References 8, 11.
10-154
-------
-------
Oil and grease, cyanides, chelator/complexants, and.competing ions for
surfactant sites (i.e., carbonates) have not been considered in literature in
any great detail. There is a real research need for work to determine whether
or not these factors will make the waste acceptable for chemical flotation
treatment. Mosc of the research to date has focused on low concentration
(1,000 nsg/L or less) simulated waste streams prepared by dissolving four or
less metallic salts (alone or in combination) in tap water. If chemical
flotation is to achieve wid'e acceptance in industry, more realistic (and
consequently difficult to separate) waste streams will have to be investigated,
Residuals produced by chemical flotation consist primarily of a metals
laden foam layer which is skimmed or drawn off the top of the reaction
vessel/column. Post-treatment typically consists of sedimentation and sludge
consolidation. The resulting hazardous sludge must often be treated
(i.e., encapsulation) and then land disposed. In addition, air stripping of
the foatnace to recover surfactant may be desirable to reduce surfactant
consumption (approximately 60 to 70 percent).
10,4.2 Process Performance
While most current research has focused on precipitate and absorbing
colloid foam flotation techniques, Eastern European research has also
encompassed the technology of ion flotation of dissolved metallic
contaminants. For example, Skyleu et al. at I. I. Mechnivkov State
University, Odessa, investigated the removal colloidal suspensions of metallic
mercury (25 to 50 ttg/L) using flotation apparatus. The collectors were
0.01 to 0.1 percent aqueous solutions of potassium salts of pentadecoanoic,
palmitic, heptadeconoic, and stearic acids. The best collector for mercury, in
ionic flotation treatment of these solutions was found .to be potassium'Stear-
ate, giving 98 percent extraction of a simulated waste stream containing 20 to
50 mg/L Hp and 10,000 to 15,000 mg/L NaCl and 78 percent extraction from a
simulated waste stream containing 50 mg/L Bg, 1,000 rag/L NaCl, 3,000 me.il NaOH,
500 mg/L Na2CO 10 mg/L CA*2, and 500 mg/L S0~ . However, as with
most ionic flotation technologies, collector consumption was greater than
stoictiiometric requirements. Actual collector consumption'reported by Skyleu
et al. was 1,9 moles of potassium stearate/mole of mercury removed. The
process was carried out at 25°C and required 10 to 12 minutes for completion. •
10-155
-------
-------
In the United States, heavy metal removal through precipitate and
absorbing colloid foam flotation treatment of industrial waste streams has
received the most attention. In the early 197Q'a, Zeitlin's group at the
University of Hawaii demonstrated the effectiveness of absorbing colloid
flotation for removing zinc, copper, phosphate, and arsenate from seawater
14 15
using dodeeyl sulfate as a surfactant. ' In the mid 1970's to the
present, Wilson'a group at Vanderbilt University investigated and refined
Zeitlin's method using sodium lauryl sulfate (NLS) as a surfactant with iron
16-19
and aluminum as coagulants.
Using the apparatus shown in Figure 10.4.3, Wilson and Thackston, et al.
investigated lead, copper, zinc, trivalent and hexavalent chromium removal
using absorbing colloid flotation on actual and simulated industrial
wastewaters. Table 10.4.4 describes optimum operating conditions as
determined by pilot plant data for copper and zinc removal. Tables 10.4.5 and
10.4.6 illustrate foam flotation results for wastes containing a binary
mixture of copper and zinc and a tertiary mixture of copper, divalent zinc,
and trivalent chromium, respectively (influent metals concentrations 20 mg/L
each). Tables 10,4.7 and 10.4.8 contain pilot plant data on hexavalent
chromium and lead removal.
According to the researchers, for mixtures of metals, the optimum pH for
metals removed is displaced to a higher value than those obtained for each
individual metal. For example in the copper—zinc system, residual copper
concentrations below 0.1 mg/L were obtained in pH range of almost one unit
when using chemical doses of 100 mg/L Fe(III), 100 mg/L AL(III), and 70 mg/L
NLS. Even when the chemical doses are reduced to 50 to 75 mg/L of Fe(III) and
AHlII), and 40 to 50 mg/L of NLS, residual copper concentrations
substantially below 0.2 mg/L were consistently achieved in the same pH range.
Residual zinc concentrations below 1.0 rag/L were obtained in a pH range of 7.4
to 8, and values close to 0.5 mg/L were obtained even at the low coagulant-
adsorbent and NLS dose concentrations indicated above for copper. When only
Fe(IIl) was used as a coagulant-adsorbent, and no Al(III) was used, poor zinc
removal was obtained. The presence of AH OH)., exhibits some complementary
beneficial effect on copper removal, and Fe(OH)~ improves zinc removal,
although the main effect is on copper.
10-156
-------
-------
TABLE 10.4.4. .OPTIMUM OPERATING CONDITIONS AS DETERMINED IN THE
29- BY 244 CM PILOT PLANT
Parameter
Copper' removal
Zinc removal
Effluent pH
6.9-7.:
7.5-7.8
Coagulant - adsorbent4
90-100 rag/L
Fe(IIi)
100 mg/L
AHI1I)
NLSC
15-20 mg/L
30-40 mg/L
Hydraulic loading
7-14 o3/m2/hr
(3-6 gal/£t2/hr)
7-14 m3/Bi2/hr
(3-6 gal/ft2/tnin)
Air flow rate
12-14 N iD3/m2/hr 12-14 N m3/iD2/hr
(40-45 ft3/ft2/hr) (40-45 ft3/ft2/hrj
afhese values correspond to an initial metal concentration of 20 mg/L.
**These values refer to a floe concentration corresponding to the Fe(IlI) and
Al(lII) doses given in the table.
Source: Reference 10,
10-157-
-------
-------
TABLE 10.4.5. FOAM FLOTATION OF WASTES CONTAINING Cu(I)+ Zn(II)a
pH
7.0
7.2
7.2
7.4
7.4
7.4
7.4
7.5
7.5
7.5
7.7
7.7
8.0
FeUlI)
(mg/L)
100
100
100
100
150
75
100
50
75
100
50
100
50
Al(III)
(mg/L)
100
100
100
0
0
75
100
50
75
100
50
100
50
NLS
(mg/L)
70
70
50
35
45
50
70
50
50
70
40
70
40
Residual
Copper
(mg/L)
0.09
0.09
0.15
0.33
0.29
0.10
0.04
0.13
0.17
0.03
0.11
0.03
0.10
Residual
Zinc
(mg/L)
2.3
1.1
1.2
9.6
6.6
1.2
0.8
1.2
0.9
0.4
0.9
0.6
0.7
aAll runs with initial copper and zinc at 20 mg/L each, influent flow rate
at 6.9 m3/m2/hr, and air flow rate at 14 N m3/m2/hr.
Source: Reference 10.
10-158
-------
-------
TABLE 10.4.6. FOAM FLOTATION OF HASTES CONTAINING Cu(II)+ Zn(H}+
PH
7.0
7.2
7.2
7,2
7.3
7.3
7.3
7.3
7.3
7.4
7.4
7.6
Fe(III)
(mg/L)
100
100
115
150
75
100
115
. 115
" 150
115
150
100
' AK11I)
(mg/L)
100
100.
75
100
75
100
75
75
100
75
100
100 .
NLS
(mg/L)
70
70
50
85
50
70
70
50
85
70
85
70
Residual
Copper
(m'a/L)
0.10
0.07
0,12
0.07
0,42
0.08
0.08
0.14
0.06
0.07
0.11
0.22
Residual
Zinc
{mg/L)
2.00
0.90
1.10
0.60
2.40
0.56
0.56
0.75
0.40
0.16
0.30
0.20
Residual
Chromium
(mg/L)
0.12
0.13
0.13
0.12
0.60
0.20
0.20
0.33
0.12
0.20
0.21
0.26
aAll runs made vith initial copper, zincj and chromium at 20 mg/L each,
influent flow rate at 6.9 m^/m^/hr, and air flow rate at 14 R m^/m^/hr.
Source: Reference 10.
10-159
-------
-------
TABLE 10.4.7. INFLUENCE OF Fe(II) DOSE AND pH ON CHROMIUM AND IRON REMOVAL3
pH controller
Lower set point
6.0 .
6.0
6.5
5.5
6.5
6.5
6.5 •
Steady-state
effluent pH
5.
5.
6.
5.
6.
6.
6.
2
2
0
9
0
0
1
Fe( II) dose
Cmg/L)
70
64
70
64
60.5
57.6 ;
51.2
Effluent Cr
Ug/L)
0.
0.
0,
0.
0.
0.
1.
25
26
17
22
25
25
10
Effluent Fe
Cmg/L)
12
7
14
. 3
2
2
1
aOperating conditions; initial Cr(VI)=20, NLS-40 mg/L, H.L.=0-45 m3/h
(2 gal/min), H.L.R. = 6.S m3/m2 h (2.8 gaL/rnin ft*), air Elow
rate = 21.5 N m3/m2 (50 SCFH).
Source: Reference 9.
10-160
-------
TABLE 10.4.8. EFFLUENT Pblll) CONCENTRATION VERSUS pH
Pb(ll) • .
pH concentration (mg/L)
• .5.5
5.6
• 5.8
6.0
6.4
6.5
6.6
6. ?
6.8
6.9
7.1 ' '
7.2
>4.0
4.0 - . '
3.4
3.1
2.0 . -
1.3 *
1.0
0.30
0.20
0.10 '
-------
In the copper-zinc-chromium system good results were obtained over a
reasonably wide pH range at 100 to 115 mg/L Fe(III) and 75 to 100 mg/L
Alt III), and adjusting the NLS dose as a function of the total floe
concentration in the system. Zinc removal was more effective at pH values
higher than 7.3 in the range studied. Copper and chromium were better removed
at pH values between 7 and 7.3, although very good resales were achieved
throughout the experimental range. Similar low effluent concentrations were
obtained for the hexavalent chromium (0,22 mg/L) and lead (<0.1 mg/L) single
component syste'ms when iron was used as the coagulant. However, Wilson
et al., did not determine the effect of greater than 100 cng/L influent metal
concentrations, metallic precipitates, and chelation agents, on filtration
removal efficiencies.
Bench-scale experimentation was performed by Brooks et al, in 1984 to
examine the effect of these factors on the potential for separation and
concentration of strategic metals, such as chromium, copper, zinc, and
nickel. Specific waste systems selected for the evaluation tests were
electrochemical machining solids from high nickel-alloy processing and
electroplating wastes, as well as brass industry pickling waste sludges.
The experiments which simulated waste metal hydroxides using alkali
precipitation from 'the salts o'f-''the individual metals (see Table 10,4.9)
indicate that without coagulants (i.e., iron or alum), only nickel provided
any measure of efficient flotation. Similarly, when chelation agents were
used in conjunction with NLS (see Table 10.4.10) only the nickel-dimethyl
glyoxime system obtained a high selectivity in flotation separation.
Clearly, while flotation techno-logies offer promise as an efficient
method for removal of low concentrations (<100 mg/L) of soluble metals in
wastewaters, further research remains to be performed.
10.4.3 Process Costs
Realistic costs for the foam flotation process can not be developed at
this time due to the lack of commercial-scale testing. It is expected that
the primary cost would be for the flotation column. Operating costs would be
expected to hs eqiiivsler.t ••its that of lime precipitation, and .savings would
be realized in reduced disposal costs and reduced purchase costs for recovered
. . 10.12
cnemicais. • ~ '
10-162
-------
TABLE 10.4.9. FLOTATION OF METAL HYDROXIDES WITH SODIUM LAURYL SULFATE
(NLS) at 200 ppm " '
Metal as
hydroxide
Nickel
Zinc
Copper
Iron
Chromiun
Metal . NLS/Metal
concentration (pptn) (Wt/Ratio)
1,000
1,000
1,000 :
200
200
0,2
0.2
0.2
1.0
1.0
pH ,
9.
8.
9.
9.
9.
5
0
5
5
5-
FlotatioD
Performance
Good
Partial
Partial
Poor
Poor
TABLE 10.4.10. METAL HYDROXIDE FLOTATION WITH CHELATION AGENTS COMBINED
WITH SODIUM LAURYL SULFATE (200 ppm)
Hydroxide
{ 1,000 ppm
metal)
Nickel
Zinc
Copper
Iron
Chromium
Chromium
Chelation Agent
Dimethyl Glyoxirae
Zincon
Neocuproine
Batbophenanthroline
Diphenylcarbazide
Aliquat 336
Weight n
Surfactant
solids
JO. 2 '
0.2
0.2
1.0
0.8
1.0 '
itios
Chelate
solids
0.3
0.36
0.4 .
2.5
1.7
1.0
pH
8
8
8
9.5 '
9.5
9
Flotation
performance
Good
Partial
Partial
Poor
Poor
Poor
Source: Reference 20.
10-163
-------
Preliminary cost estimates for adsorbing foam flotation metal treatment
systems have been prepared by the developer (Wilson, et al.) based on
pilot-scaling testing. Capital costs (1983) for a 50,000 gpd plant, creating
60 mg/L of heavy metal was hypothesized to be approximately $20,000. However,,
until a full-scale cotmnerical unit is actually in place, the economic
feasibility of adsorbing foam flotation has yet to be adequately determined.
10.4.4 Process Status
Chemical flotation, while currently at the -bench and pilot-scale level of
developments shows promise for reducing low concentrations (.— 100/mg/L) of
effluent metals to acceptable levels. The process operates at ambient
ternperature and pressure and is well suited to automatic and computer
control. Its ability to treat both singular and1 mixed metals waste 'streams
has been well demonstrated by researchers at the University of .Hawaii and
Vanderbilt University. -However, further research is necessary before this
technology is applicable on a wider scale.
Environmental impacts result primarily from the production of potentially
hazardous sludge. The sludge product is generated during the foam breaking
process in the supernatant clarifier. At this point in time, little
information is available in the literature on the quality and composition of
the sludge product produced. Research dissertations outlining experimental
results have focused primarily on metals removal. However, foam flotation is
reported to generate less sludge than comparable precipitation processes,
approximatley 2 to 3 percent of the influent volune. The demonstrated
performance, and possible lower sludge generation rates for this waste
treatment technology warrant further research efforts.
10-164
-------
REFERENCES
1. U.S. EPA. Treatability Manual, Vol. III. EPA-6GQ/18-80-042.' July 1980.
2. Sundstrom, D.W., and H.E. Kiel. Wastewater Treatment. Prentice-Hall,
Inc., Englewood Cliffs, m. 1979.
3. Kirk-Gtbmer Encyclopedia of Chemical Technology, Vol. 10, 3rd. Edition,
Joho Wiley & Sons, New York, NY. 1981.
4. G. Parkinson. Improved Flotation Routes Get Separation Tryouts.
Chemical Engineering. March 31, 1986.
5. T.D. Reynolds.1 Unit Operations and Processes in Environmental
Engineering. PWS Publishers, Boston, MA. 1982,. ,
6. 8..E. Klimpel, Dow Chemical Company, Use of Chemical Reagents in
Flotation. . Chemical Engineering. September 3, 1984.
7. Rubin, A.J., and W.L. Lapp. Foam Separation of Lead (11) with Sodium
Lauryl Sulfate. Analytical Chemistry. July 1969,
8. E.L. Thackston, et al. Lead Removal with Adsorbing Colloid Flotation.
Journal of the Water Pollution Control Federation. February 1980.
9. Huang, S., and D.J. Wilson. Hexavalent Chromium Removal in a Foam
Flotation pilot Plant. Separation Science and Technology,•19, 1984.
10, G. Mclntyre, et al. Inexpensive Heavy Metal Removal by Foam Flotation.
36th. Industrial Waste Conference, Purdue University. 1982.
11. Slapik, M.A,, Thackston, E.L., and D.J. Wilson. Improvements in Foatn
Flotation for Lead Removal. Journal of the Water Pollution Control
Federation. March 1982.
12. G. Mclntyre, et al. Inexpensive Heavy Metal Removal by Foam Flotation.
Journal of the Water Pollution Control Federation. September 1983.
13. L.D. Skrylev, et al. Flotation Separation of Colloidally Dissolved
Metallic Mercury Collected with the Aid of- Potassium Salts of Saturated
Fatty Acids. Plenum Publishing Corporation. 1985.
14. Kim, Y.S., and M. Zeitlin. The Separation of Zinc and Copper From
Seawater by Adsorption Colloid Flotation. Separation Science 1. 1972.
15. Chainc, F.E., and M. Zeitlin, The Separation of Phosphate and Arsenate
from Seawater by Adsorption Colloid Flotation. Separation Science 9(1).
1974.
10-165
-------
16. Ferguson, B.B., Hinkle, C.» and D.J. Wilson. Foam Separation of
Lead (11) and Cadmium (11) from Wasteuater. Separation Science 9(2),
1974.
17. Wilson, J.W., and D.J. Wilson, Electrical Aspects of Adsorbing Colloid
Flotation. Separation Science 9(5). 1974.
18. Robertson, R.P., Wilson, D.J., and C.S. Wilson. The Adsorbing Colloid
Flotation of Lead (11) and Zinc (11) by Hydroxides. Separation
Science 11(6). 1976.
19. Currin, B.L., Potter, F.S., and D.J, Wilson. Surfactant Recovery in
Adsorbing Colloid Flotation. Separation Science 13(4). 1978.
20. C.S. Brooks. Recycle Metals, Metal Recovery From Waste Sludges. 39th.
Industrial Waste Conference, Purdue University. 1984.
10-166
-------
-------
SECTION 11.0
BIOLOGICAL TREATMENT FOR METAL-CONTAINING WASTES
Biological treatment is a separation process rather than a destruction
technology for metal-containing wastes. Biological separation mechanisms
include sulfide precipitation, adsorption, and biofloeculation. The types of
biological treatment technology vary considerably.' Those that are considered
in this section are activated sludge, anaerobic digestion, and algal
treatment. Common parameters for the design and operation of these
technologies are outlined and information on recently developed biological
organisms is presented.
High concentrations of heavy metals are toxic to most microorganisms and
often cause serious upsets in biological systems. Thus, influent heavy metal
concentration which can be tolerated and removed is the major criterion on
which these technologies'are evaluated in"this' section." In addition, factors
such as type of influent, its strength, and the extent of system acclimation
are also used to evaluate the viability of bioloieal treatment as a
technology for the removal of heavy metals from wastes.
11.1 PROCESS DESCRIPTION
As previously stated, several mechanisms can affect the removal of heavy
metals during biological treatment include sulfide precipitation, adsorption,
and bioflocculafcion. The first mechanism, hydrogen sulfide precipitation (see
Section 10.1), is initiated by the pH dependent generation of hydrogen sulEide
by bacteria.1 Soluble metal ions react with the hydrogen sulfide and are
precipitated as insoluble metallic sulfides. The second mechanism, adsorption
of cationic metallic ions, may result from the anionic nature of certain
cellular material, clay particles, and industrial waste constituents.2
11-1
-------
Also, the organic part of. organo-netallic complexes may ie adsorbed through
the cell walls of the biological organisms, thus trapping the metals. '
The third mechanism, bioflocculation, * is related to the synthesis of
insoluble extracellular polymer strands. These extracellular polymers can act
as non-specific sorbers for metal ions.
Typically, the removal of heavy metals in a biological system and the
type of mechanism which dominates are dependent on the species of heavy metal
present (see Table 11,1.1). The distribution of a particular heavy metal
among various chemical forms, however, largely depends upon the physical and
chemical properties of the'environment established by the treatment process
itself. Upon introduction into the biological treatment system, species of
heavy metal make adjustments toward a new equilibrium state defined by
chemical environment parameters such as pH, oxidation reduction potential
(ORP), the presence of coraplexing agents, and concentrations of precipitant
stion t
10,12
D - O Q
ligands. At this point, adsorption to solid phases or biomass, and
intracellular storage can occur.
It has been found that the microbial removal of heavy aetals consists of
initial rapid uptake followed by slow, but consistent long-term uptake.
The rate 'of uptake is greatly affected by solution pH.- Sl'udge'aae, as ' '"
well as the extent of acclimation, can also affect the extent of metal removal
in an activated sludge system.
The following is a brief description of the three- main technologies used
in biological treatment of heavy metals. More details can be found in
standard texts and the references cited herein.
11.1.1 Activated Sludge
The activated sludge process uses biological populations in a completely
mixed, oxygen-rich environment to treat wastes. Dissolved"oxygen-and mixing
are provided by mechanical aerators or fine-bubble air diffusers. A settling
tank Hs then used to remove the biological floe, part of. which is mixed with
, 15
incoming waste in the aeration tank.
A variation to the activated sludge process is the use of high purity
oxygen instead of air for aerobic treatment. Oxygen can be supplied from
onsite gas generators with liquid oxygen storage as back-up. In addition to
11-2
-------
TABLE 11.1.1. POSSIBLE SPECIES OF HEAVY METALS
IN BIOLOGICAL TREATMENT
Soluble
- Ionic
- Organic complexes
- Inorganic complexes
Co-precipitate_s in metal oxides-
Precipitates
Adsorbates
- Physical
— Chemisorption
- Clay lattice
Biological residues
So'urce:' Reference 7.' '
11-3
-------
oxygen use, Che aeration tank is covered which helps to-eliminate odors and
maintain temperatures in cold-weather periods.
There are many design variations to the conventional activated sludge
process besides the use of high purity oxygen, Theae include: multiple units
with series and/or parallel flow patterns; a tapered distribution of air along
the tank length; stepwise addition of raw waste; reaeration of the recycled
"sludge before mixing with the raw influent; and extended aeration (e.g., 24
hours or longer) used for small wastewater flows.
Activated sludge systems lack stability because of the tnicrobial growth
pattern within the tank. A high rate of growth exists at the influent but
decreases along the length of the tank. This problem is greatly amplified
during flow surges. Variation of pH, temperature, and the presence' of toxic
waste constituents can all contribute to instability. Extended aeration and
the use of high purity oxygen can help eliminate the effects of shock
loadings. An extended aeration system is depicted in Figure 11.1.1.
Table 11.1.2 presents typical ranges in values for activated sludge
20 ...
system design parameters. . Additional design factors to be considered
include sludge settling and accumulation rates and air requirements. These
factors' will v-a.ry-depending ! upon5.the-type^ of;, was.t-e- to- be.handled.- Other i >.•••.
factors which may limit the viability of the activated sludge process include
climate/temperature conditions, available land area, and variations in flow
rate and organic loading.
11.1.2 Anaerobic _P ig e s t ion
Anaerobic digestion is a process commonly used to convert raw vrastewater
sludge into inoffensive forms by decreasing its organic content. The process
biologically reduces the amount of volatile suspended solids that must be
handled by subsequent dewatering and disposal operations rendering the organic
material nonputrescible. In addition, its major gaseous end product, methane,
can be harnessed to supply plant energy needs and the digested sludge can be
.. . 15 ' . '
used as a soil conditioner.
The biodegredation mechanism proceeds in two discreet steps. First,
facultative organisms called "acid formers" degrade the complex organics of
wastewater sludge to volatile organic acids, primarily acetic acid. In the •
11-4
-------
CONTROL
VALVE PRESSURE SIGNAL
OXYGEN
VENT
OXYGEN
SUFTLY
WASTEWATER
FEED -=
AGITATOR
AERATION
TANK
u
\
LI
RFTXJRN ACTIVATtO SLUDGE
TREATED
EFFLUENT
ACTIVATED SLUDGE
Figure 11.1.1. Schematic diagram, three—stage Unox system.
Source: Reference 20.
-------
-------
TABLE 11.1.2. TYPICAL ACTIVATED SLUDGE DESIGN PARAMETERS8
Process
modification
Conventional
Complete mix
Step aeration
Contact
stabilization
Extended
aeration
Pure oxygen
systems
Sludge Food to Aerator
retention microorganism loading #BDOs
time ratio-//BOD5/ 1,000 ft3
Flow regime . (days) MLVSS/day tank volume
Plug 5-15 0.2 - 0.4 30-40
Complete mix 3-10 0.2 - 0.6 50-120
Plug 5-15 0.2 - 0.4 40-60
Plug 5-15 0.2 - 0.6 30-75
Complete mix 20-30 0.05-0.15 10-15
Complete mix 8-20 0.25-1.0 100-250
reactors in
aeries
/ Mixed liquor Detention
suspended time
solids (mg/1) (hr)
1500-3000 4-8
3000-6000 4-6
2000-3500 3-6
1000-4000b 0.5-1.5°
4000-10000°
2000-6000 24
4000-8000 1-4
Recircir ration
ratio
0.25-0.75
0.25-1.0
0.25-0.75
0.5-1.5
0.5-2.0
0.25-0.5
%alues given are for organic removals only; no nitrification.
Contact unit.
Stabilization tank.
I Reference 20.
-------
-------
second step, these volatile acids are fermented to methane and carbon dioxide
by a group of strict anaerobes called "methane bacteria."
Two main anaerobic digestive processes are used: the standard rate and
the high rate systems• Schematics of these processes, as well as their
operating criteria are provided in Figure 11.1.2. A modification of these
systems, the two-stage process, has also been successfully used (see
Figure 11,1.3). A brief description of each of these systems follows.
In a standard-rate system, the tank is not mixed and, in some cases, is
not heated. Sludge is added at the top and withdrawn at the. bottom. During
progression from top to bottom of the digestion tank, the sludge is compressed
and gradually dewatered. Stratification develops- within this plug-flow system
due to a lack of mixing. As a result, much of the digester volume is wasted,
and many operational problems result. Acidification sometimes takes place in
the top and middle layers while methane fermentation is confined to the lower
layers. This can lead to areas of low and high pH in the system, which
restrict optimum biological activity. Also, chemicals added for pH control
are not dispersed throughout the tank, and their effectiveness is therefore
limited. Grease breakdown is poor because the grease tends to float to the
,top of the digester while the.methane bacteria are.confined to the.lower
levels. Methane bacteria are removed wich the digested sludge and are not
21
recycled to the top, where they are required.
The high-rate system differs from the low-rate system in that the
contents are well mixed, either continuously or intermittently, and the
digester is heated. This procedure .avoids most o£ the difficulties inherent
in low-rate systems. Consequently, this system demonstrates improved
operation at lower retention times and higher organic loadings.
The two-stage process evolved as an attempt to provide additional gas
production and a separate settling and thickening process in the secondary
digester. The process can be used successfully when-the feed consists of
primary sludge or combinations of primary sludge and limited amounts of
secondary sludge. With the advent of wastewater treatment systems that are
more efficient than simple sedimentation, large quantities of activated and
sometimes advanced waste treatment (AWT) sludges are produced. When placed in
a two-stage anaerobic digestion process, this additional sludge can cause high
operating costs and poor plant efficiencies since the'additional solids do not
21
readily settle after digestion.
11-7
-------
GAS OUTLET
SLUDGE INLET ~
SUPERNATANT
«= ! SUPERNATANT
•c ' REMOVAL
ACTIVELY
DIGESTING SLUDGE
SLUDGE OUTLET
STANDARD RATE DIGESTION
1. UNHEATtD
2. DETENTION TIME 30 - 60 DAYS
3. LOADING 0.03 - 0,10 m. VSS/eu. fi./day
4. INTERMITTENT FEEDING AND WITHDRAWAL
5. STRATIFICATION
A
C
2
O
SLUDGE
' INLET
SLUDGE OUTLET
' (B)
HIGH RATE DIGESTION - - .
1. HEATED TO 85=- 950 f
2, DETENTION TIME 15 DAYS OR LESS
3 LOADING 0.10- D.50 Ib. VSSte. ftjday
4. CONTINUOUS OR INTERMITTENT FEEDING
AND WITHDRAWAL
5. HOMOGENEITY
Figure 11.1.2. Standard rate and high rate digestion.
Source: Reference 21.
11-8
-------
GAS
RELEASE
SLUDGE
INLET
ZONE OF
ACTIVELY
DIGESTING
SLUDGE
GAS
RELEASE
MIXED
LIQUOR
SUPERNATANT
SLUDGE
SLUDGE D.RAWOFF
SUPERNATANT
REMOVAL
TO FURTHER PROCESSING
Figure 11.1.3. Two-stage anaerobic digestion.
!
Source: Reference Zl!.
-------
Anaerobic digestion i-s suitable for nontoxic, organic~contsining sludges
resulting predominately from.primary settling! It is in widespread use,
accounting for 60 to 70 percent of biological treatment applied for primary
and secondary sludge in plants having a capacity of 1 ragd or more. The
two-stage system, while having roughly twice the capital cost of single-stage.-
digestion, is gaining in popularity. This is attributable to the increased
gas production, clearer supernatant liquor, and lower heat requirements (due
to smaller tank size) associated with its use. Anaerobic digestion is very
sensitive to process upsets due to the difficulty the bacteria have in
adjusting to_ environmental changes. However,, despite its- operating
sensitivity, anaerobic digestion is widely used due to the production of
mechane.
Since a biological mechanism is involved, the applicability of this
treatment process to the digestion of any given industrial sludge can only be
determined by specific pilot plant studies. Chemical factors are of greatest
importance to industrial sludge treatment. Close pH control ie required
because methane bacteria are extremely sensitive to. slight changes in pH. The
usual pH range required is 6.6 to 7.4, and the pretreatment of incoming sludge
to "a pH.'of 7."0' I's" desirable"." ' ' ' •""' ' '
The optimum temperature for sludge digestion is related to the
temperature response of the methane bacteria. The .rate of bacterial growth
and, therefore, the rate of process stabilization increase and decrease with
temperature within certain limits. Systems operated at high temperatures cost
more to heat, but may be justified by -increased efficiencies. Essentially all
digesters in the United States .operate between 80°F and 110°F. More important
than selection of a particular temperature is maintaining it at a constant
level. A temperature change of 1 or 2°C is sufficient to disturb the dynamic
-balance between the acid and methane formers. This will lead to an1 upset
because the acid formers will respond much more rapidly to changes in
temperature than will- the methane bacteria.
•'Knowledge of the specific nutritional reauirements of methane bacteria is
limited. Domestic wastewater appears to contain all of the nutrients required
by these organisms. However, due to the uncertainty of the precise
nutritional mix required, difficulty may be encountered when treating
wastewater of industrial 'origin. , ,
11-10
-------
11.1.3 AlgalBiodegradation. Technology
Recent, research indicated that algae may be used to remove metal ions
from wastewater or possibly concentrate valuable metals.from dilute
- 22 23 23
solutions. * Filip et al, found that when algae grown in a- sewage
lagoon were mixed with heavy metal solutions and subsequently dewatered by an
intermittent sand filter, 98 percent of the copper in solution and 100 percent
24
of the cadmium had been removed. Kerfoot and Jacobs reported rapid uptake
of cadmium by algae used in the first stage of a tertiary treatment system.
Typically, algae is contacted with the influent metals-containing
wastewater in an aerated lagoon. The lagoon is usually a lined, -flat-bottom
pond enclosed by earthen dikes.. Oxygen transfer between the air'and water
is accomplished through algae photosynthesis, although platform-mounted
mechanical aerators can be used to enhance transfer. Influent wastewater
enters near the center of the lagoon and' effluent discharges at the windward
side.
Advantages of this type of system relative to previously mentioned
biological processes include lower capital and operating costs. In~addition,
operational flexibility is increased since tha effluent flow'can be
regulated. Disadvantages include extensive physical space requirements, poor
industrial waste treatment capacity, and seasonal performance variations.
Table 11.1.3 presents the major design parameters and typical values for algae
lagoon processes for aerobic and facultative systems, with and without
supplemental mechanical aeration.
11.1.4 Pretreatment and Post-Treatment Requirements—
Industrial influents to biological waste treatment plants, are often
characterized by- periodic changes in waste volume., strength, and composition,
all of which can have a detrimental impact on maintenance of desirable
conditions. Flow equalization can. be used to lessen the chance for system
upset by dampening changes in waste quantities and qualities. Similarly,
concentrated sludge discharges can be mixed with the feed to maintain constant
solids concentration.
11-11
-------
TABLE 11.1.3. EMPIRICAL DESIGN CRITERIA FOR WASTE STABILIZATION LAGOONS
i
h-»
ro
Aerated
Parameter
Flow regime
Lagoon size (acres multiples)
• Operation
Hydraulic retention time (days)
Depth (ft)
Hydraulic loading (In. /day)
DODs loading:
(Ib/day/acre)
(Ib BOD/day/1.000 ft3)
Optimum temperature (°C)
Temperature range (°C)
BOD5 removal efficiency (%)
Algal concentration (mg/1)
Coliform removal %
Algne
Hie roorgnn lams
' Otlier
Effluent BOD5 (mg/1):
Soluble BOD 5
; Insoluble BODs
Typical effluent quality, (mg/1):
• BODs
ss
IP"
. Oxygen source
Aerator design goals
Aerobic
10 acre multiples
Series or parallel
10-40
3-4
3-5
60-120
'
20
0-40
80-95
80-200
>99
(0.4-1.2)99
(0.2-0.8) (BOD;)!
(0.2-0. 5) (BODs)!
(0.1-0.4)(SS)±
(0.02-0.1)(BOD5)i
(0.3-1.0)(SS)e
15-40
25-50
'": 6.5-9.0
Algae
; —
Aerobic
Completely mixed
2-10 multiples
Series or parallel
3-20
6-20
"
20-400
-
20
0-40 '
80-95
-
-
(0.02-0.1) (BODs)i
(0.2-0.5) (BODs^
(l.l-1.4)(SS)i
(0.02-0.1)(BOD5)l
(0.5-0. 8) (SS)e
20-70
.
6.5-8.5
Aerators
Aeration plus mixing
Facultative
Mixed surf, layer
2-10 multiples
Series or parallel
7-20
3-8
- -
-
-
20
0-50
80-95
-
(0.02-0.1)(BOD5)i
(0.2-0.5)(BODs)i
(0.1-0.4) (SS)i
(0.02-0.1)(BOD5)i
(0.3-0. 8) (SS)e
20-70
-
6.5-8.5
Aerators
Aeration
influent, c - effluent.
Source: Reference 15.
-------
Where the equalization basins are not on-line with the continuous waste
flow, special overflow weir or sensor—actuated flow gates must be provided for
temporary diversion of flows to the basin structure. Where the basin is used
for protecting the on~line biological treatment processes, methods must be
provided for anticipating qualitative changes so that the appropriate waste
volume can be' diverted before the normal treatment scheme can be restored.
The acidity or alkalinity {pH} of the waste stream introduced to the
bioreactor must be maintained within a specific range to preserve microbial
populations. To accomplish this, different industrial waste streams may be
combined for treatment based on their neutralizing effects, or chemicals may
be purchased and added to the influent wastes. In the former s-ituation, lower
operational costs are realized (no chemicals must be purchased), but
maintenence of pH is dependent on the consistency of each component stream.
The stockpiling of chemicals such as sulphuric acid, caustic soda or lime will
increase chemical costs, but will also provide the capability for responding
to variations in waste stream characteristics or flows by adjusting chemical
additions. Neutralization of highly concentrated waste streams may be most
effectively achieved before they are mixed with other, more dilute waste
'streams.
.-. ^ ^ $,. , ^ , , - . .
Other pretreatment techniques may also be practiced in order to enhance
the biodegradation of problematic waste streams. The use of cooling towers
should be considered as an effective means of enhancing biodegradation. If
inlec wastes are not, cooled ,to ac least 40"C to 45°C, they may adversely
affect the microorganisms in the bioreactor. Some technologies,« such as
solvent extraction, are best applied to individual process waste streams
before they are combined with other industrial waste streams prior to
treatment. Other techniques applicable to single or mixed waste flows
include: reverse osmosis, chemical precipitation, evaporation, ion exchange,
distillation, resin adsorption, and gravity separation. Powdered activated
carbon has been shown to be effective in adsorbing and attenuating compounds
thereby limiting the. toxic effects of concentrated wastes.
Post-treatment of biological residues containing heavy metals are often
restricted by the presence of these rnetals. For example, high metal levels
can result in air pollution, ash disposal, and mechanical operating problems
during sludge incineration (see Section 12). The presence of cadreium and
.1-13
-------
other metals at excessive levels can prevent Che sludge -from, being disposed
25 ""
via land application. Depending on the level of heavy metals accumulated
within the sludge product, the most likely method of'disposal would be
solidification followed by landfilling. Note that, these sludges are unlikely
to, contain metals at concentrations which would .prohibit them from being land
disposed under the 1984 HSWA regulations. Alternatively, extraction for
metals removal or- incineration may be the most viable post-treatment methods.
11.2 PROCESS PERFORMANCE
Numerous research studies have been conducted to investigate the toxic
effects of heavy metals on conventional biological treatment,processes. The
following is a brief summary, by metal type, of the adaption of biological
systems to heavy metals removal. This is followed by a discussion of the
effects of synergistn and recent developments concerning the use of novel
organisms for biological treatment. While the available literature emphasize
activated sludge treatment, anaeorbic and algal systems have been increasingly
jy 0*
explored in recent years and are also discussed. • '
*<;;--•
11.2.1 Zinc . . . -
The percent removal of zinc in activated sludge treatment is normally
very good compared to other metals. Typical values range from as low.as
22 percent to as high as 68 percent, averaging better than 50 percent. Both
soluble and insoluble zinc is removed mainly in the aeration basin absorbed by
'•• ' 27
the microbial floe. Pilot plant studies on activated sludge treatment
showed zinc removals ranging from 74 to '95 percent at concentrations ranging
from 2.5 to 20 mg/L. BOD removals for these^zinc concentrations were only
slightly affected. It is not expected that municipal plants achieve these
removals, but it is interesting to note that"the pilot studies have no
supernatant recycle.
Anaerobic digestor operation has been found Co be tolerant of zinc
influent levels up to 20 mg/L.2? Zinc-cyanide complexes can cause digestor
problems if the digestor is not previously acclimated to low cyanide- levels.
-An al.gal system using Chlorel'Ia -pyrevoidosa was -reported- -to 'toler-ate up to
10 mg/L zinc over a-24-hour period. However, removal efficiency was poor
•y o
(29 percent) and 103 mg/L of Chlorella was required.
11-14
-------
11.2.2
In municipal treatment plants, copper is often encountered as copper
-sulphate or copper cyanide complex since these are common plating wastes. A
29
field survey of treatment plants .showed influent copper values ranging
from 0.2 to 6,8 mg/L. While copper removals of only 37.2 percent were
obtained at a full-scale facility (the Ukima Treatment facility in Japan) ,
results in Table 11.3.1 from pilot plant studies 'Show 50 to 75 percent
removals are obtainable vihen copper is provided as copper cyanide. Of
interest in this study are the "low Cu residuals obtained at 1.2 mg/L Cu feed
levels. It was also found that 1 mg/L Cu actually increased sludge
settleability leaving BOD and COD removals essentially unchanged
(Table 11.3.2). In another study, an algal system using Nostoe museormn on ,
influent copper concentration of 10 mg/L reported a 48 percent reduction over
a 24-hour period.
These results tend to support the theories of Wood et al. concerning -
the growth of filamentous organisms which can enhance metal removal through
adsorption. Sludge feed doses of copper sulfate were largely absorbed by the
activated sludge and then released slowly, minimizing toxic effects on the
microorganisms.
11.2.3 Nickel
32
"Nickel removals by activated sludge were found by McDermott, et al.,
to be roughly 30 percent for influent concentrations ranging from 2.5 to
10 mg/L. The BOD removal efficiency was reduced only 5 percent at these
concentrations, however, a sludge dose at 200 mg/L was found to have
significant adverse effects on system operation. Anaerobic digestion was
found to reduce soluble Ni content of sludge to & constant level of 8 to 14
mg/L, regardless of the initial feed levels. However, these results were
33
contrary to the findings of other investigators
Another experiment tested six strains of algae and Eugleve sp. for their
4
.ability to bioaccutnulate nickel. The researchers found that, at pK 8.0,
algae could tolerate up to 7.8 mg/L of nickel before cell lysis would occur.
-------
-------
TABLE 11.2.1. FATE OF COPPER FED AS COPPER CYANIDE COMPLEX
IN ACTIVATED SLUDGE TREATMENT
Copper in sewage feed (rag/L)
Type of check sample Location of check sample 0.4 1.2 2.5 5
Copper
outlet
fed found in
(2)
Primary sludge
Excess activated sludge
Final effluent
Unaccounted for
__
—
43
—
12.
43.
25.
20.
5
3
1
0
10.7
25.6
43.3
20.0
7
23
50
20
Efficiency of copper
removal (%)
Soluble copper in Total
primary effluent (tng/Lj Reactive
Soluble copper in
effluent (mg/L)
Total
Reactive
57 75 57 50
0.22 0.19
2.56
0.12 0.10
0.67 0.92
Source: Reference 31.
11-16
-------
-------
TABLE 11.2.2. EFFICIENCY OF ACTIVATED SLUDGE
TREATMENT OF SEWAGE-CONTAINING
COPPER FED CONTINUOUSLY
Copper
(mg/L)
0
0.1
1.2
2.5
5.0
10.0
BOD
removal, average
U)
95
95
93
91
89
88
COD
removal, average
{%)
85
85
84
: 85
76
69
Source'-: ••• " Reference-- 31. ; -
11-17
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11.2,4 Lead . ' .
34
Lead removals obtained by Brown, et al., were on.the order of
55 percent for secondary plants. Lead removal was found to -be enhanced by
longer settling times and larger floe size in the activated sludge process.
o A
Lead removals using Anabaena Flos-Aquae investigated by Sloan et al. , wen
on the order of 65 percent when initial lead concentrations were 4.0 mg/L.
11.2.5 Cadmium
Cadmium removals are normally not very high in biological treatment,
partly due to the low concentrations usually encountered. Low removals of
cadmium can be attributed.to its high solubility, ability'to form other
complexes at the pH of sewage, and competition from other metal ions in the
influent. However, when an influent concentration of 2.0 mg/L of cadmium was
contacted with the algae Anabaena Flos-Aquae, a removal efficiency of
2 8
70 percent was reported over a. 24-hour period. •
• 11.2.6 Chromi'dtn
Chromium removals by operating activated sludge plants are generally
34 35
around 40 to 60 percent. * Hexavalent chromium is normally reduced to
the trivalent state before it -is removed by microbial floes. However, in a:
n , £
pilot-plant study by Moore, et al., the prior reduction of Cr by means
of biological reduction showed approximately 92 percent removal of CR at
feed levels of 46.5 mg/L, The biological reductor is a complete mix reducing
basin, with chromate serving as the principal source of 'oxygen. The BOD
removals of the activated sludge process in conjunction with.the biological
reductor were around 94 percent, indicating little decrease in efficiency due
to presence of the metal. Gas production in the digester was also not
affected by chromium in the sludge, although these results were contrary to
o c.
those found by other investigators
11-18
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11.2.7 Synergistic Effects '
29
Earth, et al., wrote a summary report on the effect of heavy netals
on various biological units. Using Che data compiled from an activated sludge
pilot plant, they concluded that the aerobic biological treatment system could
tolerate a total heavy metal concentration of up to 10 mg/L (Cr, Cu, Ni, and
Zn) , either singly or in combination, with only about a 5 percent reduction of
29 37
the BOD removal efficiency. * They further concluded that a small dose
of metal could noticeably reduce the treatment efficiency, but this effect
diminished with larger doses, Nitrifying microorganisms were also reported to
be sensitive to heavy metals. The investigators reported that 5 mg/L of
copper in influent sewage was the highest dose'that could still allow
satisfactory anaerobic sludge digestion. Finally, they reported that limited
metals could have beneficial effects; e.g., reducing the degree of sludge
bulking problems in aerobic systems. These conclusions have been confirmed in
""3 f. "^1 C "3 Q
other studies including Dawson and Jenkins, Jenkins, and Tarvin,
11.2,8 Recent Developments
In the past 5 or 6 years, researchers have been working on the
development of specific organisms designed primarily for the bioaccutnulation
of heavy metals. The following is a brief discussion of some of the more
promising technologies.
Researchers at the Hebrew University in Rehovot, Israel, have developed a
40
ne" method for removing metals from wastewater using water ferns. Azolla,
a water fern found in Asia, East Africa and Central America, can be used to
remove metals such as copper, zinc, chromium, cadmium, nickel, silver,
titanium and uranium from industrial waste. It can be grown in settling ponds
and, when harvested and dried, used as filtering material in paint and
metals-processing plants. - '
In the United States, a process developed by Rerr-McGee of Oklahoma City,
40
Oklahoma, was used to biologically remove selenium from wastewater. The
researchers found that selenium can be 'removed from uranium-mine wastewater by
anaerobic Clostridium bacteria. At a scale of 100 gal/d, selenium
11-19
-------
concentrations have been lowered from 1.6 mg/L to below 0.5 mg/L. The
organism is proving to be more effective in this application than ion exchange
and reverse osmosis which were also tried.
At New Mexico State University in Las Cruces", Dennis'Darnell and his
research group are using green algae, Chlorel'la vulgaris, to recover metals
from waste streams at a cost that is a mere 1 percent to 2 'percent"of the cost
of exchange resins which are currencly used. The researchers immobilized the
algae on silica gel and have run as many as 20 cycles with no decrease in
effectiveness. Acidity and salt content of solutions can reportedly be
adjusted co retrieve metals selectively. Chromium, silver, mercury, platinum,
and other metals have been removed in trials, but gold was found to be the
most tightly bound by the algae. More than 90 percent of gold was removed
from a test solution, although the inlet concentration was not specified in
, 40
the reference.
Researchers at Austria's Institute of Microbiology at the University of
Innsbruck have discovered that certain fungi and bacteria are able to "filter"
silver from dumped waste material and store, it in their cells. This discovery
40
could lead to the recovery of precious metals from industrial sewage.
•-• -In TSweden,'cRolf 'OP Ha'tliber-g -h'ays-ai's"covered a'process fo'r""fe!moving "heavy7"'
metals from wastewaters containing sulfate ions by means of sulfate-reducing
41
bacteria. The bacteria can be any of the known sulfate reducers,
including Dasulfoyibrio and Desulfotomaculum. The bacteria reduce the,sulfate
to sulfides, producing hydrogen sulfide gas, leaving the heavy metals to
precipitate out as sulfides. Two vessels are used, one for culturing the
bacteria in a nutrient and the wastewater, and the other for precipitation.
Holding time in the culture vessel may be 10 to 40 hours. Aqueous solution of
hydrogen sulfide produced- in the culturing vessel is-fed continuously into the
precipitation vessel along with the remainder of the wastewater. The
resulting precipitate is flocky and settles easily.
An example presented in a. patent for the process used simulated "
wastewater with a sulfate ion content of about 600 rcg/L, 10 mg/L copper,
600 mg/L zinc, and 500 mg/L iron, Unfiltered water from the precipitation
vessel contained up to 0.1 mg/L copper, 0.1 mg/L zinc, 10 mg/L hydrogen
sulfide, and 10 to 50 mg/L iron. Iron content could be decreased to zero by
11-20
-------
adjusting pH in the precipitation vessel. Another poasible process variation
would be to aerate the output to oxidize residual hydrogen aulfide to sulfate,
and iron to Fe and Fe for reuse.
A strain of Pseudomonaa fluorescens that .reduced chromate ions to a
41
precxpitatable form has been found by Lawrence H. Bopp, The strain,
designated LB300, reduces Cr to Cr , which precipitates and is thus
removed from the wastewater. The organism can be used to detoxify chromate in
a contaminated sewage digester in which nicrofIfora have been killed by
chroroate-bearing sewage. After detoxification with LB300, normal microflfora
growth can be reestablished. .
LB300 is resistant to potassium chromate concentrations as high as
2,000 ppm in a minimal salts medium. That is high.enough to include most
industrial effluent of chromate, such as chrome plating wastes or wastes from
chromate ore processing. LB300 can. be used-under aerobic or anaerobic
, , 41
conditions at temperatures between -4° and +35DC.
11.3 PROCESS COSTS
Widespread use. of- aerobic b.iol.og,ical treatment systems- has led to well
developed cost estimation procedures basing capital, operational, materials,
and labor costs on system capacity. Estimated treatment system outlays can be
determined using Table 11.4.1 and Figures 11.4.1 and 11.4.2, although these do
not include the additional costs of seed chemostatic organisms to be used in
bioaugmented processes. More complete and up-to-date cost information can be
• • 42
found in the EPA publication " Estimating Water Treatment Costs".
Although the breadth of this document prevents its inclusion in this section,
the data presented here'do show the relative costs and scaling factors used
43 44
for various cost elements. * All costs have been updated to 1987
45
dollars, using the Chemical Engineering Index.' .
Standardized cost data for anaerobic treatment systems were not found.
An example of a modern, anaerobic system is the "Celrobic" process developed by
44
Ceianese, '• In 1983, a 1.08 million gallon/day waste stream with an
influent COD of 3.3 g/L, incurred outlays of $8,100,000 in"capital costs and
$400,000 in annual operating costs (1982 dollars). This plant was expected to
produce 220 million cubic feet of methane gas annually which considerably
reduced its net annual operating costs.
11-21
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TABLE 11.3.1. ESTIMATED CAPITAL COST FOR WASTEHATER TREATMENT UNITS
Treatment unit
Parameter
Model cost
(1987 dollars)
i
S3
NJ
Raw wastewater pumping
Screening, grit removal and
flow measurement
Equalization
Primary sedimentation or
secondary clarification
Aerat ion-bas in
V
Aeration-diffused air system
Aerat ion-surface
Trickling filter
Recirculation pumping
Sludge digesters and buildings
Lagoon
Vacuum filtration
Centrifugation
Incineration
Capacity (mgd)
Capacity (mgd)
Volume (Mgal)
Surface area (in 11,000 ft2)
Volume (in 1,000 ft3) C = 10.5 x 103
Blower capacity (in 1,000 cEm) C = 22.5 x 104 (cap)0-72
C = 6.5 x 103 (mgd)1-0
C = 67.5 x 104 (mgd)0-62
C = 18.0 x 104 (Mgal)0-52
C = 7.0 x 104 (A)0-88
Capacity (horsepower)
Media volume (in 1,000 ft3)
Capacity (mgd)
c'= 2.5 x 103 (hp)°-89
C = 8.5 x 103 (V)0-84
C = 6.3 x 104 (mgd)0-70
Sludge volume (in 1,000 ft3) C = 3.5 x 104 (v)°-64
C = 18.0 x 104 (V)0-71
C = 14.8 x 103 (A)0-67
C = 8.3 x 104 (gpm)0-54
Volume (Mgal)
Filter area (ft2)
Capacity (gpm)
Dry solids capacity (Ib/hr) C = 3.5 x 104 (cap)0-56
Source: References 43 and
-------
2,500
5 1C 20 SO 10D 200 SOO 1000
Capacity
Figure 11.3.1- Estimated annual operating and materials costs as a
function of wastewater treatment facility capacity.
Source: References 43 and 44.
11-23
-------
25,000
250
1000
Figure 11.3.2. .Estimated annual man-hours needed'for. wastewater
treatment facility operation. -
Source: References 43 and '44.
11-24
-------
Actual treatment costs will depend on specific characteristics of the
waste stream. Pertinent data needed for cost estimation are: waste stream
volumetric rate, organic compound constituents and concentrations, other waste
characteristics such as influent BOD, COD, or level of toxins,, treatment
design, and overall treatment objectives.
il.4 OVERALL STATUS OF BIOLOGICAL TREATMENT
A large number of companies exist that specialize in the design and
construction of biological treatment systems. Aerobic systems are the most
readily available, and their design and operation are complex, but
manageable. The Local number of facilities using some sort of aerobic
46
biological treatment is over 2,000, Conversely, the number of companies
offering expertise in bioaugroentation and anaerobic treatment is relatively
small, but this segment is expected to grow rapidly,'
Biological treatment of metals using conventional equipment and
acclimated strains is typically only capable of treating combined heavy metal
37
influents of 10 mg/L. While improvements in process tolerance for
inorganic priority pollutants is encouraging, most advancements are still in
' '41
the developmental stage and have yet to be widely applied.
11-25
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REFERENCES
1. Wood, D.K., et al. Trace Elements in Biological Treatment' with Specific
'Reference to the Activated Sludge Process, Proc. 29th Industrial Waste
Conference, Part II, Purdue University. 648. May 1974,
2, ' Rnauss, H.J., and S.W. Porter. The Adsorption of Inorganic Ions by
Chlorella Pyrenoidosa. Plant Physiology, 29(3). May 1954. -
3. Wood, J.M. Proc. Nobel Conference Inorganic Biochemistry,'Chemical
Service, 21, 155-160. 1983.
4, Wang, H.K., and J.M. Wood. Bioaccumulation of Nickel by Algae. . .
Environmental Science Technology, Vol. 18, No..2. 1984.
5. Dugan, P.R., and H.M. Dickrum. Removal of Mineral Ions from Water by
Microbially Produced Polymers. Proc. 27th Industrial Waste Conference,
Purdue University, Lafayette, Indiana. 1019. May 1972.
6. Unz, S.F. Microcultures Studies of Activated Sludge Bacterial
Zoogloeas. Abstr. Ann.. Amer. Soc. Microbiol. _2jf(32). 1972.
7, Hayes, T.D. et al., Department of Agricultural Engineering, Cornell
University. Heavy Metal Removal from Sludges Using Combined
Biological/Chemical Treatment. 34th Industrial Waste Conference, Purdue
University. 1979.
8. Gould, M.S., and E.J. Genetelli. Heavy Metal Cdtnplexation Behavior in
Anaerobically Digested Sludges. Water Research, 12, 505. 1978.
9. Eckenfelder, W.H. Water Quality Engineering for Practicing Engineers.
Barnes and Noble, New York, NY. 1970.
10. Doyle, J.J., Marshall, R.T., and W.H. Pfander. Effects of Cadmium on the
Growth and Uptake of Cadmium by Microorganisms. Applied Microbiology,
29(4), 562. 1975. • •
11. Hayes, T.D., and T.L. Theis. The Distribution of Heavy Metals in
Anaerobic Digestion. Journal Water Pollution Control Federation, 5011),
61. 1978.
12. Patrick, P.M., and M. Loutit. Passage of Metals in Effluents, Through
Bacteria to Higher Organisms. Water Research, 10, 333. 1976,
13. Chang, S.H., et al. Effects of CD (II) and Cu (II) on a flioFilm System.
Journal of Environmental Engineering, Vol. 112, No. 1. February 1986.
14. Sicrp, F., and F. Fransetnier. Copper and Biological Sewage Treatment.
Vom Wasser, 7. 1933.
11-26
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15. Chillingworch et al.,. Alliance Technologies Corporation. .Industrial
Waste Management Alternatives Assessment for the State of Illinois.
Volume IV. February 1981.' • • "• . ' '
16. Benjes, H.H., Jr. .Handbook of Biological Wastewater Treatment. Garland
STPM Press, New York. 1980. ' '•..;.' '••_
17. Clark, Viessman, Hammer. Water Supply and Pollution Control.
Harper & Row. 1977. . :
18, Gurnham,'C.F. Industrial Wastewater Control. Academic Press, New York.
1965. ' >
19. Shreve, R.N., and J.A. Bunk, Jr. Chemical Process Industries.
McGraw-Hill, New York. 1977.
20. Hinrichs, D.J. Inspector's Guide for Evaluation of Municipal Wastewater
Treatment Plants. EPA-430/9-79-010, U.S. Environmental Protection
Agency. April 1979.
21. U.S.- EPA Process Design Manual for Sludge Treatment and Disposal.
EPA-600/1-74-006. October 1974.
22. Basset, J.M., Jennett, J.C., and J.E. Smith. Heavy Metal Accumulation by
Algae. In: Contaminants and Sediments, R.A. Baker Editor, Ann Arbor,
Ann Arbor Science Publishers, Inc. 1980.
23. Filip, D.S. Peters, T., Adams, V.D., and E.J. Middlebrooks. Residual
Heavy Metal Removal by an Algae-Intermittent Sand Filtration System.
Water Resources. 13:305. 1979.
24. Kerfoot, W.B., and S.A. Jacobs. Cadmium Accrual in Combined Wastewater
Treatment Aquaculture System. Environmental Science & Technology.
10(7):662. 1976. .
25. Patterson, J.W., and Hao, S.S. Heavy Metals Interactions.in-the.
Anaerobic Digestion System. 34th Industrial Waste Treatment Conference,
Purdue University. 1979. . . •
26. Russell, H.H. et al. Impact of Priority Pollutants on Publicly Owned-
Treatment Works Processes: A Literature Review. 37th Industrial
Waste Treatment Conference, Purdue University, 1982.
27. McDertuott, G.N., et al. Zinc in Relation to Activated Sludge and
Anearobic Digestion Processes. Proc. 17th Industrial Waste Conference,
Lafayette, Indiana, May 1962. Engineering Ext. Series 112, Engineering
Bulletin, Purdue University, 47(2):461. 1963.
28. Sloan, F.J. et al., Clemson University. Removal of Metal' Ions from
Wastewater by Algae. 38th Industrial Waste Conference, Purdue-
University. 1983. ... • ."•••
11-27
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29, Bartb, E.F,, English, J.K., Salotto, B.V., Jackson, B.N., and
H.B. Ettinger. Field Survey of Four Municipal Wascewater Treatment
Plants Receiving Metallic Wastes. M.P.C.F. Journal _32.:1101» August 1965.
30. Rondo, J. Some Problems on the Joint Treatment of Industrial Wastes and
Sewage in the Ukitna Treatment Plant. Water Research 2-' 375-384. 1973.
31. MeBennott, G.N., et al. Effects of Copper on Anerobic Biological
Treatment, W.P.C.F. Journal, 15(2): 227. February 1963.
32. McDennott, G.N., et al. Nickel in Relation to Activated Sludge and
Anaerobic Processes. W.P.C.F. Journal, 37J.2); 163. February 1965.
33. Lawrence, A.W., and P.L. McCarty. W.P.C.F. Journal, 37(3): 392. March
1965.
34. Brown, H.G., Hensley, C.P., WcKinney, G.L., and J.L. Robinson.
Efficiency of 'Heavy Hetals Removal in Municipal Sewage Treatment Plants.
Environmental Letters, .5(2): 103-114. 1973.-
35. Moore, W.A., et al. Effects of Chromium on the Activated Sludge'
Process. W.P.C.F. Journal, 33: 54, January 1961.
3b. Dawson, P.S., and S.H. Jenkins. The Oxygen Requirements of Activated
Sludge Determined by Manometric Methods. Sewage and Industrial Wastes,
22: 490. 1950.
37. Earth, E.F. , Salotto, . B..V. , ;.McDertoott.,-G..»._,'. English, ' J .N.,* and" " ''
h.B. Ettinger.- Effects of a Mixture of Heavy Metals on Sewage Treatment
Processes. Proc. 18th Industrial Wastes Conference, Purdue University.
p. 616. May 1963.
38. Jenkins, S.H. Trade Waste Treatment. The Institute of Sewage
Purification, 28: 1371. 1956.
39. Tarvin, D. Metal Plating Wastes afld Seuage Treatment. Sewage and
Industrial Wastes, 28: 1371. 1956.
40. Greene, R. Biotechnology and Pollution Control. Chemical Engineering.
March 4, 1985.
41. Technical Insights, Inc. New Methods for Degrading/Detoxifing Chemical
Wastes. Emerging Technologies, No. 18. 1986.
42. U.S. EPA. Estimating Water Treatment Costs. U.S. Environmental
Protection Agency, Municipal Environmental Research Laboratory,
Cincinnati, Ohio. EPA-600/2-79-162(a,b). August 1979.
43. Black & Veatch. Estimating Costs and Manpower Requirements for
Conventional Waste«ater Treatment Facilities. U.S. Government Printing
Office, Washington, D.C. 1971.
11-28
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44. Sundstrom, D.W., and-H.E. Kiel.- Wascewater Treatment. Prentice-Hall,
Inc., Englewood Cliffs, N.J.S- pp. 439. 1979. : ;
45. Chemical Engineering. Chemical Engineering.. Plant'Cost Index. April 27,
1987. . . •• :';•£.;.•••.• •• / . '•::.•'.••.'
46. U.S. EPA. Background "Document-.for Solvents to Support 40 CFR Part 268
Land Disposal Restrictions. Volume II. January 1986.
47. U.S. EPA. Selected Biodegrsdation Techniques for Treatment and/or
Ultimate Disposal of Organic Materials. EPA-600/2-79-006. March 1979.
11-29
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SECTION 12.0
THERMAL T11ATM1NT/RECOVERY PIOCESSES
Processes included here may be used to treat hazardous wastes containing
many of the toxic heavy metals. Processes such as incineration and the
"pyrometallurgical" processes such as calcination or smelting, and recovery
processes such as evaporation and crystallization, may be used to separate the
metal compounds from the other waste constituents. The application of these
processes will allow either their recovery (e.g., by crystallization) or
concentration and ultimate disposal through techniques such as encapsulation
of incineration aeh that are most appropriate for concentrated wastes. Since
the metal compounds cannot be "broken down" in the same manner as, for
example, an organic compound nay be broken down through pyrolysis to simpler
compounds or oxidized to form C05 and water, the usefulness of the thermal
destruction processes as a means of concentrating metal wastes strongly
depends upon the nature of the other waste constituents. To a lesser extent,
the applicability of a thermal technology may also depend upon the volatility
of the metals, and the physical form of the waste.
In this chapter, the available technologies to be discussed are as
follows:
* Incineration;
* Calcination/Smelting (Pyrometallurgy);
* Evaporation; and
* Crystallization.
The last two processes are recovery processes which are more physical
than thermal in nature. Evaporation depends upon removal of a volatile,
12-1
-------
-------
nonmetallic component of the waste, usually water, to concentrate metallic
'salts. Crystallization, often used in conjunction with evaporation, involves
cooling a solution to reduce the solubility of the metal salts and bring about
precipitation.
While many hazardous wastes containing heavy metals may be Rood
candidates for thermal destruction processes, as classified above, there are
some which should definitely never be handled in this manner. Such wastes
include: .
Organoarsenic compounds, which when combusted yield arsenic III
oxide, which has a fairly low boiling point. To effectively.capture
this material from combustion vent gases, the gases need to be
cooled to near ambient temperatures before dry or wet collection of
particulate arsenic III oxide can be satisfactorily conducted.
Selenium compounds; the selenium dioxide generated is a low boiling
compound which, like arsenic oxide, will be difficult to contain.
Wastes containing chlorides and chromates which, during incineration,
may generate ehromyl chloride (Cr02Cl2), which has boiling point
of only 117eC. For this substance, the same problems as those
discussed for organoarsenic compounds are encountered.
Explosives; explosive mixtures should not be handled by incineration
for obvious reasons.1
In general, there are several disadvantages inherent in using, thermal
destruction processes to handle metal-bearing wastes, including the following:
• All thermal destruction processes will create air emissions, in many
cases including heavy metal particles or vapors;
* Thermal destruction processes will often form chemical by-products
such as hydrochloric acid which may 'be damaging tor- the systens
themselves, or which, like dioxins, may present serious
environmental hazards which are as significant as those posed by the
toxic heavy metal wastes;
» Thermal destruction processes require the generation or removal of
heat energy through fuel burning or consumption of electrical power
and therefore may not be economically attractive.
12-2
-------
-------
In this chapter, the amenability of metal wastes to the specific
processes will be reviewed, focusing on specific performance test results to
develop an understanding of which physical/chemical parameters dictate the
acceptability of each process.
12,1 INCINERATION
Incineration is the process of applying sufficient heat -energy and oxygen
LO cause the oxidation and/or pyrolysis of compounds such chat* they are broken
down to form more "basic" chemical species such as water, C0_, HC1,
elemental metals or metal oxides, etc. Incineration' processes are often
considered an attractive management alternative for hazardous wastes because
they possess many advantages over other technologies, including Che following:
Thermal destruction by incineration provides the ultimate disposal
of hazardous wastes, minimizing future liability from land disposal;
Toxic components of hazardous wastes can be converted to harmless or
less harmful compounds;
The volume of waste material may be reduced significantly by
incineratio.n; and- , • . -,-, _ -
Resource recovery (i.e., heat value recovery) is possible through
combustion.
Unfortunately, because metals are not destroyed by incineration and have no
heat recovery value there is little incentive to treat most metal bearing
wastes by incineration. An exception might be an organometallic compound
containing waste such as tetraethyl lead or a cyanide complex which is highly
toxic and not readily treated by more conventional methods.
Several problems must be faced when incinerating metal—wastes. A prinsary
consideration is- the extent to which air emissions of toxic heavy metal
particles or vapors will be generated. Certain metals and their oxideSsuch as
mercury, lead, selenium, and arsenic are volatile, particularly at the
elevated temperatures of incineration. A significant percentage of the input
of these relatively volatile metals will be emitted as a vapor or as fine
particulatea which are difficult to control- A second problem involved in the
12-3
-------
Incineration of such wastes ie the generation of ;an incinerator ash or sludge
containing metals or metal oxides, which will require safe disposal. Third,
wastes containing high concentrations of noncotnbustible materials require
greater energy input via auxiliary fuel combustion, thus increasing processing
costs significantly. Finally, such wastes nay be difficult to handle in
certain incineration systems. Liquid injection incinerators may not be used,
for example, should the solids content of the waste be such that the injectors
will be come clogged.
Although numerous studies of incinerator performance have been conducted
in which organic wastes containing metals were burned, the available data are
limited in content relative to the effect of metals on combustion. Based on
the available data, it does not appear as though the presence of metals in
small concentrations will hinder the destruction of organics. The data do
show, however, that certain tnetal species may present more of a concern
relative to potential air emissions than do 'others.
Incineration facilities permitted to operate by EPA under RCRA are
2
required to achieve at least a three tiered environmental standard:
1. They must achieve a destruction and removal efficiency (DRE3 -of
99.99 percent for each principal organic hazardous constituent
(POHC); . • • .
2. They must.achieve a. 99 percent HC1 scrubbing efficiency or emit less
than 4 Ibs/hr of hydrogen chloride; and
3, They must not emit partieulate matter in excess of 0.08 grains/dscf,
(0.18 gratas/dscra) corrected te 7 percent oxygen.
The HC1 and particulate matter standards will be exceeded by most uncontrolled
incinerators burning even relatively clean wastes. For, example, the HC1
standard-of 4 Ib/hr will be exceeded by units larger than 3.8 x 10 Btu/hr
burning a 19,000 Btu/lb waste containing 2 percent chlorine. The ash content
corresponding to the 0.18 g/dscm particulate emission standard would be about
• *l^Z-
0.3 percent assuming incineration of a similar high Btu fuel at 7 percent
3
oxygen. Thus, in order to comply with the particulate emission standard,
control devices will be required to reduce air emissions when burning a waste
containing higher concentrations of metals For example, a control device of
about 90 percent efficiency will be needed to achieve the standard for a waste
12-4
-------
containing 3 percent metal.. The exact value will depend.upon:the heating
value and composition of combustibles in the waste (and auxiliary fuel) and
the fraction of the metal input emitted with the flue gas. : .• •
4 .••--'•.•' ./•••-.•
According to Oppelt, more than half of the incinerators-operating in
1981 used no air pollution control system at all. These uncontrolled
incinerators would not be suitable for the incineration of metal bearing
wastes. Standards for emissions of toxic air pollutants such 88 toxic metals
may also limit incineration of taetal«bearing wastes.
12,1.1 Process Description . . , . '. .
Hazardous waste incineration technologies range fron those with
widespread commercial application and many years of proven effective
performance, to those currently in'development. As many as 6? companies may
be involved in the design and development of hazardous waste'incinerators,
with more expected as limitations on land disposal of hazardous wastes
5
increase. •
As mentioned previously, there are several incineration technologies
which have become established commercially as the primary options available
for the incineration of hazardous'wastes. These technologies have Seen
demonstrated extensively for a wide range of hazardous wastes. They comprise
ft 7 K
about 80 percent (by number) of the U.S. market. ' ' They include;
* Liquid injection incinerators;
* Rotary kilns; . ,, ,
* Fluidized-bed incinerators; .
* Fixed hearth incinerators, -particularly the starved air or pyrolysis
type units; and -
* Multiple hearth incinerators.
Liquid injection {66 percent), rotary kiln (12.3 percent), and fixed
hearth incinerators (18.5 percent) are the most widely used for the disposal
of hazardous wastes. A detailed discussion of the design and operation of
these systems can be found in Reference 3, or in the open literature.
12-5
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The relative ease with which hazardous wastes containing toxic heavy
metals or metal compounds may be incinerated has been studied within the
context of a general study of hazardous waste ineinerability conducted by
9
EPA. A summary of the "incinerability ratings developed by EPA for such
wastes is presented in Table 12,1,1. As shown, almost all of the
metal-bearing hazardous wastes are perceived to be "poor" candidates for
incineration. The results of a study analyzing "incineration risk" conducted
for EPA in 1981 by ICF, however, showed that 35 of a total of 139
hazardous waste streams currently incinerated (25 percent) contain toxic
metals or metal salts. The metal-containing wastes do not, however, account
for a significant percentage of the total volume of hazardous waste
incinerated. In face, the study showed that approximately 90 percent of the
volume of wastes incinerated are characterized as D001 and DQ02 wastes, about
which little is known. It is not expected that such waste would contain an
appreciable amount (if any) of heavy metal.
The primary characteristic of metal-bearing hazardous wastes which might
limit incinerability is the concentration of the metal itself. Other limiting
factors relate to the characteristics of the other waste constituents,
e.g., organic solvents. A detailed discussion of those components is
presented in Reference 3. Most commercial incineration facilities surveyed by
Alliance indicated metals concentration limits in the 1 to 500 ppm
range. *" Most are limited to such low feed concentrations by air
emission regulations or effluent guidelines. Thus, although a higher metals
concentration may not necessarily render a waste less coobustible, the air
emissions and/or ash or effluents generated may preclude the incineration of a
particular waste.
As a result of this limitation, many commercial facilities blend
metal-bearing wa-s-tes with other .compatible waste streams in order to achieve a
proper concentration level. Blending may also serve to enhance the
combustibility of the waste stream (i.e., raise the heat value). No other
form of pretreatment appears to be used by the commercial incineration
industry.
12-6
-------
TABLE 12.1.1. METAL WASTE INCINERAB1LITY
Waste
code
Description
Rating8
D004 Arsenic
D005 Barium
D006 Cadmium
D007 Chromium
D008 Lead
D009 Mercury
D010 Selenium
D011 Silver
F006 Hastewater treatment sludges from electroplating Poor
F007 Spent bath solutions from electroplating Poor
F008 Plating bath sludge from electroplating Poor
F009 Spent stripping and cleaning solutions from electroplating Poor
F010 Sludge from metal treating Poor
F01I Cleaning solutions from metal treating • Poor'
F012 WW trtmt. sludge from metal treating Poor
F019 Conversion coating sludge from metal treating Poor
K002 Chromium pigment production sludges Poor
K003 Chromium pigment production sludges Poor
K004 Chromium pigment production sludges Poor
K005 Chromium pigment production sludges Poor
K006 Chromium pigment production sludges Poor
K007 Chromium pigment production sludges . . Poor
K008 Chromium'pigment production sludges Poor
K021 Spent antimony catalyst Poor
K031 Cacodylic acid production by-products Poor
K046 Sludge from lead detonator production Poor
K053 Chromium trimmings from leather tanning ' Poor
K054 Chromium trimmings from leather tanning Poor
K055 Buffing dust from leather tanning Poor
K056 Screenings from leather tanning Poor
K057 WW trtmt sludges from leather tannning Poor
K058 WU trtmt sludges from leather tanning Poor
K059 WW trtmt sludges from leather tanning Poor
K060 Lime sludge from coking operations Poor
K061 Furnace dusts • Poor
K062 Spent pickle liquor Poor
KO&3 Lime treatment sludge from steel finishing Poor
KO&5 Surface impoundment solids from primary lead smelting Poor
K066 WW trtmt sludge from primary zinc production Poor
K067 Electrolytic anode sludge from primary zinc production ' Poor
K068 Cadmium plant leach residue from lead smelting - Poor
KO&9 Lead smelting dusts Poor
(continued)
12-7
-------
-------
TABLE L2.1.1 (continued)
Waste
code
K071
KQ84
R086
K.087
KlOO
R101
K102
K106
P006
P010
POli
P012
P015
P036
P038
P065
P073
P074
P087
P092
P103
P107
PiiO
P113
P114
P115
PU9
P120
P122
U032
U136
U139
U144
U145
U151
U204
U205
U214
U215
U216
U217
Description
Muds from mercury chloroalkali cell
Organarsenic production wask
Slude Erom ink and pigment manufacturing
Tars, sludges from coking operations
Lead processing leachate
Organoarsenic production waste
Organoareenic production waste
Mercury chloroalkali cell sludge
Aluminum phosphide
Arsenic acid
Arsenic pentoxide
Arsenic trioxide
Beryllium dust
Dichlorophenylarsine
Diethylarsine
Mercuric fulminate
Nickel carbonyl
Nickel carbide
Osmium tetroxide
Phenyl mercuric acetate
Selenourea
Strontium sulfide
Tetraethyl lead
Thallic oxide
Thalloue selenite
Thallous sulfate
Ammonium vanadate
Vanadium pentoxide
Zinc phosphide
Calcium chromate
Cacodylie acid
Iron dextran
Lead acetate
Lead phosphate
Mercury
Selenoua acid
Selenium disulfide
Thallium acetate
Thallium carbonate
Thallium chloride
Thallium nitrate
Rating3
Poor
Poor
Poor
Poor
Poor
Poor
Poor
Poor
Poor
Poor
Poor
Poor
Poor
Poor
Low
Poor
Poor
Poor
Poor
Poor
Poor
Poor
Poor
Poor
Poor
Poor
Poor
Poor
Poor
Poor
Poor
Poor
Poor
a Low = Low potential
Poor • Poor potential
Blank = No information provided
12-8
-------
-------
12,1.2 Performance of Incinerators Burning Metaj.-Bearing Hazardous Wastes
Numerous studies have been conducted to assess the effectiveneis of
incinerators in destroying various hazardous wastes. These studies, and the
accompanying available performance data have, however, focused upon the extent
to which organic waste constituents are destroyed (i.e., the destruction and
removal efficiency, or ORE), and thus in the majority of cases the wastes
tested are organic waste streams. While those wastes may contain metals in
very low concentrations, few if any could be conaidered equivalent to the
metal-bearing wastes generally considered here. The available data which show
the effect of toxic heavy metals on hazardous waste incinerator performance
are, therefore, ouite limited.
Although the concentrations of metals in the wastes tested are not
4
generally significant, some valuable conclusions may be drawn from an
evaluation of metals analysis data derived from several of the available
performance tests. Tables 12,1.2 and 12,1,3 present summaries of data
21 22
obtained from studies conducted by MRI and GCA Corporation. These
data indicate the fate of metals introduced with the waste feeds, showing
their resultant concentrations in stack emissions and effluents (which include
incinerator ash, and control system effluents such as wet scrubber sludges).
The data, while obviously not conclusive, suggest that many metals are
retained as bottom ash and that the amount of metal in the fly ash which is
not caught by the air pollution control systems and ie thus emitted from the
incinerator stack may be significant (i.e., as high as 10 percent or more).
These phenomenon appears to be related to both the concentration of metal in
the waste feed and the volatility of the metal species. The ratio of
emissions to input were higher for lead, for example, than other metals.
While not shown in the table, it was concluded that in all cases, the
organics under study were destroyed well beyond the required limits. On the
basis of these studies, therefore, it appears that incineration can be an
effective means of managing certain metal—bearing wastes, particularly in its
ability to significantly reduce the overall waste volume to be handled, to
convert much of the other toxic components of such wastes to harmless or less
harmful compounds, and to render a metal-bearing waste more amenable to land
disposal.
12-9
-------
-------
TABLE 12.1.2. SUMMARY OF METALS ANALYSIS DATA - MASS INPUT AND OUTPUT RATES (G/MIN) ,
FOR TOXIC HEAVY METALS STUDY
to
I
Metals
Facility
Rossb
Plant B
American
Cyanamid1'
DuPontb»d
TWld.f
Stream Arsenic
Input
Effluent from absorber
Emissions from stack
Input 0.90
Control device ash 0.03
Emissions 0.035
Inputc 0.024
Emissions
Input
Emissions11 0.049
Input 0.100
Emissions 0.020
Cadmium
5.04
23.4
0.071
0.11
0.04
0.069
0.002
0.005
0.007e
0.007
0.141
0.044
Chromium
20.1
14.4
0.187
0.84
0.1
0.47
0.555
0.188
0.016e
0.007e
0.234
0.012
Mercury
0.33
56.7
0.004
0.50
0.06
0.037
0.009
0.064
0.01
0.020
Nickel
19.7
6.08
0.032
0.16
0.03
0.32
0.377
0.188
0.155
0.032
0.050
0.004
Lead
71.7
248.0
6.93
2.1
0.9
1.5
0.002
0.016
0.117
0.146
2.51
0.810
Selenium Thallium
0.65
113.4
0.011
5.2 0.9
5.1 0.05
4.5 0.06
0.022 0.011
0.04
0.17 0.009
0.500 0.040
aWhere no data are shown, no data were available.
"Average for three test runs.
cCalculated from concentration data.
dNo effluent data given.
eOne test run only.
^Average for four test runs.
Source: Reference 21.
-------
-------
TABLE 12.1.3. CONCENTRATIONS OP METALS IN COMBUSTIBLE WASTE FEED AND
COMPARISON OF INPUT RATES TO EMISSION RATES
Average concentration Average
in combustible waste feed rate
(ug/g) (mg/min)
Arsenic
Barium
Beryllium
Cadmium
Chromium
Iron
Lead
Mercury
Selenium
19.30
121.00
4.67
4.06
166.00
20,800.00
458.00
0.52
0.50
88.90
558.00
21.50
18.70
765.00
95,900.00
2,110.00
2.40
2.30
Average
emission rate
(mg/oinj
40.10
56.20
1.31
23.70
34.20
5,370.00
2,340.00
0.02
0.82
Ratio of emission
to input frora
combustible waste
0.45
0.10
0.06
1.30
0.05
0.06
1.10
0.01
Source: Reference 22.
12-11
-------
-------
12.1,3. Coats .
The overall costs associated with hazardous waste incineration, whether
onsite or at commercial incineration facilities, are high relative to other
hazardous waste treatment or disposal methods. Incineration facilities
require large capital costs due to the size and complexity of the systems
involved, and the requirements associated uich the handling of hazardous
wastes and their combustion products. Operating costs are high due to the
large energy input reauired, and also due to high raw materials and
environmental control costs. Incineration costs are' difficult to specify in
general, because in each situation the number of factors impacting costs is
large. These factors may be classified fundamentally as follows:
* Haste characteristics;
• Facility design characteristics;
,* Operational characteristics.
A 'detailed discussion of, these factors is.presented in Reference 3,,
Costs for commercial incineration of metal-bearing hazardous wastes were
obtained by Alliance within a survey conducted for commercial
incinerators. In general, it can be stated that incineration costs are
higher and in certain cases 'much higher depending upon the type of metals
involved and the metal concentrations involved (see Table 12.1.4). It is
useful to note that the costs of incinerating; such wastes at a commercial
facility employing a cement kiln ere significantly lower than those for
standard incinerators.
12.1.4 Overall Process Status
As discussed in leference 3, there are a number of companies actively
involved in the development, manufacture, and installation of hazardous waste
incineration systems. There are also numerous commercial facilities which
operate hazardous waste incinerators capable of handling wastes containing
toxic heavy -metals, up to certain limits. A telephone survey .of several such
12-12
-------
-------
TABLE 12,1.4.
SUMMARY OF' INFORMATION OBTAINED IN ALLIANCE SURVEY OF
COMMERCIAL HAZARDOUS WASTE INCINERATORS
Type o£
incinerator
State
Types-of metal
wastes handled
Limitations
Bed
Colorado
Liquid Injection/ Illinois
Rotary Kiln
Organic wastes
(e.g., refinery
wastes), Paint
sludges, waste oil.
Blend all wastes.
Any except those
containing cyanides.
llend mil wastes.
Liquid Injection New Jersey Organic wastes
'• "" • "" "''• "' ' -"~" --'- --{no iiYQrganIc"3""'"v
Ko cyanides
Blend all wastes.
Cement Kiln
Ohio
Any- No cyanides
Blend all wastes.
Could not specify.
Hecsl concentration
will be directly
proportional to cost
increase factor.
Based on metal
concentrations.
Tvo examples:
Lead « iOO-600 ppm
Chromium " 1000 ppra
Also somewhat based
on volatility of metals
Ee*g*> metal hydrides
are more volatile, and
they may not choose co
handle them due to
potential air emission).
Will burn vastes con-
taining cyanides up tc-
2000 ppn CN,
Only limitations or.
lead and mercury, ss
shown : Lead *
if concentration is
over IOC ppm,
container size is
limited to 50 Ibs/
container.
Mercury "if cone, is
over 100 ppm, awsi limit
to 300 mg/eontainer.
Based or metal concen- Could not specify.
Typical cost is
i85/5 Ib for
lead-hearing sludges.
Costs are proportional
to tnetal concentration
Chromium * 100-200 ppffl
Lead • 100-200 ppa
Mercury " !> ng/kg
Cadmium " 5 mg/kg
Lirai tacions exist on
wastes containing Pb ,
DSr, Zn, Cr (most
prevalent, also have
linits on others).
General limit to these
species is 4000 ppm
Can handle up to
10,000 ppsi and rcduc€ by
blending. Only bulk
wastes bur ued
Based on netal
concentrator.
Typical costs as
follows:
Base * S20/gal
4-600 ppm = 425-35/gal'
6-10,090 pea »
S35-45/gal
11-20.
12-13
-------
I
facilities nationwide was conducted to determine the prevalence of
incineration of metal-bearing wastes and the associated incineration
costs. TBe results from four facilities in this survey are presented •
in Table 12.1.4.
Overall, it may be concluded that incineration appears to be a limited
and potentially costly alternative for the treatment of hazardous wastes
containing heavy metals. The wastes which may be handled in this manner are
limited to organic wastes (including, orgatiometallic compounds such as cyanides
and tetraethyl lead) which contain metals in fairly low-concentrations. Most
commercial incineration facilities will handle such wastes, but will charge a
premium based on metals concentration,
12.2 PYRQMETALLURGICAL PROCESSES
12.2.1 Process Descriptions
Most of the pyroraetallurgical processes identified .for metal waste
treatment are classified as "calcination" or "smelting" operations.
Calcination processes are generally those which form metal oxides, 'while
smeiVing" produces pure "me'ta*l'."'"'Th"e" reTirt'ionsn'rp^o'f"'"these "processes with'ity'th'e-
overall realm of the pyrometallurgical processes may be depicted as shown in
Figure 12.2.1. The first step shown in the figure eliminates volatiles from
the waste matrix. The oxidation step is similar to incineration, where
combustible materials such as organics will be eliminated. From this point
on, the various available pyrometallurgical technologies are numerous, and, in
some cases, quite different from one another.
Calcination is essentially comprised of Steps 1 and 2 of the flow process
depicted in Figure 12.2.1. The product or products of calcination are metal
oxides. Metal oxides may be separated for use by further chemical processing
or may be disposed of through encapsulation. Generally, if the purpose of
treating the metal waste stream is to render it more amenable to land
disposal, calcination is conducted. Smelting essentially involves Steps 3
and 4 of the flow diagram. The product of a smelter is the pure metal. The
feed to a smelter may be the metal oxides formed through calcination of metal
12-14
-------
STEP 1
STEP 2
STEP 3
STEP 4
CMSTAL WASTES)
APPLICATION OF HEAT
WASTE MATERIALS ARE
HEATED TO A POINT
BELOW THE TEMPERATURE
OF FUSION. VOLATILES
ARE' DRIVEN OFF.
OXIDATION
OXIDATION OF METALS,
OTHER COMPONENTS ENSUES
VIA REACTION WITH AIR,
FLUXING AGENTS.
REDUCTION
REDUCTION OF METAL OXIDE
VIA REACTION WITH REDUCING
AGENT TO--FORM "PUR! 'METAL'- "
PHYSICAL SEPARATION
METAL IS EXTRACTED
THROUGH VARIOUS MEANS
(e.g., ELSCTRQLYS'TS)
HEAT ENERGY INPUT
VOLATILES, WATER
DRIVEN OFF
-FLUXING AGENT
INPUT
(e.g., LIMESTONE)
REDUCING AGENT
INPUT
(e.g., COKE)
OTHER WASTE BY-PRODUCTS
RECOVERED METAL
Figure 12.2.1. Flow diagram for pyrometallurgical processes.
12-15
-------
wastes, or wastes containing metal oxides. In many smelting operations, coke
is employed as a reducing agent. The most common smelting operations involve
recovery of iron, lead, or copper.
Usually, the waste matrix cannot be reduced to the metal in a single
operation and a preparation process is needed to modify the physical or
chemical properties of the raw material. Furthermore, most pyrometallurgieal
reductions do not yield a pure metal and an additional step, refining, is
needed to achieve the chemical purity Chat is specified for the commercial use
24
of the metal.
Drying and calcination are usually carried out in various types of kilns '
such as rotary kilns, shaft furnaces, and rotary hearths. - Smelting1 operations
are conducted in blast or reverberatary furnaces as described in reports and
, ,_ - - 23™" 25
texts dealing with metal processing. Many nonferrous metals can be
extracted by reduction smelting: copper, tin, nickel, cobalt, silver,
antimony, bismuth, and others. Blast furnaces are sometimes used for -the
smelting of copper or tin, but reverberatary furnaces are more common for most
metals.
One of the newer pyrometallurgical processes to be developed is one which
employs the ultra-high temperatures of a plasma arc furnace. Waste dusts from
furnace operations may be fed to a plasma burner operating at temperatures as
high as 500Q°C. The high heat will pyrolyze (break apart) the molecules of
the waste mixture. Recovery may then be effected through selectively
precipitating metals at their appropriate condensation points. This is a
proposed method for handling solid, metallic wastes, particularly those in
which a variety of metals are contained, such as dusts from specialty
steelmaking furnaces. A flow diagram of one such system, the PLASMADUST
process developed'by SKF Industries, is shown in Figure 12.2.2.
Plasma Energy Corporation (Raleigh, N.C.) is testing and demonstrating a
plasma ladle—reheating system for maintaining or increasing the temperature of
molten steel drawn from basic oxygen or electric arc furnaces. The plasma
maintains the necessary temperature for vertical and horizontal continuous
casting. The company is testing prototype plasma systems for recovering
precious metals from automobile catalysts and electronic scrap, for making
fused quartz and superalloys, and for destroying PCBs. It also is conducting
refuse conversion tests for the Canadian government.
12-16
-------
REFUSE
M
I
c
GASIFICATION
SHAFT
CRACKLING
REACTOR
HEAT
EXCHANGER
SLAG
PLASMA
GENERATOR COOLING
AIR t WATER
ELECTRIC HOT DUST
ENERGY WATER ;
OR •
STEAM
COOLING
WATER
MLI tK
. L,
/
J
F
C
4
UEL GAS
ONDENSER
0°C
Od
WATER
TREATMENT
[ ^
=—
Tf
SOLID
WASTE
GAS
WATER
CLEAN
WATER
Figure 12.2.2. Pilot waste destruction system.
Source: Reference 28.
-------
12.2,2 Process Performance . . '
*
Because- the-potential for recovery of metals and other valuable waste
constituents often constitutes the primary incentive for selection of a
pyrotoetallurgical process, it is most meaningful to evaluate their performance
on that basis. While the amount of available performance data for such
systems are limited, several studies of pyrometallurgical process performance
2 A—o A
were reported in the literature. The results of three such studies
have been summarized below. Overall results are as follows;
The three test studies represent both a variety of metal-bearing
wastes and technologies. The different metal -bearing hazardous
wastes tested include electroplating bath sludges, metal
manufacturing sludges, and furnace dusts from specialty steel-making
operations. The results presented may be ske«ed somewhat, however,
due to the face that only a limited number of the metals were
represented (i.e., only the recovery of chromium, nickel, and lead
were shown) . Both bench-scale and full-scale tests were conducted.
In all cases, the percentage of metal recovered was high, ranging .
from approximately 70 percent to 99 percent. In most cases, this
represented recovery as essentially pure metal.' In certain cases,
material. Several studies were referenced indicating high recovery
of other waste constituents, primarily, acids from treatment of
plating wastes.
» Strong dependency was exhibited between metal recovery yields and
operating parameters. , •
Summary of Performance Test Results: Recovery of Chromium and Nickel from
Speciality Steelmaking, Other Wastes — ,
A series of tests were conducted by the U.S. Bureau of Mines to
assess the recovery of chromium and nickel, and other metalsj from specialty
steelmaking dusts (from processes, such as stainless steelmaking}. The
smelting process was also tested for the recovery of other metal-bearing
wastes such as sludges from electroplating.
Four bench-scale test series were conducted and the results obtained were
as follows:
12-18
-------
In the first series, waste feed containing 15 percent ADD dust,
20 percent EF dust, 20 percent- mill scale, and 40 percent grindings
dust were smelted at different temperatures (ranging from 2850° to -
3050°F), for different times (20 to-30 minutes), and with different
chromium oxide reduetants. The results showed high level recovery
of chromium and nickel (i.e., greater than 85 percent) in all
cases. The results also indicated that the recovery of chromium and
nickel is directly related to temperature, time-, and amount of
reductant (as well as to the type of reductant).
la the second series,1 feed composition was varied between a "low
mill scale" grade and a "high mill scale" grade, which employed
twice as much mill scale. Pelletizing tests indicated an optimal
value of 35 percent mill scale, although pellets containing up to
55 percent mill scale could be produced. Smelting results showed
high metal recovery [from 82 to approximately 100 percent) for both
grades. The recovery was greater, however, for the "low mill scale"
grade.
In the third series, smelting of pellets made from- another specialty
steelmaking dust was conducted. This material contained several
other types of isetals in addition to those found in the previous
tests. The results showed that metal recovery was still high for
this material, although not as high as in previous tests (i.e., 60
to 90 percent as opposed to 80 to S9 percent).
In the fourth series, the smelting process was applied to several
chromium and ..nickel-bearing s.ludg
-------
Summary of Performance Test Results: Pilot Testing of Che PLASMABUST Process—
SKF Steel Engineering has a plasma pilot plant at Hofors, Sweden, where
many different kinds of waste oxides have been tested with very promising
results. The pilot plant used 1.5 MW plasma generator for heat supply.
Pneumatic charging systems are used for the raw materials, coal, and the slag
formers. The plant is equipped with a commercially used splash-type zinc
condenser and a venturi scrubber.
All tescs in the pilot plant have shown that the valuable metale in
typical waste oxides from steeltnakine and other secondary materials can be
recovered with a high yield (96 to 99 percent) in the PLASMADUST process,
n 0
They also report that the process is stable and simple to control.
12.2,3 Costs
The economics of using pyrometallurgical processes to treat metal-bearing
hazardous wastes represent the most significant potential drawback to their
attractiveness. The systems involved are highly capital-intensive, as.
large-scale furnaces and attendant systems are used. Typically, capita^ costs,
• ' ' K_ '-;.'' •*»•** i: •>\l't-i-"r"i
-------
TABLE 12.2.1. PRICES OF METALS, METAL OXIDES
Substance
Arsenious Trioxide
Cadmium Metal
Chromium Oxide
Lead Metal
Lead Dioxide
Mercury Metal (precipitate)
Mercury Oxide
Nickel Metal
Nickel Oxide
Selenium
Thallium
' Price (S/lb)
0.42
1.20
1.90
0.28
0.66
7.89
,7,00 ..
3.45
2.60.
13.00 •
35.00
Source: Chemical Marketing Reporter, March 1987.'
12-21
-------
process design and operation. One commercial waste processor known to be
involved in this type of treatment was contacted, but was unable to give
specifics on costs without clear definition of the nature of wastes to be
handled.' The operating costs of such a facility were estimated by Higley,
et al., at $78/ton for the specialty steel-making dusts tested in their study.
In that study, the value of the recovered materials was also estimated to be
9 ft
$280/ton; thus, a resultant economic gain of $0.1Q/lb could be realized.
Arr economic evaluation of a similar thermal system was also presented in
Reference 30. Economic data from this evaluation is presented in
Table 12.2*2. This system was designed to recover acids, but as described
realizes some value from the recovery of iron oxide.
Overall, the key element in evaluating the economic attractiveness of
pyrotnetallurgical systems is the value which may be derived from recovery of
metals. However, systems which can not produce reusable materials nay be
attractive in terms of providing good volumetric reduction of wastes, but tnay
not be viable economically.
J.2.. 2L.J-K. Proce&s^g tatus., •-,.-,;».6j
Usage of pyrotnetallurgical processes for treatment of metal-bearing
hazardous wastes is not well established. Commercial waste processing by
pyrometallurgical processes is not extensive, based on a survey conducted of
waste processors. Due to the potential economic benefits associated with '
metals recovery, souse furnace dust wastes are now being recovered. A number
of examples of facilities where pyrometallurgical systems have been
23
implemented were described by Franklin Associates.
12.3 'EVAPORATION
12.3.1 Process Description
Evaporation is a common unit operation used in the chemical process
industry to separate materials on the basis of their relative volatilities.
In the metal finishing and electroplating industry, evaporation, is used to
concentrate and recover plating solutions, chromic acid, nitric
12-22
-------
TABLE 12.2.2. ECONOMIC EVALUATION OP HYDROCHLORIC ACID REGENERATION
USING THERMAL DECOMPOSITION
Item
Capital Costs
Operating Costs
Labor
Maintenance
Fuel
Electricity
Water
Cost Savings
Acid value
Iron oxide value
Treatment and
disposal costs
Cost basis
TIC* $3
1.82 of TIC
3% of TIC
12,000 Btu/gal /waste
0.10 kwh/gal waste
1 gal/gal waste
Total
, 50% of.PMV** . -
£100/ton
Caustic soda
10,000
,907,000
69,000
120,000
59,000
11,000
6,000
$265,000
. 4-5.7., 000-.^
"187,500
12,000
System size (
100,000
$14,974,000
.- 266,000
460,000
225,000
40,000
24,000
SI, 015, 000
1,875,000
120,000
«pd)
200,000
$23,487,000
427,000
720,000
363,000
65,000
39,000
$1,614,000
-•L '9.;- 133,, 000= ;
3,750,000
240,000
Total 6656,500
Net Annual Savings^ Savings - operating $391,500
Payback Period
16,565,000 $13,123,000
$5,550,000 411,509,000.
(years)
Capital-Net savings
10.0
2.7
2.0
*TIC = Total installed cost.
**PMV = Present market value.
Source: Reference 30.
12-23
-------
acid/hydrofluoril acid pickling liquors, and metal cyanides from spent baths
and rinsewaters. It is also commonly used.as a post—treatment following
reverse osmosis to concentrate metal solutions to the levels needed for
replenishment of the plating bath. Fractionation of volatile components in a
distillation column is a procedure that is seldom required when treating these
metal-bearing wastes. The usual purpose is to achieve satisfactory
concentration levels by evaporation of water from the aqueous1 solution. An
extensive discussion of evaporation/distillation processes can be found in
Reference 3.
12,3.2 Process Performance
The available "data describing the performance of systems used for -the
recovery/treatment oE metal-bearing hazardous wastes are limited to studies of
the treatment of pickle liquors, where acids are also recovered. Despite the
widespread usage of evaporation/distillation in the recovery of volatile
metals, no detailed studies of performance were described in the literature.
Numerous references , , howe.xe_r,>llndicate4-,r,,fcp.r,=!exampl-e,,_;,tJ]at,»recoveryt-p-f-iasi.Tnucht
as 95 percent of pure mercury is regularly achieved through the distillation
24,32
process,
The results of several studies of the distillation of pickle liauor •
wastes were discussed extensively in Reference 30. Summaries of these tests
are presented in Tables 12.3,1 to 12.3.3, and the results are briefly
summarized below:
• High percentage levels of acids and metals (metal salts) were
recovered;
* Low temperature operation was maintained (thus reducing energy
demand/cost);
* Environmental impacts were negligible;
m Waste volume was reduced significantly.
12-24
-------
TABLE 12.3.1. SUMMARY OF PERFORMANCE TEST: SUPERIOR PLATING INC.
Parameter
Result
Purpose of Test: Recovery of cadmium, sodium cyanide from a cadmium cyanide
plating solution
Heat Pump Capacity
Heat Pump Exit Temperature
Evaporator Capacity
Evaporation Temperature
Chiller Exit Temperature
Freon.L Cpndens,or.., .Ex-i-
Coefficient of Performance (COP)
Recovered Cadmium Cone.
Recovered Sodium Cyanide Cone.
300,000 BTU/hr
125°F
200 - 250 Btu/lb
water distilled
110DF
95°F
4.35
2.14 oz./gal.
15,3 oz./gal.
Source: Reference 30.
12-25
-------
TABLE 12.3,2, SUMMARY OF PERFORMANCE TEST: A SINGLE-STAGE EVAPORATOR
AT A CHINESE STEEL PLANT
Concentration (g/1)
Volume • •
Item (liters) H+ F- KOj Fe+2 Ni+2 Cr+3
Purpose of Test: Recovery of acid and tnetal from pickling li
-------
TABLE 12.-3.3. SUMMARY OF PERFORMANCE TEST: TESTING OF HIGH VACUUM VAPOR
COMPRESSION EVAPORATION AT THE CHARLESTON NAVY YARD
Parameter
Result
Purpose of Test: Recovery of chromic acid from a hard chromium plating
line rinse
Compressor Efficiency
Coefficient of Performance (COP)
Adiabatic Efficiency
Capacity
Total Chrome Recovered
(70 gals x 54,900 mg/l)/7484
Dragout Rate ' •
(32.1 lb/320 hrs/month)
x (1 gal/2 Ib Cr+6)
Rinse Ratio
(Ratio of plating bath concentration
to final rinse.concentration using
3 c'ountercurrent rinse tanks)
Rinse Flow Rate
Evaporator Capacity
(Required Rinse Rate)
27 gph x 0.05 gph = 1.35
Recovered Process Water
Quantity
Conductance
Operating Temperatures
Electrical Requirements
10.3
25 2
25 gph § 700 rpm speed
40 gph (8 1170 rpm speed
32.1 Ib
(513.5 oz.)
0.05 gal/hr
20,000
2? gph
per 1 Rph dragout
- 1.35 gph
•- 8.75 gph
(33,600 gpy)
10 nunho
95 - 122°F
9 kw
Source: Reference 30.
12-27
-------
12.3.3 Process Costs ' ' -
Capital costs for an evaporative recovery system will vary with the waste
type, waste Quantity, process flow rates, type of heat exchanger'employed, and
system size. Operating costs generally include 1-2 hours labor for system
maintenance and operation (labor requirements will be reduced if Che system is
operated continuously), electrical and fuel energy requirements for heat
supply, taxes and insurance, and depreciation costs. Approximately 10 Ibs of
low pressure steam (15 psig) is required for every..gallon of liquid
33
evaporated. Typical capital equipment costs are shown in-Table'12.3,4.
Evaporation/distillation processes require large amounts of heat energy,
which can make the process quite' costly. However, efficient use of energy
systems can lower these costs significantly. Waste heat from other industrial
processes (diesel generators, incinerators, boilers, and furnaces) within the
plant can be recovered for use in the evaporation/distillation system. The
use of multi-effect evaporators and vapor compression systems can also improve
thermal efficiencies. Cost savings will be realized in reduced neutralization
costs, reduced sludge disposal costs, and reduced purchase requirements for
fresh'bath 'makeup solutions.
12.3.4 Process Status
Evaporation/distillation is one of the oldest recovery techniques,1 and'is
widely used in industry. Over 600 metal waste recovery units are currently in
30 33
operation in the United States. ,* They are most commonly used in metal
finishing and electroplating industries to .recover plating solutions, chromic
acid and other concentrated acids, and metal cyanides. In addition, water
recovered from the evaporation process is of high purity and can be reused in
process waters. The percentage of these units used, in various plating
applications is presented in Table 12.3.5. These systems are most effective
in recovering acids, bases, and metals from rinsewaters. Systems can be
designed cost-effectively with capacities ranging from 20 gpb to 300 Rph.
These system are cost-competitive with conventional neutralization and
disposal technologies. Greater cost savings are realized with larger
operations.
12-28
-------
TABLE 12.3.4. TYPICAL CAPITAL EQUIPMENT COSTS FOR VARIOUS,.EVAPORATION
SYSTEM CAPACITIES
Evaporator
capacity (gph)
20
40
55
120
300
Capital costs-
(*)
25,000
' , 33,800 '
39,199
44,129
115,000
Source; Reference 30.
TABLE 12.3,5. PERCENTAGE BREAKDOWN BY PLATING TYPE OF EVAPORATION
UNITS CURRENTLY IN OPERATION
Plating chemical
Chromium
Chromium Etch
Nickel
Cyanide
Other
Percent of unics
50 -. •
10
20
10
10
Source: Reference 33.
12-29
-------
12.4 . CRYSTALLIZATION
Crystallization ie a recovery technique in which metal contaminants in
spent corrosive solutions are precipitated through temperature reduction and
then are removed by settling or centrifugation. The applicability of
crystallization as & treatment alternative for metal-bearing hazardous wastes
is limited to liquid waste with appropriate solubility characteristics. As
such, crystallization is most applicable to spent acid wastes from pickling,
plating, etching, or other types of metal finishing operations, such as
caustic soda etching of aluminum.
12,4.1 Process Description
The general process employed in crystallization is simple, focusing
primarily on controlled cooling. A typical crystallization process is
depicted in Figure 12.4.1. In this proces_s, pickle solution is pumped
directly to the crystallizer, which is essentially an insulated tank. Cooling
of the solution is effected and crystallization, of metal salts occurs. The-.
ITe -. - • "' ' ' •' •.' .1 & " -
crystals settle to the bottom and in some system, flow out of the
crystallizer. , The crystallization process often is conducted over long
periods of time, e.g., 8 to 16 hours, and involves temperature reductions of
30 to 100°F. Crystallization may be done on a batch or continuous basis.
Eventually all crystals are removed from solution and settle out. Acid is
then recovered by filtration, or centrifugation, or some other physical
separation operation, and the metal salt crystals are collected for disposal,
or in some cases, for further treatment for metals recovery.
The critical operating parameters involved in crystallization processes
include the solubility of the metal salts in solution, waste composition,
process time, and temperature. The process is more efficient when
concentrations and crystallization temperatures are high. Freezing point
characteristics are also a significant consideration, as some waste constituents
(most typically water) way begin to freeze at or before the applicable
crystallization temperature. To counter freezing, pretreatment is often
conducted. In particular, dewatering of wastes may be done through thermal
12-30'
-------
SPENT SULFUR 1C ACID FICKLE L IOUGK
IsJ
I
CHILLED
AIR
Figure 12.4.1. Flow diagram of crystallization system for recovery of sulfuric acid pickling liquor.
Source: Reference 34.
-------
evaporation, eounteircurrent rinsing, chemical treatment, or air agitation,
Processes in which evaporation is conducted prior to-crystallization are
thought of as two-stage systems.
12,4,2 Pro eg SB tg r forma nee
The performance of a crystallization system is typically evaluated on the
basis of percentage acid recovery, percentage metals removal,'overall product
quality (purity), processing time, and overall economics (i.e., recovery value
vs. operating costs). Typical performance data for crystallization are
summarized in Tables 12.4,1 and 12.4.2. In general, metal recoveries are in
the 50 to 90 percent range.
12.4.3 _Proc_e s s_ Co s t s . • '
The costs of such systems are moderate compared to other thermal-based
recovery processes, primarily due to the great simplicity of crystallization
systems. Capital costs may typically include construction of tank-type
evaporation"and crystallization units, refrigeration system,.and connections.
Operating costs are primarily based on disposal, energy, and maintenance.
Cost-effectiveness depends strongly on the value -of the- acid or other
substances recovered. A typical economic profile is presented in^Table 12.4.3,
12.4.4 Proeesjs Sja_tus
Crystallization systems have been applied on a commercial scale,
primarily by generators of large volumes of spent solutions (e.g., iron and
steel plants). There are several different commercially applied processes for
recovery of sulfuric acid from spent pickle liquor. -All processes, however,
rely upon the basic principles of crystallization of iron salts (mainly
ferrous sulfate) from the spent pickle liquor and the addition of enough fresh
sulfuric acid to return the pickling solution to its original acid strength.
These commercial acid recovery systems allow the free sulfuric acid remaining
in the spent pickling solution to be reused. The processes differ in the
methods used to crystallize the ferrous sulfate.
12-32
-------
TABLE 12.4,1. TYPICAL OPERATING PARAMETERS AND RESULTS FOR'SULFURIC ACID
RECOVERY SYSTEM USING CRYSTALLIZATION
Parameter
Result
Optimum iron content in the waste feed
Iron removal efficiency
Acidity losses in recovered acid
Average cycle time
10 to 141
80 to 85%
2 to 32
6 hrs
Source: Reference 35.
TABLE 12.4,2. TYPICAL PERFORMANCE OF A TWO-STAGE CRYSTALLIZATION
SYSTEM FOR THE RECOVERY OF NITRIC-HYDROFLUORIC ACID
Concentration, weight-percent (Ibs/hr)
Parameter • •.
Feed to evaporator
Feed to crystallizer
Condensed vapor
Residue from
crystallizer
Filtrate from
crystallizer
Total concentration
recovered
Total required
additions
, Fe
3.4
(26.5)
6.5
(26.5)
-
25
(20.0)
2.0
.(6.5) '
0.9
(6.5)
-
Cr
1.1
(8.6)
2.1
(8.6)
--
4.6
(3.7)
1.5 •
(4.9)
0.7
(4.9)
-
Ni
1,6
(12.5)
3.1
(12.5)
-
0.8
(0,6)
3.5
(11.9)
1.7
(11.9)
-
NO 3
12.0
(93.6)
22.1
(89.9)
i
(3.7)
6.0
(4.8)
26.0
(15.1)
12.7
(88.8)
(43)
.,.. F
6.0
(46.8)
. 10.1
(41.2)
1.5
(5.6)
30.8
(24.6)
5.1
(16.6)
3.2
(22.2)
(32)
Water
75.9
(592)
56.1
(228,3)
• 97.5"
' (363.7)
32.9'
'(26.3)
61.7
(202,0)
80.8
(565.73
(261.1)
Source; Reference .36.
12-33
-------
TABLE 12.4,3. ECONOHIC EVALUATION OF ACID RECOVERY SYSTEM USING
CRYSTALLIZATION TECHNIQUE
Item
Flow rate (gal/day)
CAPITAL COSTS
Equipment
Tank (2 tanks •
16,000
460,000
40,000
46,000
546,000
27,600
' 2,300
10,000
46,000
85,900
139,400
314,000
16,200
100,000 ,
569,600
483,700
1. 16 yrs
(14 months)
Large unit
(*)
30,000
850,000
75,000
85,000
1,010,000
51,000
- 4 , 250
12,000
85,000
152,250
261,375
588,750
30,375
187,500
1,068,000
915,750
1.14 vr
(14 months)
Source: References 34, 35, 37.
12-34
-------
Copper recovery from- sulfuric acid-hydrogen peroxide pickle liquors is
being used more and more in the U.S. and Europe.- One of the main advantages
of peroxide is the ability to regenerate spent liauors and to recover copper
electrolytically or by crystallization of copper sulfate. Copper recovery
regeneration of the sulfuric acid has been accomplished for many years with
simple sulfixric acid pickling solutions. This method is suitable also for
23
peroxide pickles and low peroxide concentration.
12-35
-------
REFERENCES
1. Versar, Inc. Technical Assessment of Treatment Alternatives for Wastes
Containing Metale and/or Cyanides. Draft Final Report. Versar, Inc.,
Springfield, VA. Prepared for U.S. Environmental Protection Agency,
Office of Solid Waste, Washington, D.C. EPA Contract No. 68-03-3149.
October 31, 1987.
2. Federal Register 1982, 47 27516-35.
3, Breton, M. et. a 1. Technical Resource Document, Treatment Technologies
for Solvent-Containing Wastes. Final Report. GCA Technology Division,
Inc., Bedford, MA. Prepared for U.S. Environmental Protection Agency,
Hazardous Waste Engineering Research Laboratory, Cincinnati, OH. EPA
Contract No. 68-03-3243. August, 1986.
4. Qppelt, E. T. Hazardous Waste Destruction, Environmental Science and
Technology, Vol. 20, No. 4. 1986.
5, U.S. EPA . National Survey of Hazardous Waste Generators and Treatment,
Storage and Disposal Facilities Regulated under RCRA in 1981, U.S.
Government Printing Office Order No. 055000-00239-8. U.S. Environmental
Protection Agency, Office of Solid Waste, Washington, D.C. 1984.
6. MITRE Corp. A Profile .of Existing Hazardous Waste Incineration
Facilities and Manufacturers in the United'States, PB-84-157072, MITRE
CorpcraticS', McLean, VA. Prepared-for U.S .'"Environmental Protection
Agency, Office of Solid Waste, Washington, B.C., 1984.
7. Incineration and Treatment of Hazardous Waste: Proceedings of the 8th
Annual Research Symposium. EPA-600/9-83-003. Article cited:
Frankle, J,, N. Sanders, and G. Volgel, "Profile of the Hazardous Waste
Incinerator Manufacturing Industry";
8. MITRE Corporation. Survey of Hazardous Waste Incinerator Manufacturers,
1981. MITRE Corporation, METREK Division, McLean, VA. 1982.
9. Advanced Environmental Control Technology Research Center. Research
Planning Task Group Study - Thermal Destruction. EPA-6QQ/2-84-025.
Prepared for U.S. Environmental Protection Agency, Industrial Research
Laboratory, Cincinnati, OH. January 1984.
10. IGF Incorporated. RCRA Risk/Cost Policy Model - Phase III Report.
Prepared for U.S. Environmental Protection Agency, Office of Solid Waste3
Washington, D.C. 1984.
11. Young, C. Telephone conversation with M. Kravett, Alliance Technologies
Corporation. Waste-Tech Services, Inc., March 1987,
12. Mullen, D. Telephone conversation with M. Kravett, Alliance Technologies
Corporation. SCA Chemical Services, March 1987.
12-36
-------
13. Anderson, R. Telephone conversation with M. Kravett, Alliance
Technologies Corporation. IT Corporation, March 1987.
14. Warren, P. telephone conversation with M. Kravett, Alliance Technologies
Corporation. Stable* Corporation, March 1987.
15. Garcia, G. Telephone conversation with M. Kravett, Alliance Technologies
Corporation. TWI, Inc., March 1987.
16. Bell, R. Telephone conversation withrM. Kravett, Alliance Technologies
Corporation. SYSTECH Corp., March 1987.
"•\
17. Klotzbach, K. Telephone conversation with M. Kravett, Alliance
Technologies Corporation. Rollins Environmental Services Inc.,
March 1987.
18. Frost, D. Telephone conversation with M. Kravett, Alliance Technologies
Corporation. Rollins Environmental Services, Inc., March 1987,
19. Bush, R. Telephone conversation with M. Kravett, Alliance Technologies
Corporation, IT Corporation, March 1987.
20. Cooper, D. Telephone conversation with M. Kravett, Alliance Technologies
Corporation. Rollins Environmental Services, Inc., March 1987.
21. Trenholm, A., Gorham, P., and G. Sungclaus. Performance Evaluation of
Full Scale Hazardous Waste Incinerators. EPA-60Q/2-84-181. Midwest
Research ,I_ns.titute, Kansas..City,, MO... -. Prepared'.,for-U.S.-..Environmental.:-
Protection Agency, Office of Research and Development, Cincinnati, OH.
November 1984.
22. Hall, R.R., et al. Union Chemical Trial Burn Sampling and Analysis.
Draft Final Report. GCA Technology Division, Bedford, MA. GCA Report
No. GCA-TR-84-22-G. Prepared for Union Chemical Co., Union, ME.
February 1984.
23. Franklin Associates, LTD. Industrial Resource Recovery Practices;
Metals Smelting and Refining (SIC 33). Final Report. Franklin
Associates, LTD, Prairie Village, KS. Prepared for: U.S. Environmental
Protection Agency, Office of Solid Waste, Washington, D.C. EPA Contract
No. 68-01-6000. January, 1983.
24, Duby, P. Extractive Metallurgy. In: Kirk-Othermer Encyclopedia of
Chemical Technology, Third.Edition, Volume 9. John Wiley and Sons,
New York, NY. 1980. pp. 739-767.
25. Shamsuddin, M. Metal Recovery from Scrap and Waste. Journal of
Metals. February 1986. pp. 24-31.
26. Higleys L.W. , Crosby, R.L., and L.A. Neaumeir. In-Plant Recovery of
Stainless and Other Specialty Steelmaking Wastes. Report No. 8724.
U.S. Department of the.Interior. Bureau of Mines, Washington.,--^.C, 1982.
12-37
-------
27, Bolto, B.A., Kotowski, M., and L. Paulowski. Recovery of Chromium from
Plating Hastes. In: ASTM Special Publication No. 851, "Hazardous and
Industrial Waste Management and Testing; Third Symposium,. March 7-10,
1983, Philadelphia, PA". American Society of Testing Methods,
Philadelphia, PA. 1984.
28, Berlitz, H., and S. Erikson. Metal Recovery,from Hazardous Baghouse Dust
using Plasma Technology. SKF Steel Engineering, Inc., Sweden, In:
Proceedings, Second Conference on Management of Municipal, Hazardous, aod
Coal Wastes. DOE/METC-84-34. Miami University. September•1984,.
29. Alben, P. Telephone conversation with M. Kravett, Alliance Technologies
Corp. World Resources Corp. March 1987.
30. Wilk, L. et. al. Technical Resource Document; Treatment TechnoLoEies
for Corrosives-Containing Wastes, Volume II. Final Report. Alliance
Technologies Corporation, Bedford, MA. Prepared for U.S. Environmental
Protection Agency, Hazardous Waste Engineering Research Laboratory,
Cincinnati, OH. EPA Contract No. 68-03-3243. October, 1986.
31. Stephenson, J.B., Hogan, J.C., and H. S. Kaplan, Recycling and Metal
Recovery Technology for Stainless Steel Pickling Liquors. U.S.
Department of the Interior, Bureau of Mines, In; Environmental
Progress, Volume 3, No. 1. February 1984. pp. 50-52.
32. Jones, H.R. Mercury Pollution .Control. Polluciqn.-ControL Review -No. 1;
. ' . ."-Koyes "Data1' Corporation;--Park- Ridge', "NJ ,' ' 1977. : ' " •'•-*••"•
33. Constantine, D, Corning Process Systems, -Corning, NY. Technical Data
Sheet No. RT-1: Rinse Theory. May 12, 1980.
34. Camp, Dresser, and McKee, Inc. Technical Assessment of Treatment
Alternatives for Wastes Containing Corrosives, Camp, Dresser, and McKee,
Inc., Boston, HA. Prepared for; U.S. Environmental Protection Agency,
Office of Solid Waste, Washington D.C. EPA Contract No. 68-01-6403.
September 1984,
35. Luhrs, R. Telephone conversation with L. Milk, CCA Technology Division,
Inc. Acid Recovery Systems, Inc. September 1986.
36. Krepler, A. Apparatus for Recovery Nitric Acid and Hydrofluoric Acid
from Solutions. -U.S. Patent No. 4, 252, 602. Assigned- to Ruthner
Industrieanlagen-Aktiengasellschaft, Vienna, Austria. February 24, 1981.
37. Crown Technology, Inc. Product Literature: Crown Acid Recovery
Systems. Received in cocnnunication with L. Mi lie, 'GCA Technology
Division, Inc. July 1986.
12-38
-------
SECTION 13.0
PHYSICAL TREATMENT PROCESSES FOR CYANIDES
The two treatment processes discussed in this section are based on
physical methods of separation and do not result in destruction of the
contaminants in the waste feed stream. The processes discussed are:
13.1 Ion Exchange
13.2 Flotation/Foam Separation
Both of these processes are used to some extent for the treatment of
cyanide wastes, but differ in their applicability to various types of wastes-
-and.,thei-c ..needv.fox--,p-fetr-ftatment.rand-t=pQ.s>&T-t-if.ea'.tiaeiTt-!,procedure*.: .-However;'- since-
successful physical treatment merely concentrates the free and complexed
cyanides into smaller volume residuals, some sort of secondary .treatment will
be required prior to disposal.
13-1
-------
13.1 ION EXCHANGE
Ion exchange has successfully removed metal cyanide complexes from
plating, coke plant, and gold mill effluents. However, backwashing of the
cyanide complexes from the strong base anion resins has often proved
difficult, resulting in a continual loss of capacity through repeated
2
cycles. This problem has been apparently overcome by the use of weak base
anion resins or using a three-bed ion exchange installation consisting of a
strong acid cation, weak base anion, and a strong base anion resins.
Laboratory experiments and pilot testing programs have demonstrated the
removal of ferrocyanide from synthetic solutions and industrial wastewaters to
4 • • '
below 1 mg/L of total cyanide.
13.1.1 Proces_s Descrj.ptj.oji
Ion exchange, is the process of removing undesirable anions and cations
from a wastewater by bringing the wastewater in contact with a resin that
""• , ..15
.exchanges ..the. ions int^the.wastewater .with*, a^.set 'o!f ^/substitute- dons.. .',.,.. ...„•.. „ ••
There are three principal operating modes for ion exchange systems:
cocurrent fixed-bed, countercurrent fixed-bed, and continuous
countercurrent. Figure 13.1.1 illustrates the three principal operational
modes, while Table l3.1;l contains a comparison summary.
Fixed-bed ion exchange operations require a cylindrical ion'exchange bed,
tanks for solution storage, and pumps. * The choice of materials is
governed by the chemical environment. Continuous ion exchange systems are
much more complex, requiring solids handling equipment, and more intricate
control systems. Table 13.1.2 gives design parameters and a: range of typical
design values for ion exchange.
In selective cyanide removal through ion exchange, free cyanide is often
first ccmpiexed with iron and then contacted with a basic anion exchanger
which is highly selective for ferrocyanide. The ion exchange column removes
, ., 6
cyanides as .follows:
13-2
-------
Street In
i Rxid&td&fedt,
Step
Caumcfeuffent Cantingouj Mode
Oovdirit Jov* Type!
Service ift *~»-
Service Out «.
i-
tnilng Swrtioft
ttint 1ft —*
l Oui
Out
J
Walh To Remove ?ints
Puls« Gf Deration Section
3-+ R«9trtr*ni Out
Figure 13.1-1. Operational modes for ion exchange
Source: Reference 1.
13-3
-------
TABLE 13,1,1, COMPARISON OF ION EXCHANGE OPERATING MODES
Criteria
Cocurrent
fixed bed
Countercurrent
fixed bed
Countercurrent
continuous
•••Capacity for high feed
flow and concentration
Effluent quality
Regenerant and rinse
requirements
Equipment complexity
Least
Fluctuates with
bed exhaustion
Highest
Simplest; can use
manual operation
Equipment for continuous
operation
Relative costs (per
unit volume:
Investment
Operating
Multiple beds,
single regenera-
tion equipment
Least
Middle
High, minor
fluctuations
Somehwat less
than cocurrent
More complex;
automatic -con-
trojls^ J.or ^ ;(i
regeneration
Multiple beds,
single regener
tion equipment
Middle
Highest
High
Least, yields
concentrated
.regeneration
waste
Most complex;
completely
automated .,.
Provides con-
tinuous service
Highest
Highest chemicals Less chemicals, Least chemicals
and labor; highest water, and labor and labor;
resin inventory than cocurrent lowest'resin
inventory
Source: Reference 1.
13-4
-------
TABLE 13.1.2. TON EXCHANGE DESIGN CRITERIA"
Units
Design criteria
Ion exchange operation
Bed height
Wastewater loading rate
Pressure drop
Cycle time
Regenerat ion
Solution flow rate
m
(ft)
bed volume/hour
cm of water/ra
(in. of water/ft)
bed volumes*
bed volumes**
bed _yplumes/hpur
"iit'ef'/sec'/m^ '
(gal/min/ft2)
percent of treated
wastewater (or 10
bed volumes)
hours
liter/sec/m*
(gal/min/£t2) .
1.2 to 1.8
(4 to 6)
7-5 to 20
11
(-8.4)
100' to 150
200 to 250
Total solution volume
Cycle time backwash
(4 to 8)
2.5 to 5
1 to 3
5
(8)
*For one 1.8 m (6 ft) bed.
**For two 1.8'in (6 ft) beds.
Source: Reference 1.
13-5
-------
Cl"j * Fe(CN)*~ = [(Resin-N-R*) + Fe(CN)*~] + xCl"
o ,5xo
where x = 4 in ferrocyanide and x = 3 in ferricysnide. Once the resin is
exhausted, it can be regenerated with aqueous sodium chloride as follows:
((Resin-N-R*) +
-------
TABLE 13.1.3. CHELATTNG AND ANTON EXCHANGERS USED
Ion exchanger Produced by
Amber lite 7RA-45 Rohm Haas, U.S.A.
Amberlite 1RA-93
Amberlite IRA-94S
Amberlite 1RC-50
Amberlite 1RC-84
Amberlite IRO718
Amberlyst A-21
Chelex 100 " Bio-Rad, U.S.A.
Diaion CR-10 Mitsubishi Chemical
Diaion CR-20 Industries, Japan
Diaioa CR-40
Diaion WA-11
Dialog WA-21 '
Duolite A-6 Dia Prosim, France
Duolite A-7
Duolite S-3Q
Duolite S-37
Duolite ES-63
Duolite A-368
Duolite ES-346
Duolite ES-465
Duolite ES-466
Duolite ES-467
Tmac A-20S Imacti, Holland
Imac GT-73
Tmac TMR
Matrix
structure2
S
S
s
M
A
S
S
S
s
A
S
F
F
F.
F
S
S
S
S
s
s
s
s
+ DVB
+ DVB (M)
-*- DVB (M)
+ DVB
+ DVB
+ DVB .
-
+ DVB
+ DVB'
+ DVB
-
I- DVB
+ DVB
+ DVB
+ DVB
- '
+ DVB
+ DVB
+ DVB
+ DVB
+ DVB (M>
+ DVB
Functional
group
-NH2, -NHR, -NR
-NR2
-NR,
carboxyl
carboxyl
iminodiacetic
weak base
iminodiacetic
iminodiacetic
polyamine
po ly ethyl en imine
weak base
weak base , '
-NR2
-NH2, -NHR, -NR2
polyphenol
-NHR, -NR
phosphor, ic
' -NR2
'araidoxime
thiol
irainodiacetic •
aminophosphotiic
-NR2
weak acid type
-SH, -SQ-jH
(continued)
13-7
-------
TABLE 13.1.3 (continued)
Ion exchanger
'Kastel A-101
Lewatit MP-62
Lewatit TP-207
Merck II
Relite 4MS
Relite MG1
"Thiol resin
Wofatit AD-41
Wofatft MC-50 .
Zerolit HXiP
Zerolit 216
Zerolit S-1006
Matrix
Produced by , structure3
Montedison, Italy S
S
' S
Merck, F.R.G.. S
Resindion, Italy S
A
Chemical Industry Works
A.E.j Poland
Veb. Chemiekombinat S
Bittp.rfeld, G.D.R. - S
Permutit, England S
F
S
+ DVB '
+ DVB (M)
+ DVB (M)
+ DVB '
+ DVB ' •
+ DVB ..
-
+ DVB (M)
+ DVB .
+ DVB
+ DVB
Functional
group
-NR2 •'•
weak base
iminod race tic
-NH2, -NHR, -NR2
-NR2 , •
moderate base
-SH
-NR2 .
imincd1" ?c^t i" andareti
-NR2, -NR3OH
phenylcarboxylic
EDTA type
aS + DVB = Copolyrner of styrene with divinylbenzene; •
S -t- DVB (M) = copolyrner of styrene with divinylbenzene of macroporous
matrix structure;
M + DVB = copolymer of methacrylic acid with divinylbenzene;
A + DVB = copolymer of polyacrylic acid with divinylbenzene;
F = polycondensation exchanger of phenol formaldehyde matrix structure;
- = no data available.
'Source: Reference 7. '
13-8
-------
Other pretreatment requirements include flow equalization for waste streams
experiencing flow or pollutant concentration surges and oil separation to
prevent resin fouling.
Waste streams from the ion exchange process include: spent regenerant
solution, wash streams, and solids from the filtering system. Typically,
since both the spent regenerant solution and the wash stream contain cyanides
these streams will require treatment and disposal, although in some cases the
recovered cyanide can be reused or marketed. Solids from the filtering system
12
can generally be land disposed without further treatment. The quantities
of wastes generated will depend on the types and concentrations of
contaminants present in the solution being treated.
13.1.2 Process Performance
t
Numerous studies have been conducted by various researchers into the
effectiveness of ion exchange for treating metal/cyanide-bearing waste
streams. In 1979, Traehtenberg and Murphy described studies on iron cyanide
removal from leachate from a storage dump coi tav-inp. discarded ' linings from \-
aluminum reduction cells. Data from the full-scale treatment system
showed an average reduction from 48 mg/L of ferrocyanide to 0.5 mg/L
(99 percent removal) at a hydraulic loading of 0.13 mL/min/mL resin. However,
no information was provided pertaining to wastewatcr volumes treated before
regeneration or exchange capacities.
Bessent et al. reported on the use of ion exchange for the treatment of
14
coke plant wastewaters. Pilot-scale, glass columns, 6 in. x 6 Et high
with metal headers, were used to simulate the filtration and ion exchange
systems. Initially, sand was used as the media-in the filter column. This
was replaced, following filtration problems, with a media consisting of sand,
anthracite coal, garnet and granite. In the case of the ion exchange system,
early pilot study runs utilized only one resin column. Later pilot study runs
utilised two ion exchange columns operated in an alternating series mode.
Backwash, regenerant and rinse facilities consisted of various sizes of tanks
and containers applicable co the specific operation being performed.
Figure 13,1.2 presents a schematic ,of the cyanide removal pilot .plant.
13-9
-------
NB RINSE
(SfBviCE WATER)
RiW W&STC
RECTCLtO R£G£N6R*MT
FRESN RI5ENERANT
BACKWASH WATER
SPtHJ REOCNEHUKT
Figure 13.1.2. Schematic of pilot-plant cyanide removal system.
Source: Reference 14,
13-10
-------
The coking effluents were first treated with ferrous suifate Co convert
the free cyanide to ferrocyanide. The ferrocyanide was then removed by an ion
exchange process employing a strongly basic ion exchange resin, Awberlite
IRA-958. Based on the pilot-scale studies, they concluded that for a
full-scale application, the average resin loading would be 11 BE CN_ /mL
resin at a nominal concentration breakthrough of 10 mg/L of iron ferro-
cyanide. Breakthrough would occur after 200 bed volumes if the influent
cyanide concentration was 80 mg/L and the hydraulic loading rate was 0.13
mL/min/mL resin,
Table 13.1.4 presents a summary of the test conditions at the pilot plant
while Table 13.1.5 presents a summary of the test results. Total cyanide in
the polishing column effluent was consistently below 10 Dg/L and in most
cases, the free cyanide concentration was less than 2 tng/L. Runs 6 and 7
closely simulated the equipment configuration, hydraulic loading, operational
mode, and performance recommended for a full-scale system. In runs 6 and 7, a
filter (sand, anthracite, garnet, and granite) followed .by two exchange
columns in series was evaluated. The filter replaced a sand filter in an
attempt to: (a) assess performance of an alternative filtration media, and
(b) provide greater protection for the ion exchange resin from solids with
maximum run times between backwashes. • .
Resin capacity to cyanide breakthrough points were calculated for various
ion exchange runs from cyanide breakthrough data. Average resin capacity was
•3
determined to be 17.6 kg cW/M . This approaches the lower limit of
published capacity for Amberlite IRA 958. It should be noted that the runs
with the 32 bed volume (BV)/hr feed rate showed the highest resin capacity of
3
approximately 20.98 kg CN/M . Final hydraulic loading requirements will
depend on desired throughput.
In 1985, Vachon investigated the removal of iron cyanide from synthetic
and actual/gold mill .effluents', using the strongly basic anion exchanger
Amberlite IRA-958. As indicated in Table 13.1.6, ion exchange was
effective in removing iron cyanide to concentrations of less than 3 mg/L. The
general trend observed was that exchange capacity increased under conditions
of increasing cyanide concentration, lower hydraulic loading, and increasing
13-11
-------
-------
TABLE 13.1.4. PILOT STUDY RUN SUMMARY - RUNS 1 THROUGH 7
Purpose
Flou
Run length
FeSO(, dose
An ionic
polymer dose
Cationic
polymer dose
Type of
filtration
Ion exchange
Resin col. 1
Ion exchange
Resin col. 2
Ijj Regeneration
1 mode
I—1
Run la
8 BV/hr
0.6 gpm
400 BV
500 mg/L
None
None
• Sand
Lead/polishing
0.59ft3 resin bed
None
4 BV fresh
BaCL (reg.)
Run 2b
16 BV/hr
1.2 gpra
400 BV
300 mg/L
2 mg/L
None
Sand
Lead/polishing
0.59ft3 resin bed
None
2 BV recycled
(reg.)
2 BV fresh (reg.)
Run 3C
8 BV/hr
0.6 gpm
400 BV
300 mg/L
2 mg/L
None
Sand
Lead/polishing
0.59ft3 resin bed
None
2 BV recycled
(reg.)
2 BV fresh (reg.)
Run 4d
32 BV/hr
2.4 gpm
400 BV
300 mg/L
3 mg/L
15 mg/L
Sand
Lead
0.59ft3 resin bed
Polishing
0.59ft3 resin bed
2 BV recycled
(reg.)
2 BV fresh (reg.)
Run 5e
32 BV/hr
2.4 gpm
400 BV
300 mg/L
3 mg/L
15 mg/L
Send
Polishing
0.59ft3 resin bed
Lead
0.59ft3 resin bed
2 BV recycled
(reg.)
5 BV freeh (reg.)
Run 6f
16 BV/hr
1.2 gpm
400 BV
300 mg/L
3 mg/L
15 mg/L
Qufld-medi fl*1
Lead
0.59ft3 resin bed
Polishing
0.59ft3 resin bed
4 BV fresh (reg.)
Run 78
16 BV/hr
1.2 gpra
400 BV
300 mg/L
3 mg/L
15 mg/L
Quad-media*1
Pol ishing
0.59ft3 resin bed
Lead
0.59ft3 resin bed
None
Development of cyanide breakthrough curve to determine initial cyanide loading to the column.
Determine effect of increasing hydraulic loading relative to run length and cyanide removal.
CVerify breakthrough curve of run 1; determine resin deterioration.
Observe Z-column operation; determine effect of high hydraulic loading on the columns relative to cyanide removal and run length on a virgin resin.
Determine performance of exposed resin columns operated at a high hydrau Lie loading; duplicate run 4 cyanide breakthrough curve.
Evaluate multi-media filtration system and ion exchange system operated at its high recommended hydraulic loading.
^Duplicate cyanide breakthrough curve of run 6; determine evidence of any resin deterioration.
Sand p anthracite, garnet, and granite.
Source: Reference 14.
-------
-------
I
t-1
U)
Run No.
FeSCV,
s t u il y
1-eSO^
scutly
I
1
2
3
It
5
6
7
Site
No.
1
1
1
2
3
4
1
2
3
4
1
3
4
1
3
4
1
2
3
4
1
2
3
4
1
2
3
4
1
2
4
Total
cyanide
(mg/L)
36.0
65.0
53.0
64.0
63.5
1.0
72.5
80.5
70.5
12.5
74.0
55. I
31.5
67.0
70.0
9.0
34.0
68.0
61.9
2.0
38.0
-
54.0
2.5
-
-
62.0
2.0
61.0
-
~*
Free.
cyanide Ammonia
(mg/L) (mg/L)
20.0
48.0
42.0 1,300
-
0
Ob 1,006
1,024
-
Ob
Ob 922
-
35.0
Ob
-
-
-
-
-
2
3.6 1,015
-
-
-
-
-
-
-
-
19.0
-
1,722
Grease
& oil
(mg/L)
_
136
38
47
16
38
28
77
38
31
63
57
24
43
21
7
47
26
20.3
12
-
-
-
-
64
22
13
-
-
-
~
Phenol
(mg/L)
_
450
480
'
-
326
475
-
-
364
-
-
-
-
-
325
-
-
0
1,020
-
-
-
812
-
-
-
-
-
-
142
Th iocynates
(mg/L)
_
430
-
186
164
170
28
255
192
275
-
-
-
-
240
190
-
-
310
230
-
-
-
310
-
-
-
-
-
-
~
Total
suspended
sol ids
(nig/L)
41.0
-
24.0
110.0
14.0
11.0
35.0
196.0
226.0
240.0
42.0
8.0
48.0
35.0
12.0
3.0
18.0
127.0
24.0
21.0
-
25.0
56.0
-
39.0
31 .0
40
2
26
36
2
Total iron
(mg/L)
4.9
4.1
-
-
13.0
12.0
-
55.0
33.0
45.0
6.1
5.8
14.0
3.6
7.5
5.3
5.3
6.0
17.0
0.2
6.1
-
-
1.9
5.4
10.2
11.6
1.03
6.0
-
~
Chemical
oxygen
demand
(mg/L)
_
3,200
3,100
-
_
2,400
2,117
_
_
2,195
4,185
_
5,952
2,880
-
2,680
5,214
-
-
4,547
3,919
-
-
3,298
1,680
-
-
1,416
.-
-
—
aSite I = Ra« phenol pit waste; Site 2 = Clarili^r effluent; Site 3 = Filtration effluent/ion
exchange influent; Site 4 = Final ion exchange effluent.
"Number actually is negative due Co interferences; reported as zero.
Source: UoCetencr. 14.
-------
TABLE 13.1.6. RESULTS OF CONTINUOUS FLOW TESTS ON GOLD MILL
EFFLUENTS USING A STRONG BASE ANION EXCHANGER
Breakthrough point Effective
Hydraulic exchange capacity
Initial loading No. of (mg/mL resin)
Feed [CNT] (mL/inin/mL [CNp,J [CNT] bed
Test solution (mg/L) resin) (mg/£) (mg/L) volumes CNT CNW CNFe Fe CU Zn
1 Ferrocyanide 110 0.41 1.8 1.8 280 30 30 11
2 Ferricyanide 290 0.41 3.0 3.0 120 34 34 13
3 Ferro/Ferri. 210 0.31 2.6 2.6 120 25 25 8.2
cyanides
4 Iron, copper, 470 0.30 03 50 24 11 13 4.7 6.1 8.0
zinc cyanides
5 Raw barren 378 0.30 0.5 64 75 26 20 5.7 2.0 14 l.l
b leed
6A Raw barren 365 O.'iO 0.1 -62 90 27 20 6.8 2.6 17
bleed
7 Haw barren 370 0.2B 2.2 31.6 70 26 20 6.0 1.7 13 2.6
bleed
8 Raw barren 365 0.09 0 109 100 26 18 7.6 2.9 15
bleed
9 Raw barren- 378 1.7 10 245 80 -19-14 -5 -2 -9 -3
bleed
10 Treated 54 0.30 3.3 3.3 280 14 14 4.5
barren blend
11 Tailings Pond 15 0.37 0.4 0.6 1,480 21 -1 20 10
decant -
12 Field test 20 0.38 3.0 16.5 1,840 ,26 6.7 19 7.8 4.3
Source: Reference 15.
-------
pH, However, for. feed solutions containing copper and sine cyanides (teats 4
through 9), the number of bed volumes treated prior to breakthrough was
significantly lowered due to competition for available exchange capacity.
Limited testing was done using cyanide stream from the cyanide .leaching
process to evaluate the resin exchange capacity after tnulticycling. The data
presented in Table 13,1.7 indicates a 25 percent loss of resin capacity for
total cyanide after the first regeneration with subsequent losses of about
1 percent/cycle. This phenomena has been reported by numerous investigators
and presents a continuing problem in using strong base anion exchangers to
remove cyanides. ' ,
Earlier, Union Carbide investigated several anion exchange resins
(16 to 50 mesh) in the chloride form for effectiveness in treating zinc and
zinc cyanide electroplating wastes. The anion exchangers evaluated were:
* Dowex 1 - "A strongly basic anion exchanger; " . " •
• Dowex NC-20771 - A weakly basic anion exchanger;
• •' Amberlite IRA-93 - A weakly basic anion exchanger; and
• Amberlite XE-275 - A macroreticular, weakly basic anion exchange'
resin possessing tertiary amine functionality in a cross-linked
acrylic matrix. ••
Removal of zinc and zinc cyanide (Table 13.1.8) from the electroplating
waste solutions was greater than 97 percent of equilibrium pH for all four of
the anion exchange resins tested. However, as indicated in Table 13.1.9,
relatively large concentrations of stripping solution (NaOH) were required to
regenerate the column except in the case of Amberlite XE-275.
13.1.3 Process Costs
Process costs for ion exchange processes have been provided in detail siri
Section 8.2.
13.1.4 Overall Process Status
Typically, ion exchange for cyanide removal has been applied as a
polishing step to sorb any ferricyanide or other comple-xed cyanide residuals
13-15
-------
TABLE 13.1.7. RESULTS OF MULTICYCLE ION EXCHANGE STUDIES ON RAW BARREN BLEED
I
I—'
cr>
Test
6A
6B
6C
GD
6E
6F
6G
Cycle
1
2
3
4
5
6
7
Initial
[CNT]
(mg/L)
365
365
365
365
365
365
365
Hydraulic
loading
(mL/min/ml
resin)
0
0
0
0
0
0
0
.40
.40
.40
.40
.40
.40
.40
Breakthrough point
' [ GNFe
(mg/L)
0
1
12
8
11
5
11
[CNT]
(mg/L)
62
134
159
144
161
152
159
No. of
bed
volumes
90
90
90
90
90
90
90
CNT
27
21
19
20
18
19
19
Effective
exchange capacity
(mg/mL resin)
CNw
20
14
13
14
13
13
13
CNFe
6.8
6.8
5.8
6.1
5.9
6.4
5.9
Fe
2.6
2.5
2.3
2.3
2.2
2.2
2.2
Cu
17
12
11
12
9.2
11
9.5
Source: Reference 15.
-------
TABLE 13.1.8. SORPTION OF ZINC CYANIDE AND CYANIDE FROM AN
INDUSTRIAL ELECTROPLATING WASTE SOLUTION
BY VARIOUS ANION EXCHANGERS
Percent sorted
AnLon exchanger
XE-275
•
Dowex-1
IRA-93
NC-20771
Equilibrium pH Zn
11.8 7.0' -
10.4 48.0
9.9 72
9.8 87
9.4 94
11.8 97
9.9 98
8.9 - 7.5 99
8.4 - 7.3 91.0
8.4 - 7.6 87
ch>
8.0
53.0
97
97
97
' 97
97
97
97
97
Source: Reference 16.
-------
FABLE 13.1..9. STRIPPING "OF. ZINC CYANIDE FROM VARIOUS JANION ..
' EXCHANGERS AS A FUNCTION OF NaOH CONCENTRATION
Zinc stripped (percent) ' • '. . •.
NaOH, M
10
8
6
4
2
1
0.5
o.i
XE-275
91.
92.
98.
99.
98.
98.
97.
96.
8
1
0
2
8
7
9
7
Dowex-.l
97.
96.
96.
93,
88.
74.
49.
8.
4
9
3
8
8
2
7
9
IRA-93
97.
97,
- 98.
98.
97.
92.
79.
40.
0
7 •
3
2
1
5
6
0
NC-20771.
96
•' 97
; "; ' 97
97
97
-V?*
89
-. 45
.5
.0
.7 ,
.6
.9
".7,
.0
.5 •
Source: Reference 16.
13-18
-------
from oxidation processes such as alkaline chlorination. .The environmental
impact from this technology is that ic concentrates cyanides in the
regeneration step, creating a secondary stream that needs to be treated.
The advantages of this technology are that it has been demonstrated at
both the bench-scale and pilot-scale. The equipment is compact, versatile,
and is generally applicable to many different waste treatment situations. -
Limitations include the high cost of regenerative chemicals and the waste
streams originating from the regeneration process are relatively high in
pollutant concentration. In addition, if more than 25 mg/L of suspended
solids and/or more than 20 mg/L of oil exists in the influent, filtration is
required as pretreatment. Also, the stream to be treated should not contain
any materials that cannot be removed by the backwash operation. Some organic
compounds, particularly aromatics, will be irreversibly adsorbed by the
resins, and this will result in decreased capacity.
13-19
-------
REFERENCES
1. U.S. EPA. Treatability Manual, . Volume III. EPA-600/8-BO-042. July 1980.
2. Kunz, R.E., Casey, J., end.J, Huff. A Review of Cyanide in Refinery
Wastewaters. 3rd. Annual Conference on Treatment and Disposal of
Industrial 'Wastewaters and Residues. 1978.
3. ,J. Ciancia. New Waste Treatment Technology in the Metal Finishing
Industry. Plating 60 (10). 1973.
4, Avery, N.L., and W. Fries. Selective Removal of Cyanide from Industrial
Waste Effluents with Ion-Exchange Resins. Industrial Engineering
Chemical Products Research and Development 14 (2). ,1975. "•
5. Wilkj L. , Palmer, S., and M. Breton, Alliance Technologies Corporation.
Treatment Technologies for Corrosive-Containing Wastes. Contract
No. 68-02-3997. October 1986.
6. Avery, N.L., and W.H. Waits. Ion-Exchange Treatment Process for
Selective Removal of Cyanide. Amber-Hi-Lites, Rohm and Haas Technical
Brochure. 1977.
7. Z. Hubicki. Purification of Nickel Sulfate Using Chelating Ion
Exchangers. Hydrometallurgy, 16. 1986.
8. T.J. Reynolds. Unit Operations in Environmental Engineering. Rheinhold
Publishers, New York,' NY. 1982.
9. Robertson, W.M., Ja.mes, C.E.,, and J.Y. Huang, Recovery and Reuse of
Waste Nitric Acid From an Aluminum Etch Process. 35th. Industrial Waste
Conference, Purdue University. 1981,
10, C. Fontana, Eco-Tech, Ltd. Personal communication with L. Wilk, Alliance
Technologies Corporation. August 21, 1986.
11. U.S. EPA. Control and Treatment Technology for the Metal Finishing
Industry-Ion Exchange. EPA-625/8-81-007. June 1981.
12. C. Fontana, Eco-Tech, Ltd. Personal communication with L. Wilk, Alliance
Technologies .Corporation. August 26j 1986,
13. Trachtenberg, J.J., and M.A. Murphy. -Removal of Iron Cyanide Complexes
from Wastewater Utilizing an Ion Exchange Process. Light Metal. 1979.
14. R.A. Bessent, et al. Removal of Cyanides 'from Coke Plant Wastewaters by
Selective Ion Exchange Results of Pilot Testing Program. 34th.
Industrial Waste Conference, Purdue University. 1979.
13-20
-------
15. D.T. Vachon. Removal of Iron Cyanide From Gold Mill Effluents by Ion
Exchange. Water Science Technology, Vol. 17. 1985.
16. F.L. Moore. Oak Ridge National Laboratory. An Improved Ion Exchange
Resin Method for Removal and Recovery of Zinc Cyanide and Cyanide from
Electroplating Wastes. Journal of Environmental Science, (7). 1976.
13-21
-------
13.2 FLOTATION/FOAM, .SEPARATION •. - ..-•''•
Flotation/foam seps-ration -is the separation- of finely divided solid
particles from a bulk solution by attachment to fine air. bubbles introduced
into the solution. The bubbles contact the suspended solids and bring them
123'
Co Che liquid surface where they are retained as a foam, ' '
The mechanism of bubble attachment is accomplished through the addition
3 4
of,,,a suitable surfactant called a collector. ' . The principles and physical
models used to describe the attachment of contaminant particles- to air—water
interfaces in the presence of a surfactant are well understood -and- have been
previously described in Section 10.4. With respect to cyanide removal through
flotation/foam separation, iron salts are introduced to the wastewater stream
to complex free cyanide and reduce its toxicity. When precipitated with
excess iron, the iron—cyanide complexes can be removed by flotation using a
cationic surfactant (see Table 13.2.1 for list of commonly used flotation
3
surfactants). , ..'.''•
A disadvantage of this process is that flotation like ion exchange is
physical separation technology. ' • Therefore, use of this technology will
result in a low volume, but highly concentrated toxic by-product wastestream.
This wastestream will require either some sort of secondary oxidative
treatment (ozone and ultraviolet radiation, wet air oxidation, etc.) or
solidification/encapsulation prior to land disposal.1
Currently, flotation/foam separation of cyanide bearing .wastestreams- is
still in a preliminary stage of development. Research into possible
applications has been ongoing for over 15 years, but no large'scale commercial
applications have been reported in the literature. Therefore, when "
considering flotation for possible industrial'utilization, it is important to
"note that further research will be needed to determine its applicability to
specific waste streams,
13.2.1 Process Description
The general process equipment used for the•flotation of complexed
cyanides is similar to equipment utilized in the flotation of complexed
2 5-7 - . ' '
metals. * Figure ,13.2.1 illustrates a. simple flotation system used to
13-22
-------
TABLE 13.2.1. TYPICAL FLOTATION SURFACTANTS
Type
Formula3
Charge on the soluble ion
'Sulfhydryl collectors^
xanthate
dithiophosphate
tnonothiocarbamate
thiol (mercaptan)
dixanthogen
thiocarbanilide
ROCSSNa
(RO)2PSSNa
RHNCSOR
RSH
(ROCSS)2
{C6H5NH}2CS
anionic
an ionic
Colloidal electrolythesc
fatty acids and their soaps
alkyl or aryl alkyl sulfonates
alkyl sulfate
primary amine salt
secondary amine salt
quaternary ammonium salt
RCOOH, RCOONa
RS03Na
ROS03Na
RNH3C1
R2NH2C1
RNCCH3)3C1
anionic
anionic
anionic
anionic
anionic
anionic
aR = CH3(CH2)n
''For sulfides, R = Cj - Cj.
cGenerall straight chain C^2 t°j ^18> or a
may be incorporated into the S. group.
Source: Reference 3.
or naphthalene ring
13-23
-------
-------
Collector
Addition
Fe Addition
LJ
Is}
Waste
Water
Uiderflow
Froth
Product
Flotation
Column
Meter
Figure 13.2.1. Flotation system.
Source: Reference 8.
-------
-------
treat cyanide wastestreams. In this system, the solutions to be treated are
Q
conditioned over a given period of time in an agitated conditioning cell.
The resulting slurry is then introduced to the top of a flotation column. Air
is introduced through a sintered glass diffuser at the bottom of the column
and the froth product is removed at the top of the column as overflow.
Important parameters which affect cyanide removal and which must be
experimentally determined prior to full-scale application include: type of
surfactant, conditioning tank retention time, flotation column retention time,
9
air flow rate, feed concentration, and 'feed pH, The type and dosage of
surfactant added are important since at low surfactant dosages the recovery is
impaired because there is not enough surfactant present to react with all the
influent ferrocyanide. At high surfactant dosages the feed becomes emulsified
and restabilized thereby limiting separation efficiencies. Condition tank
retention time will vary with the influent feed and type of conditioning
chemicals used, but generally recovery will increase rapidly with increasing
9
conditioning time until a steady state is reached (10-30 minutes).
Residuals produced by chemical flotation consist primarily of the
cyanide-laden foam which is skimmed or drawn off the top of the reaction
89
vessel/column. ' Post-treatment typically consists of sedimentation and
sludge consolidation. The resulting hazardous sludge or by-product
wastestream must often be treated (e.g., oxidation encapsulation) and then
discharged to the sewer or land disposed depending on the post-treatment
method utilized.
13,2,2 Process Performance
One of the first experimental investigations into flotation of cyanide
bearing wastestreams was performed by Battelle Laboratories in 1971. The
experimental apparatus for the study consisted of a specially designed bench
scale glass flotation cell. In the first series of experiments several
anionic collectors were screened for flotation effectiveness. The compounds
selected consisted of primary, tertiary, or quaternary-ammonium compounds
while the complexing material consisted of either 10.8 ppm cadmium or 5.64 ppm
nickel in an aqueous stream.
13-25
-------
The results of .the first series of tests" ±B shown in Table 13'.2.2. It
was found that the nickel cyanide complexes could be renewed much more
effectively than cadmium cyanide complexes. Subsequent experiments examined
the effect of a quaternary-ammonium compound collector (tetradecylamine) on
iron cyanide solutions at various feed pHs and cyanide concentrations.
The results presented in Table 13.2,3 are nonconclusive, but do show some
general trends. For example, high extractions were obtained only when the
solution was prepared by adding ferrous iron to a basic cyanide solution. In
slightly acid solution, the complex between ferrous or ferric ion and cyanide
did not occur and low extractions were obtained.
Later investigators such as Clarke and Wilson, Bucsh and Lower, and
Szarawara built on this earlier work by further studying the utility of
flotation for treatment of cyanide bearing wastestreams.. Clarke and Wilson
reported that adsorbing colloid flotation could remove'92 percent of available
12
free cyanide at an optimum pH of 5.5. In this technique, the iron cyanide
precipitate was adsorbed onto ferric hydroxide flox using sodium lauryl
sulfate as a collector. Bucsh and Lower used ion flotation to concentrate
ferrocyanide (pH 4-10) using Aliquot 336 as a surfactant. Removals of
approximately 70 percent were achieved for both ferri and ferrocyanide while
free cyanide had only a 28 percent removal. Szarawara reported that upon
addition of ferrous iron to cyanide solutions, the cyanide concentration would
be at a minimum between pH 8 and 9 as a result of the formation of the complex
Fe
-------
TABLE 13.2.2 EXPERIMENTAL DATA ON VARIOUS COLLECTORS FOR FLOTATION OF CADMIUM
CYANIDE AND NICKEL CYANIDE COMPLEXES
Cadmium cyanide runs Nickel cyanide runs
(1)
(2)
(3)
t-'
OJ (/,)
M
-J (5)
(6)
(7)
(a)
Col lucLor Used
Di)d.:cylamii)C IIC1
Tiit i-.idecyl aniint! IICI
llexjtdecy I aiiune 1IC1
N ,N •ninictliyldodccylamine IICI
l)i;cyl t r ime t hylaminon ium bronii de
K t hy llicxadccy 1 J imethyl nnimon ium
lh*x;idccy 1 py r i J in ium chloride
Solutions were mndc bv tlissolv
Amount
used
cc(a)
I,
2
2
4
4
bromide 0
0
.0
.0
.0
.0
.0
.5
.5
ine collector in
ppm CN Indicated Amount ppin CN Indicated
Inltial(b)
10. 0
10.0
10. 0
10.0
10.0
10.0
10.0
isonronanol
Final removal cc'a' Initial'0' Final removal
5.25 47
6.25 37 2.0 10.0 0.75 93
7.50 25
6.75 32 3.0 10.0 0.75 93
',. 25 57 2.0 10.0 0.50 95
7.75(d) 22 0.5 10.0 0.50^d} 95
7.50(d> 25 1.0 10.0 1.00(d) 90
to ,i [>M or 7.
(l>) Initial solutions also contained L0.8 ppm cadmium as cadmium chloride.
( C) ~ 1 n i I i ft I solutions also contained 5.64 ppm nickel as nickel sulTdte.
(.1) Kxcrss i vo foiimi ng occn rrcil dur i ng these runs causing loss of some so 1 ut ion.
Source : Reference 11.
-------
TABLE L3.2.3. FLOTATION DATA'ON IRON CYANIDE SOLUTIONS
Expt.
No-
16A
153 '
16C
17A
17B
17C
17D
18A
-18B
18C
1813
Initial Solution
ppm
10.
10.
10.
10.
10.
10.
10.
10.
10.
10.
10.
CN
0
0
0
0
0
0
0
0
0
0
0
ppm Fe
3.
3.
3.
3.
3.
3.
3,
3.
• 5.
7.
3.
3.
58
58
58
58
58
58
58
58
3?
16
58
58
ferrous
ferrous
ferric
ferrous
ferrous
ferrous
ferrous
ferrous
ferrous
ferrous
ferrous
ferric
pH During! ,
pH<.a) Flotation^3-1
Basic 4.0
Acid 4.3 .
Basic 4.0 ' •
Basic • 8-4
Basic 6.5
Basic . 5.1
Basic 4.0
Basic =4
Basic =4
Basic =4
• Acid -4
Final
Solution
Analyses-, '^) .
- ppm CS - ' -
1
7
6
2
1
2
2
2
1
I
9
,25' '
, 75
, 25 ^ '
.35 .
.95 ,
.10
.65
.70.' -• ; -
.05
.25
.50
Apparent
Percent
Extraction.
'. 87,
•22.
37.
76.
80.
79.
73.
• 73.
89.
• 87.
5,
5
5
5
5
5
0
5
0
5
5
0
(a) Adjustments of pH were made by adding dilute HCI or SaQH.'
(b) Solutions were floaced by adding 0.5 cc of tetradecylamine 'collector
and aerating.~fsr 10 minutes. .. . .
Source: Reference 11. ' •'
13-28
-------
TABLE 13.2.4 FLOTATION"REAGENTS
Reagent name
Manufacturer
M.W.
Formula
ARMAC 12D
Dodecylamine
Ethyl acecate
Sodium Lauryl
Sulface
Aliquat 336
Tricaprylyl Methyl
Atnmoniuin Chloride
4-Methyl '2-Pentanol
Armour Industrial
Chemical Company
A'ldrich Chemical
Company
General Mills
Chemical Division
Unknown
245
288
X = 442
1-2
CB3('CH2)uOS03Na
R3NCH3Cl
R = Cg - CJQ carbon
(CH3)2CHCH2CHOHCH3
TABLE 13.2.5 EFFECT OF TYPE OF COLLECTOR
100
215
Fr =
R, =
I
Run
29
30
31
32
33
34
36
37
mg/L Total
mg/L Total
fraction
Surfactant
NLS
12D
NLS
12D
NLS
12D
NLS
12D
CN
Fe 11
removed
removal factor = 1 - cyanide
cyanide
Level mg/L (mM)
46.2(0.16)
156(0.64)
23.1(0.083
31.3(0.13}
46.2(0.16)'
156(0.64)
23.1(0.08)
31.3(0.13)
in trie underflow
in the feed
PH
4
4 .
4
4
6
6
6
6
Fr
.45
.90
.42
.87
.91
.94
.84
. ' .90 '
Rf
.28
.88
.28
.85
.89
.80
.80
.89
Source: Reference
13-29
-------
from cyanide solutions containing little or no competing ions. For actual
cyanide wastewaters such as coke plant effluents, maximum efficiencies were
reduced to 91 percent. •• ' '
13.2.3' Process Cost
Presently, little cost data has been reported on flotation/foam
separation as a treatment process for the physical removal of cyanides from
wastestreams. While capital and operating costs associated with this
technology are expected to be low, no precise costs have'-been developed, A
primary cost (and environmental liability) anticipated from the use of
flotation/foam separation is for the secondary treatment and final disposal of
the iron-cyanide flotation sludge. Secondary treatment will consist either of
a destruction technology (e.g., wet air oxidation or UV ozonation) which can •
successfully treat iron cyanides or a,solidification/encapsulation technology
which will immobilize the cyanide pollutants contained in the flotation
sludge. The inclusion of these secondary treatment costs are expected to add
significantly to the overall treatment costs,
13.2.4 Overall Process Status
Flotation/foam separation of cyanide bearing wastewaters has not yet been
tested on a pilot-scale at an actual commercial facility. Most of the
research that has been performed to date with flotation has focused on
equipment development and process parameter definition. Although preliminary
research has demonstrated the technical feasibility of the process,
pilot-scale testing is needed to determine if sufficient cyanide recoveries
can be achieved. Flotation could prove to be a cost-effective alternative to
conventional treatment practices because of its minimal operating requirements.
13-30
-------
As with all physical separation processes, any process which will
concentrate the .cyanide waste material should be followed bys"a process which
will detoxify or immobilize the concentrate. Since the concentrate of the
process considered here contains precipitated ferri- and ferrocyanide which
are not amenable to conventional oxidation technologies such as alkaline
chlorination, alternate technologies sych as wet air oxidation or UV/ozonation
may be more appropriate. In addition, solidification or encapsulation of the
residuals may be required prior to land disposal.
13-31
-------
REFERENCES
1. U.S. EPA. Treatability Manual. Volume III. EPA-600/8-80-042. July
1980. ' .
2. Sundstrom, D.W., and H.E. Kiel. Wastewater Treatment.' Prentice-Hall,
Inc., Englewood Cliffs, NJ, 1979,
3. Kirk-Othmer Encyclopedia of Chemical Technology. Vol. 10, 3rd. Edition,
John Wiley & Sons, New York, NY. 1981.
4. R,R. Klimpel, Dow Chemical Company. Use of Chemical Reagents in.
Flotation, Chemical Engineering. September 3, 1982.
5. T.D. Reynolds. Unit'Operations and Processes in Environmental
Engineering. PWS Publishers, Boston, MA. 1982.
6. E.L. Tbackston, et al. Lead Removal with Adsorbing Colloid Flotation.
Journal of the Water Pollution Control Federation. February 1980.
7. Huang, S., and D.J. Wilson. Hexavalent Chromium Removal in a Foam
Flotation Pilot Plant. Separation Science and Technology, 19. 1984.
8. Lower, G.W., and D.J. Spoctiswood, Michigan Technological University.
Cyanide Removal from Coke Making and Blast Furnace Wastewaters.
EPA-600/2-83-066. August 1983.
9. Busch, R.O., Spottiswood,. D.J., and G.W,Lower. ' Ion Precipitate
Flotation of Iron-Cyanide Complexes. Journal of the Water Pollution
Control Federation, December 1980,
10. Currin, B.L., Potter, F.J., and D.J. Wilson. Surfactant Recovery in
Adsorbing Colloid Flotation. Separation Science 13(4). 1978.
11. Battelle Laboratories. An Investigation of Techniques for Removal of
Cyanide from Electroplating Wastes. Water Pollution Control Research
Series. 12010 EIE. November 1971.
12. Clark, A.N., and D.J. Wilson. The Removal- of Mettsllo-Cyanide Complexes
by Foam Flotation. International Conference on Management and Contr-ol. of
Heavy Metals in the Environment. September 1979.
13. Bucsh, R.O., and G.W. Lower. Cyanide Removal from Coke Making and Blast
Furnace Waste Waters. EPA-600/9-81-017. March 1981.
14, J, Ssarawara, et al. Studies of the Equilibria of Complex Formation of
Cyanides with Ferrous Sulphate. Chemia Stosowana, IVT, 79. 1972.
13-32
-------
14.0 CHEMICAL-DESTRUCTION OF CYANIDES
The cyanide destruction processes discussed in this section are based on
chemical methods of separation and destruction o£ cyanide contaminants in the
waste feed stream. These unit processes are: .-
14.1 .Alkaline Chlorination
14.2 Ozonation
14.3 Wet Air Oxidation ' . . •
14.4 Sulfur-Based Technologies
14.5 Miscellaneous Processes
The cyanide waste streams treated by these processes are produced by
several industries--including ore extraction- (cvanide* le'acbing), 'photographic"'''
processing, synthetic organic and inorganic compound manufacturing, and metal
finishing. The most significant source of hazardous cyanide waste is the
metal finishing industry. Aqueous solutions with free cyanide, ionic
cyanides, and highly soluble metal cyanide complexes are of major
environmental concern. Aqueous cyanide waste solutions from the metal
finishing industry include contaminated rinse water and spent process
solutions. -
14.1 ALKALINE CHLORINATION '
14.1.1 Process Description
Alkaline chlorination of dilute cyanide waste streams is a waste
treatment technology which has been in commercial use. for over 25 years.
The process is suitable-for destroying free dissolved hydrogen cyanide and for.
14-1
-------
oxidizing all simpLe;-and moat complex inorganic- cyanides"; in .''aqueous media,
The process is operated at ambient temperature; "with •good,;'-'pH.''and
oxidation-reduction -potential (ORF) control , v-'the-fef f luent:'::eypically contains
" " 1 ") ' ' •' ':' ":' ".-• '• .""!"'"< '."'"7.''"---. '""' .-'*•-.' "•
less than 1.0 ppm cyanide. '.' • ."';"''• " ;'"';;-'---7:l'7;'"r,"v:-l:::l'v"., •
The destruction reaction is an oxidation process in ''which one or more
electrons are transferred from the chemical being oxidized -(cyanide) to the
34'"''-
chemical initiating the transfer (oxidizing agent). ' --..'-Chlorine in
elemental form or hypochlorite salt are the two most "common oxidizing agents
used in industrial cyanide oxidation systems.' '•" ''' •' -. ."" •"'."';',
The mechanism of cyanide destruction by alkaline chlorination is 'shown by
the following equations: '•'--;•'
C12 (g) * NaCM = CNC1 -i- NaCl : '..."'.
CNC1 + 2NaQH = NaCNO + NaCl + H20 . ;. . '',
3C12 (g) + 2 NaCNO +• 6 NaOH = 2NaHC03 +. N2 (g) +'6 Nad + 2H20
In this reaction chlorine gas (C1»J plus sodium hydroxide (NaOH) are
used to oxidize cyanides to cyanates "HCNO ) and ultimately to' carbon dioxide
and nitrogen. The formation of cyanide chloride (CNCl).is essentially
instantaneous. Sodium hypochlorite (NaOCl) is .of ten used in place of chlorine
gas due to the danger and higher equipment costs involved --with chlorine
usage. The stoichiometry is the same in terms of equivalents of chlorine
added, but alkali additions and unit reagent costs (sodium hypochlorite is
approximately -twice as expensive as chlorine gas)- will • vary- with the oxidizing
agent used, - ••'."•• ..'.' ,. ' .-
Alkaline chlorination treatment of cyanide solutions can be conducted in
125 ~ ' ' ' .-•.-'•-'• "' ••
,one or two stages, ' * In the more commonly used two stage process,
solution pH is initially raised to a pH of 10 of .higher-. ^ Hydrolysis of the
cyanogen chloride complex is rapid and the reaction is typically 80-90 percent
complete within two minutes. In the second stage, the pH of the solution is
reduced to the 8.0-8.5 range for rapid oxidation of cyanate. Retention time
in the second stage is generally 30 minutes 'to 1 hour 'in 'order to ensure
complete cyanide destruction. Alternatively ,- an intermediate pH between
8.5 to 10.0 can be maintained in a single tank for. simultaneous completion of
14-2
-------
both stages. In the single stage system close pH control is essential, and-
retention time will depend upon the selection of pH and the amount of
hypochiorite present. Figure 14.1.1 illustrates a conventional two-stage ,
cyanide1oxidation system. The system features separate pH controlled addition
of caustic and ORP controlled addition of chlorine to each stage if
necessary. Table 14.1.1 presents treatment levels for cyanide wastewaters -
using both single and two-stage chlorination processes.
Reagent requirements for the theoretical oxidation of cyanide to cyanate
are 2.7 Ibs. of chlorine and 3.1 Ibs of caustic per pound of cyanide.
Overall reagent requirements for the complete destruction of cyanides are
6.8 Ibs of chlorine and 7.3 Ibs of caustic per pound of cyanide. Practical
experience, however, has demonstrated that typically 8 Ibs of chlorine or more
(a 10 percent excess) are required to completely destroy cyanide and meet
effluent guidelines. The excess chlorine is used to account-, for side
-reactions (organics and reduced metals) and ensure rapid and complete
hydrolysis of cyanogen chloride.
The rate equation for the hydrolysis of cyanogen chloride to cyanate
is:
-d iCNClJ/dt = kl ICNC1J [OHJ
As indicated by the presence of the hydroxyl group fOHj, the rate equation
shows cyanogen hydrolysis to be pH dependent. The greater the concentration
of hydroxyl ions, the more rapid the reaction rate. The reaction has been
found experimentally to be most rapid above pH 10, a region of high alkalinity
(i.e., excess of hydroxyl ions).
Hydrolysis of cyanogen chloride is greatly accelerated by the presence of
hypochlorite, which apparently has a catalytic effect. Competition between
CNC1 and CNO for excess hypochlorite may result 'in incomplete cyanogen
hydrolysis at low pH values. However, at high pH values '(greater Chan
pH 10.0) cyanogen hydrolysis is complete before significant CNO oxidation
7
occurs.
14-3
-------
FLOW
EQUALIZATION
1st STAGE ;
OXIDATION ;!
2nd STAGE
OXIDATION
CAUSTIC CHLORINE
ACID
RAW
WASTE WATER
pH =10.0-11.0
pH =8.0-8.5
TREATED WASTEWATER j
(TO SUBSEQUENT
TREATMENT)
SLUDGE
(WHERE APPLICABLE)
Figure 14.1.1. Treatment flow schematic for 2-stage oxidation process.
-------
TABLE 14.1,1. TREATMENT LEVELS FOR CYANIDE WASTEWATEiS
Cyanide Concentration mfi/L)
Treatment process
Alkaline chlorinationa
Alkaline chlorinat iona
Alkaline chloriaation'5
Alkaline chlorinatioti'5
Alkaline chlorination
Alkaline chlorination
Initial Final
1.
0.
0.
700 0.
32.5 0.
5.1 0.
7
I
4
0
0
1
Percent
Removal
100
100
98
aSingle-sCage cfclorination.
^Two-stage chlorination.
Source: Reference 2.
14-5
-------
-------
"'. -
General processing equipment and construction materials for cyanide
oxidation units are identical to those of precipitation, reduction, and
coagulation/flocculaticm processes. ^A fully engineered two-stage cyanide
fc 128
oxidation system would consist of the following coraponenta: '
* 2 treatment tanks
* 3 reagerft storage tanks (caustic, chlorine, acid)
• 5 agitators •{-,-'•'''•''
• 6 pumps /I" -"*:"
• 2 pH controller/probes '_..'• ! •
• 2 OKP controller/probes '••• •
» piping and valves .'. •
* electrical fit-up
The two treatment tanks can be fabricated from a wide range of
construction materials, but most industrial systems use fiber reinforced
9 10
plastic CFRP). ' host vessels are of a flat-bottomed configuration,
equipped with air tight covers or air ducts to minimize exposure to any
volatile, toxic reaction products which might be evolved. Each stage should
be designed to provide approximately 1-bour retention volume.
Agitation serves the purpose of equalizing the concentration profile
within the reaction vessel as the influent is dispersed in the reaction
tanks. Vessels with large stagnant areas provide little nixing between
reactants and causes large disturbances when concentrated materials are
released into the system. For accurate process control, Hoyle has suggested
that agitator capacity should be measured as a ratio of the system dead time .
(the interval between the addition of a reagent and the first observable
process change) to the retention time (volume of the vessel divided by the
flow through the vessel). A ratio of dead time to retention time of 0.05
approaches an optimum value. Typically agitation is provided overhead in line
with the vertical axis. In addition, mechanical agitation should be provided
in the reagent storage/slurry tanks to maintain reagent homogeneity.
14-6
-------
-------
Pumps and piping are required for all aspects of fluid transfer within
Che cyanide oxidation system. Pumps are accessary to transport the cyanide
waste to the first stage, pump it to the second stage, and then displace the
treated fluid from the second tank to whatever post-treatment processes may be
appropriate. In addition, a separate chemical metering pump is required to
transfer reagent from each of the reagent storage tanks (in smaller system it
is sometimes possible to meter directly from a 55-gallon drum) to the
treatment system. The many different factors influencing the final choice of
pump type and size for fluids are discussed in detail by Peters and Timmerbaus
in Reference 13.
At the heart of the alkaline chlorination cyanide destruction system are
the pH and ORP control systems. The pH control systems for batch
precipitation processes can be quite simple with only on—off control provided
via solenoid or air activated valves. Control system designs for continuous
flow cyanide oxidation system are more complicated because the wastewater
feeds often fluctuate in both flow and concentration. Systems currently
available include: proportional, cascade, feedforward, or feedback pH
control. Each system has distinct advantages and disadvantages which have
been reviewed in the literature. ' ' Both pH and ORP control systems
consis't of a probe (to take the reading), monitor.^Cto .compare-,, the reading set
point and make the appropriate adjustment), and a recorder to visually display
the resultant data. In addition, there is typically a control panel with an
indicator, starters and controls for metering pumps, all relays, high/low
alarms, switches, and mixer motor starters.
Table 14.1.2 summarizes typical operating parameters for a two-stage
alkaline chlorination systems. Improper chlorination of cyanide ion, hydrogen
cyanide, or thiocyanate ion, particularly under conditions below pH 10, will
result in increased evolution of cyanogen chloride, a gas which is considered
to be at least as hazardous as hydrogen cyanide. Cyanide in combination with
nickel, cobalt, silver, or gold is oxidized slowly, but is still treatable if
2
sufficient time is provided.
A pretreataent in itself, alkaline chlorination is usually applied to
cyanide bearing aqueous waste streams segregated from other process
flowstreams. Segregation is essential to prevent the formation of difficult
to treat wastes or the evolution of toxic gases.
14-7
-------
TABLE 14.1,2.
'TYPICAL OPERATING PARAMETERS OF A TWO-STAGE ALKALINE
CHLORINATION CYANIDE DESTRUCTION UNIT ,
Parameter
Unit
Range
Influent
Cyanide concentration
Influent
Flowrate pressure
Pressure
Temperature
Agitation
First-Stage
pH
ORF
Chlorine
Caustic
Retention time
ag/L
atm
"C
turnover/minute
Mv
Ib/lb CN
Ib/lb CN
Min
- l»000a
10 - 350
1
20-22
1
9.5 - 11
350 - 400
2.7 - 3.0
3.1 - 3.4
30 - 60
Second-Stage
pH
OR?
Chlorine
Caustic '
Retention time
Effluent
Cyanide
Mv
Ib/lb CN
Ib/lb CN
8.0 - 8.5
600
4.1 - 4.5
4.2 - 4.6
30 - 60
Bg/L
alnitial cyanide concentrations of up to 5,000 mg/L are possible, but
require batch treatment. Optimum influent cyanide concentrations for
continuous systems are < 100 mg/L.
Source: Adapted from References 1, 2, 6, 9.
-------
For example, alkaline chlorination cannot effectively oxidize stable iron
and nickel cyanide complexes. As most cyanide discharge limits are based on
total cyanide levels, provisions should be made Co ensure that cyanide
solutions do not mix with iron and nickel compounds. Similarly acid-bearing
waste streams should be segregated from the cyanide bearing wastestream to
prevent pH depression and the evolution of toxic hydrogen cyanide gas IHCN).
Following successful cyanide destruction, the treated cyanide wastestream may
then be combined with other waste streams for subsequent treatment (i.e.,
metals precipitation, coagulation, filtration, etc.).
Other properties of the waste being treated that can affect alkaline
chlorination performance include;
• Flo« variations
• pH variations.
• Presence of chelators/complexants
• Competing nonpriority oxidizable species
• Oil and grease concentration
In facilities which experience a wide variation in flow rates, pH values,
or pollutant concentrations of the wastewater, flow equalization as
1 9
pretreatment is often used. ' A variety of process options exist (see
Section 10.1) but all systems basically provide some sort of flow resistance,
stream segregation, or influent concentration averaging to prevent
wastetreatment system overloading. , In all methods of flow equalization, care
must be exercised during the wasteuater analysis to completely characterize
any peak flows or concentrations. In addition, flexibility in system design
should be provided for any future expansion, change in location, or deviation
in flow rates.
Oil and grease, chelator/complexants; and nonpriority oxidizables, are
all factors which will increase reagent consumption and impede if not prohibit
chetnicsl oxidation operations. Oil and grease removal is typically the
first process step in any waste treatment train. The removal of
chelator/cotnplexants and nonpriority oxidizable compounds present more
difficult problems since many of these compounds are often an intesral part of
14-9
-------
the cyanide wastestream. The presence of organic compounds.and reduced metals
can increase chlorine or sodium hypochlorite consumption by as much as 25 to
2
100 percent over stoichiometric requirements. In addition the presence of
cupric cyanide can cause precipitation during the chlorination 'process. This •
results in a sludge containing cyanide complexes that may require separate
post-treatment. Other inorganic salts which cannot be effectively treated by
this process and may require segregation and/or pretreatment include ferro and
ferric-cyanides, nickel cyanide, and zinc cyanide.
Residuals generated in the alkaline chlorination process occur from, the
use of caustic with chlorine gas. Smaller quantities of residual product will
result from alkaline chlorinations using hypochlorites. The sludge product ,
consists• primarily of insoluble hydroxide compounds generated during the
hydrolysis of cyanogen chloride in che first-stage reactor. Therefore some
provision for sludge removal or batch clean-out should be provided. However,
alkaline chlorination post-treatment is more likely to consist of such unit
processes as precipitation, coagulation/flocculation/ sedimentation, and
sludge consolidation. The resulting toxic sludge must often then be treated
(i.e., encapsulation) and land disposed.
14.1.2 Process Performance - •
Alkaline chlorination with chlorine or hypochlorites has become the most
widely accepted conventional method of cyanide destruction. The stoichiometry
and rate factors in cyanide destruction by alkaline chlorination have been
researched -and thoroughly reported in the literature. Use of this method
however becomes increasingly difficult as cyanide and stable- inorganic
salt-cyanide complex concentrations increases.
Table 14.1.3 summarizes effluent cyanide concentrations for 15 metal
Q
finishing plants reviewed in the literature. Total cyanide influent
concentrations ranged from 0.045 to 1,680 mg/L, with a median, of 77.4 mg/L.
As can be seen, alkaline chlorination was successful in reducing 65 percent of
the total cyanide waste streams to a final effluent concentration of less than
0.10 mg/L. However, two of the facilities were unable to detoxify total
cyanide concentrations to less than 1.0 mg/L. If was' postulated that
inefficient operation, the presence of stable inorganic complexes (i.e. iron,
14-10
?^-r?*i:>^ '" '" • _ O.
-------
TABLE 14.1.3. 'EFFLUENT CYANIDE PERFORMANCE DATA USING ALKALINE CHLORINATION
Total Cyanide3 Amenable Cyanide'5. •
mean effluent ' mean effluent
Plant ID concentration (mg/L) concentration (mg/L)
.1
2
3
A
5
6
7
8
9
10
11-
12 .
13
14
15
0.04
0.15
0.09
2.20 .
0.09
0.10
1.21
0.05
-0.001
0.13
0.46
0.04
0.01.
0.06
._ ' .
. :
.
0.09
0.004
• ' . • -
0.007
aAverage daily total cyanide influent concentrations ranged from 0.045 -
• 1,680 mg/L with a median concentration of 77.4 mg/L.
Amenable cyanide influent concentrations ranged'from 0-1,560 mg/L with a
7.63 mg/L median concentration.
Source: Reference 8. -. . . •
14-11
-------
nickel, zinc cyanide), or excessive influent total cyanide.concentrations
t1,000 mg/L or greater) were responsible for the poor removal efficiencies
experienced at these plants.
Table 14.1,4 presents detailed alkaline chlorination operation and
performance data for five more facilities. Facilities A, B, and D are batch
processes used in lieu of continuous alkaline chlorination. * *'
Facility A uses its system to collect and batch treat spent cyanide baths and
floor spills. Therefore, equipment usage is intermittent and process
conditions are variable. Facility D op'erated a batch pilot plant with limited
throughput to determine treatment feasibility. Facility B is a commercial
wastetreatment plant which accepts and treats large volumes of concentrated
cyanides with lime and sodium hypochlorite. The process is limited by two
factors. First the initial content of cyanide (CN ) must not' exceed
2,000 mg/L in order for the process to achieve a final cyanide concentration
0.5 mg/L, Secondly the total amount of Cl_ used should not surpass
6,000 mg/L in order to limit the levels of cyanogen chloride formed during the
process. If either process parameter is exceeded, a dilution operation is
performed.
1 8 20
Facilities C and E are continuous alkaline chlorination operations.'' '
Facility E treats cyanide contaminated ore leaching wastewater generated
durinR gold milling operations (see Figure 14.1.2). The gold in the ore is
mainly locked in fine grained arsenopyrite (FeAsS), but also contains copper
and zinc. The capacity of the treatment plant was 2.5 to 16.8 gpm with tanks
constructed of protected (lined) mild steel and of plastic. Plastic piping
and rubber hose were used for ease of changing flow patterns. Process
operations consisted of oxidation of reduced species (cyanides and arsenites),
alkaline precipitation of metallic hydroxides, ferric sulfate precipitation of
pentavalent arsenic, and liquid-solid separation. The levels achieved are as
follows:
Cyanide 1.0 mg/L
Arsenic 0,2 mg/L
Copper 0.3 mg/L
Zinc 0.2 mg/L-
Iron 3.0'mg/L
14-12
-------
TABLE 14.1.4. ALKALINE CHLORINATION PERFORMANCE DATA
Parameter Aa
Was test ream Metal
Finishing
Influent batch
Flowrate (gpm) (1000 gal)
1st stage
pH
Retention
Time (mio) —
Reagent
2nd Stage
pH NA
QRP(MV) NA
Retention Time (Min) MA
Reagent Sodium
hypochlorite
Sodium
hydroxide
(NaOH)
Influent Total
Cyanide Concentration
(rng/L) 0.5-6.8
Effluent total
Cyanide Concentration 0.1
Bb Cc
Commercial Coke and
wastetreat- coke by-
Facility products
batch 167
11
250-350
60
Lime/
Sodium
hypochlorite
8.5 9.0-9.5
NA 120-180
5-60 90
Waste acid Chlorine
NaOH
2,000 83-104
0.5 4.7
Dd Ee
Gold Gold mill
mill effluent
barren
bleed
batch 2.5-16.8
12 11.2-11.8
90 82
Lime/ Lime/
Sodium Sodium
hypo— hypo—
chlorite chlorite
8.5 7.5-10.4
•
60 100
Sulfuric NA
acid
63 300
0.4 0.07
Reference 16.
''Reference 17.
eReference 18.
^Reference 19.
Reference 20,
14-13
-------
-------
CO!
DARREN Ui^Tt mice
10
UnBtltCKOUMB
STOIUSE
SIR! An)
Figure 14.1.2. Gold processing flow diagram.
Source: Reference 20,
14-14 '
-------
-------
Facility C was one of the few facilities Co report an inability to
achieve effluent limits with alkaline chlorinatiotu High influent ammonia and
thiocyanate concentrations were felt to have reacted with some of the excess
chlorine. In addition, Facility C reported difficulty in maintaining
efficient automatic ORP control. Subsequent test results indicate that
chlorine dosage rates of less than 2,000 mg/L should be aufficient to oxidize
the cyanide to pernit ted levels while 2500 ng/L was sufficient to oxidize
thiocyanates to below detection limits. Once the chlorination system is
effectively automated, it is anticipated that effluent guidelines for cyanide
will be met.
While most research on alkaline chlorination has focused on
stoichiometry, rate factors, and destruction efficiencies, little work has
been performed on sludge generation and handling characteristics. Researchers
at the University of Tennessee and Illinois have investigated sludge and
supernatant quality following cadmium cyanide destruction and precipitation.
The first objective in the investigation was to examine the alkaline
chlorination of cadmium cyanide solutions in the pH region of carbonate
precipitation. Previous research and field data have shown that carbonate
precipitation results in reduced metal solubilities and improved sludge
21 22
characteristics (see Section 10.1.3). ' As shown previously, an
equivalent level of carbonate is produced from the destruction of the cyanide
radical.
The second objective was to investigate the effects of two forms of
hypoehlorite on cadmium solubility and solid phase characteristics. The two
forms of hypoehlorite investigated were sodium (NaOCl) and calcium
(Ca(OCl)_) hypoehlorite. Previous work hss indicated that sludge produced
from calcium hypoehlorite oxidation dewater more effectively than the more
gelatinous sodium hypoehlorite oxidation sludges. This is primarily due to
the coprecipitation of calcium carbonate and metallic carbonate which due to
calcium granular nature results in distinctly different filtersbility
characteristics,
In the pH region between 7 and 10 for the calcium hypoehlorite system
both cadmium carbonate and calcium were formed as separate crystals. However,
in the optimum range for cyanate oxidation (pH 8.5-9.0) twice as much calcium
carbonate as cadmium carbonate was precipitated (on a molar basis). This
resulted in a dry weight sludge product of only 22-30 percent cadmium. In
U-15
-------
-------
contrast sodium hypochlorite cyanate oxidation in-"the pH-range of 8.5-9.0
resulted in 70 to 100 percent of the precipitate formed consisting of cadmium
carbonate. This represents a dry weight sludge yield of approximately
65 percent cadmium. Therefore while calcium-based hypochlorite systems may
produce a precipitate which filters more readily and to a higher solids
concent, sodium-based hvpochlorite systems theoretically yield a sludge which
is more amenable to metals recovery.
14.1.3 Process Costs
Figure 14.1.3 illustrates the process flow schematic' developed for the
continuous alkaline chlorination system costs contained in this section. The
influent waste water stream is assumed to contain 50 og/L of cyanide ion and
200 mg/L of heavy metal ions. Three flow rates were coated 1,000, 10,000 and
100,000 gallons per hour. These systems were assumed to operate 24 hours per
day, 300 days per year. Complete reaction in the cyanide chlorination tanks
is assumed to occur and the heavy .metals are rendered insoluble in the
precipitation reactor. ' .••.-••.
Cost data and design and operating cost assumptions for the equalization
" "tank/'precipitation""reacror, flocculato'r/clarif"i'er, sludge holding tanks, and
5 13 23 24
filter press have been presented previously in Section 10.1. ' * * The
capital costs for the alkaline chlorination unit has been adapted from
Figure 14,1.4. The unit uses sodium hydroxide for pH adjustment and sodium
hypochlorite as the oxidizing agent. The operations are conduc-ted in two
series-connected reaction tanks in which reagent demand in each stage is
determined by measuring pH and ORP. The reaction time.in each stage is
assumed to be 60 minutes to ensure complete cyanide destruction. The cost for
the system also includes storage and feed systems for the treatment reagents.
Table 14.1-5 contains the capita-1 and operatia-g--costs for the continuous
alkaline chlorination system developed for this section. It is immediately
apparent that at the higher flow rates chemical and sludge disposal costs can
constitute up to 60 percent of the total annual costs. In addition, the
presence of other oxidizable species, stable complexes, or higher influent
cyanide concentrations could render chis process economically.nonviable.
-------
WAST! WATER-
EQUALIZATION
TANK
ALKALINE
CHLORiNATJON
TANKS
PRECIPITATION
TANK
CLAR1FIER
•CAUSTIC AND
SODIUM HYPOCHLORITE SOLUTION
HYDRATED L-ME
OVERFLOW
UNDERFLOW
FILTER
AQUEOUS
PHASE
SF'LUES'T
TREATED
SLUDGE
Figure 14.1.3. Alkaline chlorination. process.
-------
50 r-
10
Legend:
—— Total installed cost
-^ — Hardware cost
Moves:
Balch units;
Installed cost = 2x hardware cost-
Unii consists of tv/o d-hour reaction
tanks with necessary auxiliaries
Continuous units:
Installed cost = 1.25x hardware cost.
FLOW RATE (gal.'min)
Figure 14.1.4. Investment cost for cyanide oxidation units.
-------
TABLE 14,1.5. CONTINUOUS ALKALINE CHLORINATION TREATMENT COSTS3
Purchased Equipment and Installation (PESI]
Equilization Tank
Cyanide Oxidation Units
Precipitation Reactor
Flocculator/Clarif ier
Sludge Holding Tank(s)
Filter Press • -
Total Capital Investment (360% PE&I)
Annual Operating Costs CS/Yr.)
Operating 'Labor C$20/hr.)
Maintenance (6% TCI")'
General Plant Overhead (5.8% TCI)
Utilities (2% TCI)
Taxes and Insurance (1% TCI)
Chemical Costs:
NaOH <$175/ton)
NaOCl (S0.38/gal)
Lime ($40/ton)
Sludge Transportation { $Q.25/ton-mile)
Sludge Disposal ($200/ton)
Annualized Capital (CFR-Q.177)
Total Cost/year
Cost/1000 gallon
1,000
)
17,000
28,000
24,000
18,000
3,000
10,000
100,000
360,000
72,000
21,600 '
20,900
7,200
3,600
1,900
7,800
500
200
12,000
63,700
211,400
' "T9
Flow rate Cgpi
10,000
29,000
61,000
40,000
50,000
6,000
25,000
231,000
831,600
72,000
49,900
48,200
16,600
8,300
19,200
78,200
5,300
2,300
.'-. 120,000 -
147,200
567,200
•:'• ;• g
h)
100,000
50,000
267,000
160,000
203,000
48,000
100,000
828,000
2,980,800
.'
72,000
178,800
172,900
59,600
29,800
191,700
781,700
53,000
22,500
: 1,200,000
527,600
3,289,600
5
a!987 dollars.
-------
14,1.4 Status of Technology • - -- • • . •
Alkaline chlorination systems have generally proven reliable if well
maintained and equipped with'well-designed-OB.P control. The' treatment
technology cannot oxidize stable cyanide complexes such as ferrocyanides and
has difficulty treating nickel cyanides. The most widespread application of
cyanide oxidation through alkaline chlorination is in facilities using
cyanides in electroplating operations.
The evolution of toxic hydrogen cyanide gas may be a problem if pH levels
are lowered excessively. In cases where alkaline chlorination is used to -
treat dissolved complex cyanides and dissolved cyanides of heavy metals,
sludges of metal hydroxides and carbonates are generated. These sludges'can
be recovered by filtration and treated by chemical fixation/solidification.
Many of the chemicals used in this process have potential for hazardous
and or toxic effects if catastrophically released during shipment, storage, or
handling. Liquid sodium hydroxide (greater than <4Q percent) and concentrated
911
sulfuric acid are extremely corrosive. " Chlorine gas and hypochlorite
salts are powerful oxidizers and must be segregated to avoid -reaction with
other chemicals. For a summary of the advantages and disadvantages of -
alkaline chlorination see Table,. 14.1.. h. - • , ; , . - - -
•-14=20;..
"
-------
TABLE 14.1.6. ADVANTAGES AND DISADVANTAGES OF ALKALINE CHLORINAT10N
Advantages
Proven technology with documented cyanide destruction efficiencies.
Operates at standard operation temperatures and pressures and is well
suited to automatic control.
Modular design allows for plant expansion and can be used in
different configurations.
When treating dissolved HCN, calcium, potassium, or sodium cyanide r;a
sludges are generated.
Pisadvantages
.Need for careful pH and ORP control.
Possible chemical interference in the treatment of mixed wastes
(i.e., large oxidation chemical excesses required for complete
, _ _. _ reactions,^). . • •• - - •
Process is not selective and therefore restricted to specific product
wastescreams.
- Potential hazard of shipping, storing, and handling of chlorine gas,
hypochlorite salts, sodium hydroxide, and concentrated sulfuric acid.
Unable to treat ferro and ferricyanides and has difficulty treating
nickel cyanide.
Potential for creating toxic residue which will require
post-treatment (.i.e., fixation/solidificacion/ encapsulation).
Source: Adapted from References 1, 2, 6, ?, and 11.
-------
REFERENCES
1. U.S. EPA. Treatability Manual, Volume III. EPA-6Q0/8-8G-Q42. July 1980.
2. Cushnie, G. C. CENTEC Corporation, Navy Electroplating Pollution Control
Technology Assessment Manual. CR 84.019. February 1984.
3. Sundstrom, B. W., H, E. Kiel, Wastewater Treatment. Prentice-Hall.
Englewood Cliffs, N.J. 1979.
4. Chillingworth, H. A. Alliance Technologies Corporation, Industrial Waste
Management Alternatives and their Associated Technologies/Processes.
GCA-TR-80-80-G. February 1981.
5. U.S. EPA. Reducing Water Pollution Costa in the Electroplating
Induitry. EPA-62S/5-85-016. September 1985.
6. Lanoutte, K. H. Heavy Metals Removal. Chemical Engineering.
October 17, 1977.
7. Butcher, B. 1., R, A. Minear, and R, B. Robinson. University of
Tennessee, Destruction of Cadmium Cyanide Waste by Alkaline Chlorination
Treatment. 17th Mid-Atlantic Industrial Waste Conference. 1984.
8. U.S. EPA. Development Document for Effluent Limitations Guidelines and
Standards for the Metal Finishing Point Source Category (Proposed).
EPA-440/l-82-091-b. August 1982.
9. Mitre Corporation. Manual of Practice for Wastewater Neutralization and
Precipitation. EPA-600/2-81-148. August 1981.
10. Cushnie, G. C. Removal of metals from wastewater: Neutralization and
Precipitation. Pollution Technology Review, No. 107. Noyes Publication,
Park Ridge, NJ. 1984.
11. Kirk-Othmer Encyclopedia of Chemical Technology. Vol. 14, 3rd Edition.
John Wiley and Sons, New York, NY. 1981.
12. Hoyle, D, L. Designing for pH control. Chemical Engineering.
November 8, 1976.
13. Peters, M. S., and K. D. Timnerhaus. Plant Design and Economics for
Chemical Engineers. McGraw-Hill Book Company, New York, NY. 1980.
14, Hoffman, F. How to select a ;pH control system for neutralizing waste
acids. Chemical Engineering. October 30, 1972.
15. Jungels, P. R., and E. T. Eoytowicz. Practical pH Control. Industrial
Water Engineering. February/March 1972.
14-22
-------
-------
16. Martin, J. J. Chemical Treatment of Hating Waste for Removal of Heavy
Hetals. EPA-R2-73-044. May 1973.
17. Sund, S, Physical/Chemical Processing Options. Hazardous Waate and
Hazardous Materials. Vol. 3, No. 2. 1986.
18. Zwikl, J. R.f N, S. Buchko, and D. R. Junkins. Physical/Chemical
Treatment of Coke Plant Wastewaters.
19. Schmidt, J. W. L. Simovic. and E. E. Shannon. Development Studies for
Suitable Technologies for the Removal of Cvanide and Heavy Metals from
Gold Milling Effluents. 36th Industrial Waste Conference, Purdue
University. 1981.
20. Erkku, H., and L. S. Price. Treatment of Gold Milling Effluents. 34th
Industrial Waste Conference, Purdue University. 1979.
21. HSU, D. Y., et «1» Soda Ash Improves Lead Removal in Lime Precipitation
Process. 34th Industrial Waste Conference, Purdue University. 1978.
22. Mabbett, Cappacio and Associates. Industrial Wastewater Pretreatment
Study: Preliminary Engineering Design Report. January 1982.
23. Versar Inc. Technical Assessment of Treatment Alternatives for Hastes
Containing Metals and/or Cyanides. Contract No. 68-03-3149.
U.S. EPA/OSW. October 1984.
24. U.S. EPA. Economics of Waetewater Treatment Alternatives for the
Electroplating Industry: Environmental Pollution Control Alternatives.
EPA 625/5-79-016. 1979.
14-23
-------
-------
14.2 OZONATION -
Chemical oxidation has the potential for removing from wastewaters
organic materials which are resistant to other treatment methods,
e.g., refractory materials which are toxic to biological systems. Ozone
(0,,} is one of the strongest oxidants available, as shown in Table 14.2,1,
which lists the oxidation potential and relative oxidation power of a number
of oxidizing agents. Ozone, as an oxidant, is sufficiently strong to break
many carbon-carbon bonds and even to cleave aromatic ring systems.
Ozone has been used for years in Europe to purify, deodorize, and
disinfect drinking water. More recently, it has been used in the waste
treatment area to oxidize cyanide wastewaters. Cost and mass transfer
considerations restrict usage of ozone to the treatment of wastewaters with -
23,
1 percent or lower contaminant concentration levels. * Since oxidation by
ozone occurs nonselectively, it is also generally used only for aqueous wastes
which contain a high proportion of hazardous constituents versus nonhazardous
oxidizable compounds, thus focusing ozone usage on contaminants of concern.
Ozonation may be particularly useful as a final treatment for waste streams
which are dilute in oxidizable contaminants, but which do not 'quite meet
standards.,
14.2.1 Process Description
Ozone is generated on site by the use of corona discharge technology.
Electrons within the corona discharge spilt the oxygen-oxygen double bonds
upon impact with oxygen molecules. The two oxygen atoms formed from the
molecule react with other oxygen molecules to form the gas ozone, at
equilibrium concentration levels of roughly 2 percent in air and 3 percent in
oxygen (maximum values of 4 and 8 percent, respectively). Ozone must be
produced onsite (.ozone decomposes in a matter of hours to simple, molecular
oxygen } and ozonation is restricted co treatment of streams with low
quantities of oxidizable materials. Using a rule of thumb, two parts of ozone
are required per part of contaminant. A large commercial ozone generator
producing 500 Ib/day of ozone could treat 1 million gallon/day of wastewater
containing 30 ppm of oxidieable matter, or equivalently, 3., 000 gallons/day of
-------
TABLE 14.2.1, RELATIVE OXIDATION POWER OF OXIDIZING'SPECIES
Species
Flourine
Hydroxyl radical
Atomic oxygen
Ozone
Hydrogen peroxide
Perhydroxyl radicals
Permanganate
Hypochlorous acid
Ch lorine
Oxidation
potetitail ,
volts
3.06
2,80
2.42
2.07
1.77 '
1.70
1.70
1.49
1.36
Relative
oxidation
power* ' '
2.25
2.05 '
1.78
1.52
1.30
1.25 .'
, ,.1.25--
1.10
1.00
aBased on chlorine as reference (= 1.00).
Source: References 1 and 2.
-------
wastewater containing i percent of oxidizable matter. Extensive
information related to the generation of ozone and its application to the
treatment of industrial wastewaters can be found in References 5 through 9.
While direct ozonation of industrial uastewater is possible and is
practiced commercially, other technologies- have been contained 'with ozonation
to enhance the efficiency and rate of the oxidation reactions. These
technologies, which supply additional energy to the reactants, involve the use
of ultraviolet light or ultrasonics.
Cyanides are decomposed by ozone according to the general rate expression:
"d [°3] - fc [CN ,0-63 ± 0.04
dt C (1)
where [0,J and [CN ] are the concentrations of ozone and the total cyanide
(including thiocyanates) and k is the reaction rate constant. Pilot plant and
bench-scale data indicates that the reaction is first order with respect to
ozone, and fractional order with respect to , the cyanide ion. This
fractional order of the cyanide ion indicates the ozone-cyanide reaction is
not a simple, bimolecular reaction but involves the formation and reaction of
free radicals. Reactions of OH and HO with 0, can initiate the
*"" " 10 . -
radical chain reactions. Therefore, pH considerations, as' indicated by
the following rate relation for the decomposition of ozone, are important in
determining the overall rate equations:
'V
dt (2)
However, it should be noted that the limiting factor in ozone rate
equations is the mass transfer of ozone gas to the liquid phase. Pilot plant
data will be required to determine uass transfer characteristics. Research at
Drexel University has focused on these rate relations in an efforts to
generate fundamental kinetic and mechanistic data for the reactions of ozone
with cyanide by distinguishing between mass transfer of ozone and the
11 12
oxidation and decomposition reactions of ozone. ' Figure 14.2.1
illustrates the profiles obtained for total cyanide, cyanate, and ozone
residuals using an ozone bubble column and pHs of 11.2, 7.0, and 2.5, The
results show that reaction rate increases with increasing pH and demonstrates
a varying dependence on cyanide concentration at different pH values.
-------
pH=ll.2
•i
4.0 -
pH=7.0
pH=2.5
OZONE ~
. o —o —J —
TOTAL
CYANIDE
0 5 10 13 20 25 30 0
5 10 15 20 25 30
TIME , mlnut«&
CrANATE
15
10
0 5 10 15 20 25 30
Figure 14.2.1. Profiles of total cyanide, cyanate, and ozone residual in the bubble column
Eor pH 11.2, 7.0, and 2.5.
-------
Upon oxidation of each mole of cyanide, .1.2 + 0.2 Bol of ozone is
consumed and I vole of cyanate is produced .as the reaction product. At pH
!!•£, the removal rate of cyanide ia mass transfer limited because of its very
high oxidation rate with ozone, as indicated by the zero-order behavior of the
cyanide profile. Cyanate appears in the solution at a rate which is equal to
the rate of removal of cyanide. After cyanide is oxidized completely, cyanate
starts to react with ozone at a much slower rate. During the course of the
experiment at this pH» ozone does not appear in solution because of its rapid
consumption by the oxidation and the decomposition reactions. It is postulated
that if the ozone and cyanate were allowed to react further, the cyanate would
be completely decomposed into harmless constituents.
The removal rate of cyanide at pH 7.0 is equal to the rate at.pH 11.2 and
is mass transfer limited for the first 12 min. of ozonation. However, after
the total cyanide concentration is reduced to about 0,8 mM, the oxidation
reaction becomes the rate-limiting step. Ozone appears in solution as soon as
the system becomes reaction rate limited and accumulates until reaching a
plateau at about 14 mg/L. Cyanate is produced at an equal to cyanide rate
oxidation; however, oxidation of cyanate starte while cyanide still exists in
solution.
At pH 2.5, volatilization of HCN contributes more to the removal of
cyanide than its oxidation by ozone, as demonstrated by independent
experimentations with pure oxygen. Nevertheless, oxidation of cyanide
produces equal moles of cyanate. Due to slow oxidation and decomposition of
ozone at this pH, ozone appears in solution instantaneously and stabilizes at
11 12
a saturation value of about 14 mg/L. '
To effectively bring about the reaction of ozone with reactive
contaminants, it is important that mass transfer of ozone and its reactants
through the gas-liquid interface be maximized. Also, to increase ozone
solubility in water, temperatures should be maintained as low as possible and
pressures as high as possible. Under conditions leading to maximum reactivity
rates, costs may also increase due to less efficient use of ozone. Decisions
will have to be made on a case-by-case basis to establish the most effective
operating conditions.
Several commercial designs are available for the conduct of gas/liquid
reactions which bring reactanti into contact as effectively as possible (see
Table 14,2.2 for a list of some commercial equipment vendors). The types
14-28
-------
-------
TABLE 14.2.2. MAJOR U.S. MANUFACTURERS'OF OZONE GENERATING EQUIPMENT
1 I
N>
vO
Manufacturer
Crane Cochrane
Emery Industries, Inc. ,
Ozone Technology Group
Ozone Research &
Equipment Corporation
PCI Ozone Corpora t ion
We I a bach Ozone Systems
Corporation
Infilco Degremont,
Union Carbide
Linde Division
EnvironmentaL Systems
U.S. Oconair Corp.
Address
P.O. Box 191
King of Prussia, PA 19406
(2L5) 265-5050
4900 Eutce Avenue
Cincinnati, OH 45232
(513) 482-2100
3840 North 40th Avenue
Phoenix, AZ 85019
(602) 272-2681
West Ca Id veil, NJ 07006
(201) 575-7052
3340 Stokely Street
Philadelphia, PA 19129
(2L5) 226-6900
Bo« K-7
Richmond, VA 23288
(804) 285-9961
P.O. Box 44
Tonauanda, NY 14150
(716) 877-1600
464 Cabot Road
S. San Francisco, CA 94080
(415) 952-1420
Equipment
Concentric tubea
SS/glaBB /aluminum
Series C - cabinet
Series P - akid
mounted
Concentric tubea
SS/glaBS/nichrome
Skid mounted
Concentric tubea
SS/glass/SS
Series V, B & D
cabinet
Seriea H skid
mounted
Concentric tubes
SS/glasB/fiilver
Series G - cabinet
Series B - skid
mounted
Concentric tubea
SS/glaoo/SS
Series CLP & GLP
Both skid mounted
SS/g la eo /aluminum
Skid mounted
Parallel ceramic
coated steel
Lowther plates
Concentric tubea
Titanium/ceramic/
aluminum
Hodels-
capacit iee
Ib 03/doy-
air feed
Series C,
1-18 Ib/day
Series P,
18-122 Ib/day
Series 9270,
1-23 Ib/day
Series 9260,
21-400 Ib/day
Series B & V,
1/4-2 Ib/day
Series D & H
4-250 Ib/day
1-28 Ib/day
Series B,
35-1400 Ib/day
Series CLP,
24-127 Ib/day
Series GLP,
170-322 Ib/day
designations ,
10-600 Ib/day
No model
des ignat ions ,
1-L200 Ib/day
Series HF.
5-570 Ib/day
Typical Oj
Cool ing tion in air,
method percent
Water on 1
outer
electrode
Hater on 1
outer
electrode
Water on 1
outer
electrode
inner
electrode
Oil on
miter
electrode
Water on 1
outer
electrode
outer
electrode
Air on 1
outside both
electrodes
Water on 2
inner
electrode
Air on outer
electrode
Source: Reference 13.
-------
-------
of reactor designs available range from mechanically agitated reactors to more
complex spray, packed, and tray type towers. Their advantages and limitations
are discussed in detail in many standard texts and publications (for example,
see References 2 through 5).
The process of UV/ozone treatment operates in the following manner. The
influent to the system is mixed with ozone and then enters a reaction chamber
where it flows past numerous ultraviolet lamps as it travels through the
chamber (see Figure 14.2.2). Flow patterns and configurations in the UV
exposure chamber are designed to maximize exposure of the total volume of
ozone-bearing vastewater to the high energy UV radiation. Although the nature
of the effect appears to be influenced by the characteristics of the waste,
the UV radiation enhances oxidation by direct dissociation of the contaminant
molecule or through excitation of the various species within the waste
stream. In industrial systems, the system is generally equipped with recycle
capacity. Gases from the reactor are passed through a thermo catalytic unit,
destroying any volatiles, replenished with ozone, and then recycled back into
the reactor. The system has no gas emissions.
Another alternative process involves the coupling of ultrasonic energy
with ozonation. It has been shown that significant increases in the rate of
oxidation can be obtained by the use of ultrasonic energy as apposed to ozone
alone. Experimental details were not available in Reference 3, although
different oxidation pathways were reported operating in the presence or
absence of ultrasonics. Regardless of the reaction mechanisms, there appears
to be no doubt that the combination of ozonation with either UV or ultrasonic
excitation leads to increased oxidation rates. Typical design data for one
14
40,000 gal/day UV/ozone treatment process are shown in Table 14.2.3.
In addition to reactor design, contactor system optimization and UV
radiation utilization, two key factors in ozone equipment selection and design
13
are power consumption and ozone generator cooling. Typicallyj the major
operating cost for ozone manufacturing is the cost of electric power. Power
consumption figures in facilities using air as the source of ozone range from
6 to 8 kWh/lb 03 for the ozone generator alone, and 10 to 13 kWh/lb 0^
total consumption including air handling and preparation. Using oxygen as
feed gas reduces these ranges to 3 to 4 kWh/lb 0~ for ozone generation and
7 to 12 kWh/Lb CU total consumption (depending on the source of
13
oxygen). However, when pure oxygen is used as the ozone manufacturing
reagent, chemical costs will also have to be included.
14-30
-------
-------
UV
S~~ -v'
o
o
o
o
o
d>
o
o
o
o
N
\ ^~^.
P
0
o
o
o
o
o
o
o
o
X
/ — -v.
o
o
0
o
o
_^
o
o
o
o
o
Flow distributor
W«s:e witef in
Figure 14.2,2. Schematic of top view of ULTROX pilot plant by
General Electric (ozone sparging system omitted)
(Edwards, B. H.> 1983).
Source; Reference 4. -
14-31
-------
-------
TABLE 14.2.3, DESIGN DATA FOR A 40,000 GPD (151,400 L/DAY) ULTROX PLANT
Reactor
Dimensions Meters: (LxttxH) 2.5 it 4.9 x 1.5
Wet volume, liters 14,951
UV lamps:
Number of 65 watt lampe 378
Total power, RW 25
Ozone generator
Dimensions Meters: (LxWxH) 1.7 x 1.8 x 1.2
gms ozone/minute 5.3
kg ozone/day 7.7
Total power, kW 7.0
Total energy required (RWH/day) 768
Source: Reference 14.
14-32
-------
-------
Ozone generator cooling costs arise since ozone generators must be
continuously cooled to maintain optimum efficiency and to avoid deterioration
of the dielectric. Generators are usually air- or water-cooled. All
manufacturers but one use water to cool their medium and large size
generators. Typical generator cooling fluid requirements are 100,000 ft
air/lb 0. or 500 gal water/lb 0- for systems using air as the ozone source.
Due to the nonselective nature of the ozonation reactions it is important
that the concentration levels of nonhazardous, but oxidizable, contaminants in
the feed stream be reduced as much as possible prior to treatment. The strong
electrophilie nature of ozone imparts to it the ability to react with a wide
variety of organic functional groups, including aliphatic and aromatic
carbon-carbon double and triple bonds, alcohols, organometallic functional
groups, and some carbon-chlorine bonds. It is important to recognize that
many functional groups can be present which compete for the ozone reactant and
can add significantly to the cost of the treatment.
The waste to be treated should also be relatively free of suspended
solids, since a high concentration of suspended solids can foul the equipment
normally used to bring about contact between ozone and the aqueous phase
contaminants,.^....IBhen ,t?zonation. is,.combined .with UY_ radiation- or ultrasonics, -a.--.-
high concentration of suspended solids also can impede the passage of UV
radiation or attenuate the energy supplied by,ultrasonics to enhance the
oxidation, rate. Other pretr'eatments include flow equalization, neutralization,
and oil and grease removal.
Post-treatment of industrial wastewaters that have been contacted with
ozone will involve elimination of residual ozone, usually by passing the
effluent through a thermocatalytic unit. Some by-product residuals may be
formed in the feed water and some contaminants, if present, will not undergo
reaction. Compounds considered unreactive include many chlorinated aliphatic
compounds. If these compounds are present in the waste, technologies other
than ozonation should be considered.
14.2.2 P r oc e s s Pe r f orm an c e
Although there has been a great deal of research into the ozonacion of
cyanide in the last 30 years, only a few commercial plants have been
installed. Data are limited and additional studies are needed to establish
the utility of ozone treatment.
14-33"
-------
-------
Two applications of ozone for the oxidation;of cyanides that have been
reported in the literature are those at San.Diego Plating, San Diego,
California, and Sealector Corporation, Mamaroneck, New York. The San Diego
Plating System was installed by Ozodyne Corporation to treat waste-waters from
an automobile recycling operation. The Sealector system was installed by PCI
Ozone Corporation under funding from EPA's R&D branch in Cincinnati.I5,Jo
San Diego Plating's ozone system consists of 300 gallon reactor, a vacuum
precoat filter, and a solids collection unit. Prior to the ozone reactor,
ozone gas under negative pressure is drawn into the waste stream to be
treated. The wastewater containing dissolved ozone and ozone gas is then
formed into fine particles to enhance mass transfer using a spinning dial
type aspirator. The treated wastewater is then pumped to the filter where the
solids are dewatered while the filtered effluent is discharged to the sewer.
Table 14.2.4 summarises sampling results for the ozone system at San Diego
Plating. Cyanide effluent concentrations were consistently reduced to very
low levels. In addition, oxidation of the metal hydroxide solids reduced the
degree of hydration and improved the dewatering characteristics (74 percent
solids versus 20-30-percent for conventional precipitation solids).
The treatment system at Sealector was similar'except that the ozone
reaction tank consisted of two separate compartments. One tank was used to
treat the wastewater with ozone while the second tank recovered unreacted
ozone from the off-gas and recycled it back to the "incoming cyanide waste
stream. However, constant equipment failure, operation problems, and
process unreliability resulted in the ozone system being replaced with more.
conventional alkaline chlorination technology.-^ San Diego Plating has also
taken its ozone system out of cyanide oxidation operations and replaced it
with a batch chlorine unit.18 The main drawback of the commercial systems
discussed above was high capital investment and operational costs.
Figure 14.2.3 which compare conventional waste treatment costs with those o£
ozone oxidation at San Diego Plating. Figure-14.2.3 indicates that ozone
oxidation appears to be more expensive than conventional treatment systems
over the range shown.16
-------
TABLE 14.2.4. SUMMARY OF SAMPLING RESULTS - SAN DIEGO PLATING
Parameter
Cyanide
Total chrome
Copper
Nickel
TSS
pH
Range
3.75 -
6.62 -
33.0 -
60.0 -
559
12.2 -
influent
0.05
0.82
5.05
10.2
35
3.4
Average
1.02
1.41
9.45
20.32
135
6.4a
. Effluent
Range
0.87 - <0.02
1.55 - 0.05
1.32 - 0.04
0.37 - <0.10
93 - <1
12.4 - 5,8
Ave
0
0
0
0
11
8
rage
.08
.40
.05
.13
.6
.4a
Average
and
removal
>92.5
>71.6
99.5
>99.4
>91.5
—
aMedian.
Average solids content of sludge = 74 percent.
Influent and effluent values, except pH, in mg/L.
Source: Reference 16.'
-------
K
>i
"V
J
o
Q
CL,
O
w
Q
to
O
H
O
, tJ
o
u
t,
o
Q
w
Z
s
80
60
20
Note: Conventional treatment-
includes chrome reduction,
neutralization,, cyanide
destruction, precipitation,
clarification and settling.
(1984 costs)
25 50 75 100
WASTEWATER FLOW RATE (GPK)
Figure 14.2.3. Annualized operating costs for conventional
treatment. Ozodyne treatment system at
San Diego Plating.
Source: Reference 16,
-------
14.2,3 Process
Table 14.2.5 lists the costs for a 40,000 gpd UV /Ozone plant for which
design data were shown in Table 14.2,3. Cost estimates were based on
wastewater containing 50 ppra PCB, designed to achieve an effluent
concentration of 1 ppm. Costs were considered to be competitive with
14 19
activated carbon. ' The unit cost for treatment of the waste is greatly
affected by whether or not the cost for a monitoring system is included. The
cost of PCB destroyed is in excess of $10/paund, PCB data were used for
costing purposes because of its availability. However, the costs -will increase
substantially if ozonation is to be used as treatment for a . waste containing
1 percent organic contaminants. This is 200 times the concentration used to
develop the costs in Table 14.2.5. Assuming capital equipment costs follow a
20
simple "sixth-tenths" factor scaling relationship, the costs of the
reactor and generator would be about $3,000,000 (or 24 times the costs shown
in. Table 14.2.5) for treatment of this higher concentration. Scale factors
would be variable for the operating and maintenance cost items listed in
Table 14.2.5. However, the net result of scale-up to handle the more
concentrated waste would drastically increase the cost/1,000 gallons, treated,
but would also result in far lower costs when calculated on the basis of the
amount of contaminant destroyed. Costs of roughly $10/pound of contaminant
destroyed would be reduced to an estimated $l/pound, assuming comparable
efficiencies. However, destruction efficiencies may be adversely affected at
higher concentrations due to mass-transfer and other. considerations. Thus,
the cost benefits per pound of contaminant destroyed, as stated above, may not
be fully achievable. An optimal tradeoff must be made on the basis of
pilot-scale or full scale test results.
14.2.4 Overall Status of Process
Availability —
Ozonation equipment is available commercially from several manufacturers'
within the United States. The Chemical Engineering Equipment Buyers' Guide
published by McGraw Hill iiscs nine manufacturers of ozone generators and
10 manufacturers of ozonators. The latter classification includes firms that
usually provide the ozone generator, the reactor, and auxiliaries such as the
-------
TABLE 14.2.5. EQUIPMENT PLUS OPERATING AND MAINTENANCE
COSTS: 40,000 GPD UV/OZONE PLANT
Reactor $ 94,500
Generator 30,000
- £124,500
0 & M costs/day
Ozone generator power $ 4.25
UV lamp power 15.00
Maintenance 27,00
(Lamp replacement)
Equipment amortization
(LO years at 10 percent) 41.90
Monitoring labor 85.71
Total/day: $ 173.86
Cost per 1,000'gallons (3,785 liters)
with monitoring labor . $ 4.35
Cost per 1,000 gallons without
monitoring labor $ 2.20
Source: Reference 14.
M--3&
-------
catalytic unit for destruction of ozone from the treated.stream. The status
of UV/ozonation is far less advanced. Processes such as the llltrox
4 14 19 • '
process ' * have been concerned with highly refractory compounds such aa
PCBs, Equipment specifically designed and available for UV/ozonation of
industrial wastewaters, is not available as a standard commercial item.
Application-—
Ozonation appears best suited for treatment of very dilute waste streams,
similar to those streams treated by the ozone based water disinfection
processes now used in Europe. It does not appear to be cost competitive or
.technically viable for most industrial waste streams where organic
concentration levels are 1 percent or higher. However, it may be viable for
certain specific wastes with high levels of a contaminant of special concern
and high reactivity,
Environmental Impact—
Assuming adequate destruction of a contaminant by ozonation, the
principal environmental impact would appear to be associated with ozone in the
effluent vapor and liquid streams*. However, thermal decomposition of ozone is
effective and is used commercially to destroy ozone prior to discharge,
Unreacted contaminants or partially oxidized residuals'in the aqueous effluent
may be a problem necessitating further treatment by other technologies,
Presence of many such residuals will generally result in selection of a more
suitable alternative technology.
Advantages and Limitations—
There are several factors which suggest that ozonation nay be a viable
1 4
technology for treating certain dilute aqueous waste streams: *
Capital and operating costs are not excessive when compared to
incineration provided oxidizable contaminant concentration levels
are less than 1 percent.
The system is readily adaptable to the onsite treatment of hazardous
waste because the ozone can and must be generated onsite.
-------
* Ozonation can be used as a final treatment for certain wastes since
effluent discharge standards can be met.
• It can be used as a preliminary treatment for certain wastes (e.g.,
preceding biological treatment).
However, there are limitations which often will .preclude use of ozonation
as a treatment technology. These include:
Ozone is a nonseleetive oxidant; the waste stream should contain
primarily the contaminants of interest,
Certain compounds because of their structure are not amenable to
ozonation, e.g., chlorinated aliphatics.
Ozone systems are generally restricted to 1 percent or lower levels
of toxic compounds. The system is not amenable to bulky wastes.
Toxic intermediates nsay persist in the waste stream effluent.
Ozone decomposes rapidly with increasing temperature, therefore,
excess heat, muse be removed rapidly.
Ozone oxidation is currently not as cost effective or reliable as
alkaline chlorination.
-------
REFERENCES
1. Prengle, H. W., Jr. Evolution of the Ozone/UV Process for Wastewater
Treatment. Paper presented at Seminar on Wastewater Treatment and
Disinfection with Ozone, Cincinnati, Ohio. September 15, 1977.
International Ozone Association, Vienna, VA.
2. Harris, J. C. Ozonation. In: Unit Operations for Treatment of
Hazardous Industrial Wastes. Noyes Data Corporation, Park Ridge, NJ.
1978.
3. Rice, R. G. Ozone for the Treatment of Hazardous Materials, In:
Water-1980; AICHE Symposium Series 209, Vol. 77. 1981.
4. Edwards, B. H., J. N. Paullin, and K. 'Coghlan-Jordsn. Emerging
Technologies for the Destruction of Hazardous Waste - Ultraviolet/Ozone
Destruction. In: Land Disposal; Hazardous Waste. U.S. EPA
600/9-81-025. March 1981.
5. Ebon Research Systems, Washington, D.C, In: Emerging Technologies for
the Control of Hazardous Waste. U.S. EPA 600/2-82-011. 1982.
6. Rice, R. G., and M. E, Browning. Ozone for Industrial Water and
Wastewater Treatment, an Annotated bibliography. EPA-600/2-80-142,
U.S. EPA RSNERL, Ada, OK. May 1980.
7. Rice, R. Gi, and M, E. Browning..- Ozone for. Industrial Water and
Wastewater Treatment, A Literature Survey. EPA-600/2-80-060. U.S. EPA
RSKERL, Ada, OK. April 1980.
8. International Ozone Institute, Inc., Vienna, VA. First International
Symposium on Ozone for Water and Wastewater Treatment. 1975.
9. International Ozone Institute, Inc., Vienna, VA. Second International
Symposium on Ozone Technology. 1976.
10. Gurol, M. D., W. M. Bremen, and T. E. Holden. Drexel University
Oxidation of Cyanides in Industrial Wastewaters by Ozone. Environmental
Progress. February 1985.
11. Gurol, M. D., and W. K. B-remen. Drexel -University Kinetics and Mechanism
of Ozonation of Free Cyanide Species in Water. .Environmental Science and
Technology. Vol. 19, No. 9. 1985.
12. Gurol, M. D., and P. C. Singer. Drexel University Kinetics-of Ozone
Decomposition: A Dynamic Approach. Environmental Science and
Technology, Vol. 16,. No. 7. 1982.
13. Derrick, D. G., and S. R. Perrich. Guide to Ozone Equipment Selection.
Pollution Engineering. November 1979.
-------
14. Arisman, R. K,, and Rt C» Mueick. Experience in Operations.of a UV-Ozone
Ultrox Pilot Plant for Destroying PCBs in Industrial Waste Effluent.
Paper presented at the 35th Annual Purdue Industrial Waste Conference.
May 1980.
15. Militello, P. Centec Corporation, Assessment of Emerging Technologies
for Metal Finishing Pollution Control: Three Case studies.
EPA-600/2-81-153. August 1981.
16. Cushnie, G. C. Centec Corporation, Navy Electroplating Pollution Control
Technology Assessment Manual. Cfi 84.019. February 1984.
17« Sacco, S. Sealector Corporation. Telephone conversation with Stephen
Palmer. February 26, 1987.
18. Needham, B. San Diego Plating. Telephone conversation with Stephen
Palmer. February 26, 1987. •
19. Zeff, J. D. Westgate Research Corporation. Ultrox Process Treatment o£
Organic Wastewater. Third Annual Conference on Treatment and Disposal of
Industrial Wastewaters and Residues. 1978.
20. Peters, M, S., and K. D. Timmerhaus. Plant Design and Economics for
Chemical Engineers. McGraw-Hill, New York, NY. 1980.
-------
14.3 WET AIR OXIDATION ' -
Wet air oxidation tWAO) is the oxidation of dissolved or suspended
contaminants in aqueous waste streams at elevated temperatures and pressures.
It is generally considered applicable for the treatment of certain
organic-containing media that are too toxic to treat biologically and yet too
1 2
dilute to incinerate economically. * A leading manufacturer of commercial
available WAO equipment reports that WAD takes place at temperatures of 175 to
320°C (347 to 608°F) and pressures of 2,169 to 20,708 kPa (300 to 3,000 psig}.1
Although the process is operated at subcritical conditions (i.ei, below 374°C
and 218 atmospheres), the high temperatures and the high solubility of oxygen
in the aqueous phase greatly enhances the reaction rates over those
experienced at lower temperatures and pressures.' In practice, the three
variables of pressure, temperature and time are controlled to achieve' the
desired reductions in contaminant levels.
In addition to serving as the source of oxygen for the process, the
aqueous phase also moderates the reaction rates by providing a medium for heat
transfer and heat dissipation through vaporization. The reactions are
exothermic and proceed-without the need for auxilltary fuel at feed chemical
oxygen demand (COD) concentrations of 20 to 30 grams per liter.
14.3.1- Process Description
A schematic of a continuous WAO system is shown in Figure 14.3.1. The
Zimmerman WAO System, as shown in the figure, has been developed by Zimpro,
Inc. Rothschild, Wisconsin. It represents an established technology for the
treatment of municipal sludges and certain industrial wastes. While.
industrial applications of cyanide destruction through wet air oxidation have
been few, a wet-air oxidation unit developed by Zitapro Corporation was placed
3 A—S3
into operation-in 1983 for commercial off site treatment. * During test
runs the unit e-ffectively treated cyanide wastes; a destruction efficiency of
99.7 percent of the influent cyanide was achieved with cyanide concentrations
of 25,000 mg/L reduced to 82 mg/L,
In the KAO process shown in Figure 14.3.1, the'waste stream containing
oxidizable contaminants is pumped to the reactor using a positive
displacement, high pressure pump. The feed stream is preheated by heat
-------
OXIDIZABLE
WASTE
FEED
PUMP
PROCESS
HEAT
EXCHANGER
REACTOR
AIR COMPRESSOR
PCV
WASTEWATER
Figure 14.3.1. Wet air oxidation general flow diagram.
Source: Reference 4.
-------
exchange with the hot, treated effluent stream. Steam is added as required to
increase the temperature within the reactor to a level necessary to support
the oxidation reactions in the unit. As oxidation proceeds, heat of
combustion is liberated. At feed COD concentrations of roughly 2 percent the
heat of combustion will generally be sufficient to bring about a temperature
rise and some vaporization of volatile components. Depending upon the
temperature of the effluent following heat exchange with the feed stream,
energy recovery may be possible or final cooling may be required. Following
energy removal, the oxidized- effluent, consisting mainly of water, carbon
dioxide, and nitrogen, is reduced in pressure through, a specially designed
automatic control valve. The effluent liquor is either suitable for final ,
discharge (contaminant reduction achieves treatment standards) or is now
readily'biodegradable and can be piped to a biotreatment unit for further
reduction of contaminant levels. Similarly, noncondensibie Rases can either •
be released to the atmosphere or passed through a secondary control device
(e.R., carbon adsorption unit) if additional treatment is required to reduce
Q
air contaminant emissions to acceptable levels.•
The continuous reactor can reportedly take two forns:• & to»er reactor
or a reactor consisting of a cascade o.f completely stirred tank reactors
(CSTRs). The bubble tower reactor available commercially from Zimpro is a
vertical reactor in which air is passed through the feed. The reactor is
sized, based on feed rate, to provide the holding time recuired for the
reactions to proceed to design levels. The stirred tank cascade reactor
consists of a series of horizontal reactor chambers contained within a
horizontal cylinder. The wastewater cascades from one chamber to the next,
and then is released for"discharge or post-treatment. Air is generally
injected into each of the CSTRs.
Although operation of a WAO system is possible, by definition, under all
subcricical conditions; i.e., below 374°C and 218 atm (.3220 psig), commercially
available equipment is designed to operate at temperatures ranging from 175 to
320°C and at pressures of 300 to 3,000 psig.
Of all variables affecting WAO, temperature has the greetest effect on
reaction rates. In most cases, about 150°C (300°F) is the lower limit for
appreciable reaction. About 250BC (4S2eF) is needed for 80 percent reduction
-------
of COD, and at least 300°C (5?2°F) is needed for 95 percent reduction of COD
within practical reaction times. Destruction rates for specific constituents
may be greater or less than that shown for COD reductions.
Initial reaction rates and rates during the first 30 minutes are
relatively fast. After about 60 minutes, rates become so Blow that generally
2
little increase in percent oxidation is gamed in extended reaction times,
An increase in reaction temperature will lead to increased oxidation but
generally will require an increase in system pressure to maintain, the liquid
phase and promote wet oxidation. A drawback to increasing the temperature and
pressure of the reaction is the greater stress placed on the equipment and its
components, e.g., the increased potential for corrosion problems,
In addition, increased temperatures and pressures increase both capital
and operating costs as well as greatly decreasing liquid phase equilibrium
oxygen concentration. Decreased equilibrium oxygen concentrations decrease
gas mass transfer rates which thereby restricts overall reaction rates.
As noted by Zimmerman, et al., the object of WAO is to intimately mix the
right portion of air with the feed, so that under the required pressure,
combustion will occur at a speed and temperature which will effectively reduce
the organic waste to desired levels. Pressures should be maintained at a
level that will provide an oxygen rich liquid phase so that oxidation is -
maintained. Charts and curves are provided in this reference to aid in
the determination of waste heating valve, stoichioraetric oxygen requirement,
and the distribution of water between the liquid and vapor phases at given
temperatures and pressures. More information can be obtained•from the
manufacturer. ^ ' .
A model has been developed to gain insight into the key system parameters
using a common industrial waste stream and fixed temperature, residence time,
and COD reduction. The model was used to estimate costs for the' system.
Its value, as a predictive tool, along with that of supplementary -kinetic
12
studies of batch wet oxidation, is limited by the sparsity of experimental
data concerning reaction products and their phase distributions Et the
elevated temperatures and pressures encountered during WAO.
Very little discussion is found in the literature concerning the physical
form of wastes treatable by WAO, However, WAO equipment and designs have been
used successfully to treat a number of municipal and industrial sludges.
According to a representative of the leading manufacturer of WAO systems,
wastes containing up to 15 percent COD (roughly equivalent to 7 to 8 percent
organics) are now being treated successfully in commercial equipment,
1-4=46 ••
-------
Treatment of solid bearing wastes is dependent upon selection of suitable
pump designs and control devices. HAD units used for activated carbon
14
regeneration now operate at the 5 to 6 percent solids range. Treatment of
higher solid levels is not precluded by fundamental process or design
limitations. Column design must also be consistent with the need to avoid
settling within the column under operating flow conditions. Thus, pretrestment
to remove high density solids (e.g., metals by precipitation) and accomplish
size reduction (e.g., filtration, gravity setting) would be required for some
slurries. It should be noted that the WAO unit operated by the California
facility does not accept slurries or sludges for treatment. This may be a
result of design factors precluding their introduction into the system.
Under typical WAO operating conditions it is likely that both contaminant
residuals and low molecular weight process by-product residuals may be present.
While it is entirely possible that imposition of more stringent operating con-
ditions will serve to reduce these residuals to acceptable levels, the manufac-
turers and users of commercial WAO system stress that the major applications
involve the pretreatment of waste, usually for subsequent biological treatment.
Even under conditions that are favorable for wet oxidation,,it is also
likely that certain contaminants, particularly'some of the more volatile
components, will partition"between the vapor phase and the liquid phase.
Empirical testing will be necessary to establish vapor and liquid phase
residuals and some post-treatment of both streams may be necessary. Existing
post-treatment methods for the liquid generally involves bacteriological
treatment. Although the results of post-treatment schemes for vapors from the
WAO system have not been found in the literature, a two-stage water
scrubber/activated carbon adsorption system has been used to treat WAO vapor
emissions. Presumably carbon as
routinely employed if necessary.
emissions. Presumably carbon adsorption or scrubbing systems could be
14.3.2 Process Performance
Tables 14.3.1 and 14.3.2 summarize cyanide wet air oxidation
demonstration cest results conducted in 1983 at a commercial waste treatment
facility in California. The WAO unit is currently still in operation at
this site. Table 14.3.3 contains cyanide oxidation data for the WAO treatment
of spent caustic scrubbing liquor from a natural gas based ethylene plant.
-------
TABLE 14.3.1. RESULTS OF WET AIR OXIDATION UNIT - OXIDATION OF CYANIDE WASTE
Sample
COD, g/L
COD Reduction, 2
Cyanide, mg/L
Cyanide Reduction, %
PH
Zinc, mg/L
Nickel, ng/L
Copper, mg/L.
Log Book No.
Influent
32.2
-
28,630
_
12.9
15,700
120
1,900
,2082-66-1
Effluent
9.3
71.1
0,82
99.99
9.1
3,500
15
536
2082-66-2
Source: Reference 16.
-------
TABLE 14.3.2. WET AIR OXIDATION DEMONSTRATION OF CYANIDE WASTSWATER
Oxidation Conditions
Oxidation Temperature
*Nominal Residence Time
Waste Flowrate
Reactor Pressure
495°F (257°C)
80 minutes
7.5 gpm
1200 psig
^Nominal Residence Time = Reactor Volume Divided by Waste Flowrate
Oxidation Results
COD, g/L
COD Reduction, %
Total Cyanide, rag/L
Total Cyanide Reduction, I
Off-Gas Grab Sample Analysis
Carbon Dioxide
Oxygen
Nitrogen
Carbon Monoxide
Methane
Total Hydrocarbons
Raw Influent
37.4
25,390
Oxidized Effluent
4.2
88.8
82
99.7
1.5%
8.5%
82.82
Not detected
9.0 ppm
61.1 ppra as methane
.Source: Reference 16.
-------
TABLE 14,3.3. TREATMENT OP SPENT CAUSTIC SCRUBBING LIQUOR FROM NATURAL GAS
BASED ETHYLENE PLANT, 608°F, 3000 PSIG (320"C, 210 Kg/cm2)
Influent
liquor
Effluent
Reduction
pH
12.7
12.5
COD, g/L
21.0
1.0
95.2
6005, mg/L
650
Sulfide Sulfur, mg/L
5.8
<0.001
>99.99
Cyanide, mg/L
110
0.035
99,97
Source: Reference 17.
C14-50
-------
As shown in the tables, the treatment -of cyanide bearing waste streams at
high temperatures achieves almost complete cyanide destruction in addition to
high sulfides and COD removal. As indicated by the data, cyanide destructions
of 99.7- percent are typical and, in some cases, total cyanide reductions as
high as 99.995 have been observed.
At present two commercial Zimpro wet oxidation units, one in Japan and
the other Europe, are engaged in treating spent cauatic scrubbing liquors from
petrochemical plants. Another unit, presently under construct ion'in the U.S.,
will be using pure oxygen as the oxygen source. In addition, five WAD units
are currently in service in Japan for the treatment of cyanide-bearing
1 R
wastewaters from acrylonitrile production plants.
14.3.3 Process Costs
Treatment costs for wet air oxidation systems will be affected by a
number of parameters including the amount of oxidation occurring, the
hydraulic flow, the design operating conditions necessary to meet the
treatment objectives, and the materials of construction. These factors
account for the band of capital costs shown in Figure•14.3.2.. The figure was
taken from Reference 2 and updated to reflect changes in the 1982 to 1986
Chemical Engineering (CE) plant cost index. The costs do not include any
costs associated with pretreatment of the feed or post-treatment of the vapor
phase component of the treated liquor. However, post-treatment costs were
included in another capital cost estimate of $2.45 million (adjusted to 1986
4
using the CE plant cost index) for a 20 gpm plant. This estimate is within
the capital cost band shown in Figure 14.3.2.
Operating costs for the wet oxidation unit area shown in Figure 14,3.3.
These data were also derived from data given in Reference 2 vich adjustment .
made for the costs of labor and cooling water. As noted in Reference 2, power
accounts for the largest element of cost. This power cost is primarily the
result of air compressor operation. Additional power for supplying energy for
the oxidation of very dilute wastewaters would be at most 500 Btu/gallon. The
associated costs for this energy would be less than one (1) cent/gallon.
The use of pure oxygen instead of compressed air will help to lower power
costs, particularly with respect to handling and consumption. However,
increased reagent costs may. more> than offset decreased operating costs and
decisions should be made on a case by case basis.
-------
o
_J
_J
trj
O
o
t-
2
O
LJ
1
I
10 20 30 40 50 60
WET OXIDATION PLANT CAPACITY, gpm
70
Figure 14.3.2. Installed plant costs versus capacity,
Source: Reference 2.
-------
IS
ce
UJ
a.
O
4.5
4.0
5,5
3.0
2.5
2.0
f.5
1.0
0.5
1986 cost; Pover-$0.05/KWH; C.W.-$0,25/1000 gallon;
maintenance-1% capitol cost; labor-$30,000/yr/operator.
_ : ; : ' _ POWER
— MAINT,
--LABOR
T C.W.
4.5
4'.0
3.5
3.0
E.5
2.0
1.0
0.5
0 JO 20 30 40 ' 50
WAO UNIT CAPACITY (US 6-PM)
SO
TO
Figure 14.3.3, Unit operating costs versus unit flow rate.
Saurce: Referenda 2.
-------
Total costs, capital plus operating, on a per unit of feed basis,
requires assumptions on life cycle, depreciation, taxes, and current interest
rates for the capital cost. One avenue for financing that has been used
commercially, common lease terms, are 5 years and 20 percent value at end of
term.* Table 14.3.4 illustrates the effect on total costs per unit of feed.
TABLE 14.3.4. WAO COSTS VERSUS FLOW
Hydraulic
flow (gpra)
2.0
10
20
40
Cost elements
Operating ,
23
6
3
2-3
per gallon,
Capital
31
7
5
4-5
cents
Total
54
13
8
6-8
•At Gasmalia Resources, the prices, (April, 1985) for treatment of wastes
are computed based on the oxygen demand of the material. Prices range from a
minimum of $120 per. ton to..a maximum of $700 per, ton versus $15 per ton for
the land disposal of low risk wastes.
14.3.4 Status of Technology
The WAO process is available commercially, and reportedly well', over 150
units are now operating in the field treating municipal and various industrial
14
sludges. The process is used predominately as a pretreatiaent step to
enhance biodegradability. Only a few units are now being 'used to treat
industrial cyanide wastes. These include the unit in California and six other
units currently operating in Japan and Europe.
*Assume lease charges of $17/1,000 per month based on total installed cost,
; 14-54':
-------
The oxidation of specific contaminants in waste streams by the wet
2> ant' particulate. Water scrubbing
and, if need be, carbon adsorption or fume incineration are used to
reduce hydrocarbon emissions or odors.
5. WAO also has application for inorganic'compounds combined with
organics. The oxidation cleans up the mixture for further removal
of the inorganics. WAO can detoxify most of the EPA priority
pollutants. Toxic removal parameters are in the order of
99* percent using short-term, acute, static toxicity measurements.
-------
Limitations of the WAO process relate to the sensitivity of destruction
efficiency associated with the chemical nature of the contaminant, the
possible influence of metals and other contaminants on performance, the
unfavorable economics associated with low and high concentration levels, and
the presence of residuals in both the vapor and liquid phases uhich may
require additional treatment. Costly materials of construction and design
features may also be required for certain wastes which will form corrosive'
reaction products or require extreme temperature/pressure conditions to
achieve destruction to acceptable treatment standard levels. In particular,
chlorinated aromatic compounds are more resistant to degradation and can
result in the production of HC1 byproduct.
-------
REFERENCES
1, Dietrich, M. J., T. L. Randall, and P. J. Canney. Wet Air^Oxidation of
Hazardous Grganics in Wastewater, Environmental Progress, Vol. 4, No. 3.
August 1985.
2, Freeman, H. Innovative Thermal Hazardous Treatment Processes, U.S. EPA,
Hazardous Waste Engineering Research Laboratory, Cincinnati, Ohio, 1985,
3. California Air Resources Board. Air Pollution Impacts of Hazardous Waste
Incineration: A California Perspective. , December 1983.
4, Wilheltni, A. R., and P. V. Knopp, Wet Air Oxidation - An Alternative to
Incineration, Chemical Engineering Progress. , August 1979.
I
5. Zimmerman, F. J., and D, G, Diddams. The Zimmerman Process and its
. Applications in the Pulp and Paper Industry, TAPPI Vol. 43, No. 8.
August 1960,
6. Copa, W,, J. Jeimbuch, and P. Schaeffer. Full Scale Demonstration of Wet
Air Oxidation as a Hazardous Waste Treatment Technology. In:
Incineration and Treatment of Hazardous Waste, Proceedings of the Ninth
Annual Research Symposium, U.S. EPA 600/9-84-015. July 1984.
7. Copa, W., M. J. Dietrich, P. J, Cannery, and T. L. Randall.
Demonstration of Wet Air Oxidation of Hazardous Waste. In Proceedings of
... Tenth Annual Research^Symposium, U.S.: EPA .600/9-84-022. September 1984.,
8, U.S. Environmental Protection Agency, Background Document for Solvents to
Support 40 CFR Part 268, Land Disposal Restrictions, Volume II.
January 1986.
9. Radinsky, J,, et al. California Department of Health Services, Recycling
and/or Treatment Capacity for Hazardous Waste Containing Cyanides. Staff
Report. March 1983.
10. Baillod, C. R., and R. A. Lamporter. Applications of Wet Oxidation to
Industrial Waste Treatment. Presented at 1984 AICHE National Meeting,
Philadelphia, PA. August 19-22, 1984.
11. Baillod, C. R., R. A. Lamporter, and B. A. Barna. Wet Oxidation for
Industrial Waste Treatment, Chemical Engineering Progress. March 1985.
12. Baillod, C. R.,' B, M. Faith, and D. Masi. Fate of Specific Pollutants
During Wet Oxidation and Ozonation, Environmental Progress. March 1985,
13. Randall, T. R. Wet Oxidation of Toxic and Hazardous Compounds. Zimpro,
Inc. Technical Bulletin 1-610. 1981.
14. Telephone conversation with A, Wilhelnti on April 3. 1986.
-------
15. Metcalf & Eddy, Inc. Hazardous Waste Treatment Storage and Disposal
Facility ~ Site Evaluation Report, Casmalia Resources, Casraalia,
California, Publication NS J-1074, April 8, 1985.
16. McBride, J. L.» and J „ A. Heimbuch. Casmalia Resources. Hazardous Waste
Treatment Using Wet Air Oxidation of Casmalia Resources. 1983.
17. Treatment of Spent Caustic Liquors by Wet Oxidation. Zimpro Inc.
Technical Bulletin 3260-T,
18, Wilheltni, A. R. Zinpro Inc. Telephone conversation with Stephen Palmer
of Alliance Technologies Corporation. March 3, 1987.
14.^58'-•
-------
14.4 SULFUR-BASED CYANIDE TREATMENT TECHNOLOGIES
In recent years, sulfur-based cyanide treatment technologies have been
the focus of an increasing number of research efforts and commercial
applications. Sulfur-based cyanide treatment technologies have shown
potential for removing cyanide from aqueous waste streams and are not subject
to many of the limitations associated with more conventional cyanide treatment
technologies. For example, alkaline chlorination, the most common treatment
procedure for cyanide wastes, has the potential for generating hazardous
and/or toxic by-products (i.e., chlorinated aliphatic hydrocarbons) while
2 3
oaonation and wet air oxidation require large capital investments. *
The three sulfur-based cyanide treatment technologies which have shown
the most promise are polysulfide treatment, the INCO process, and ferrous
sulfate treatment. The use of polysulfides for treating cyanide waste streams
4
was first reported in 1940. Polysulfide solutions have been recently
adapted to scrub hydrogen cyanide from fluid catalytic cracking and coking
gases, treat concentrated cyanide electroplating solutions, and remove
4
cyanides present in coal gasification wastewaters. The INCO process and
ferrous sulfate treatment have also shown promise in treating a wide variety
of cyanide wastewaters such as ore leaching and electroplating effluents.
14.4.1 Process Description
14.4.1.1 Polysulfide Treatment —
Polysulfides species are formed when neutral sulfur atoms combine with
monosulfide species. They .can be represented by the chemical formulas-,
H^S , HS . where x = 2 - 5. Equilibrium calculations show that the,
2 x. x
tetrasulfide and pentasulfide species should be the predominant polysulfide
forms in neutral and slightly alkaline solutions, but recent experimental work
has detected only the pentasulfide species.
In the cyanide-polysulf ide reaction it has been postulated that 1 mole
cyanide reacts with 1 mole of polysulfide to produce 1 mole of less toxic
thiocyanate.
CW
SXS~2 = SON" + Sx_1S-2
However, it should be noted that in sufficient quantities, thiocyanates
can cause toxic inhibition to biological treatment systems.
-------
During the reaction, one polysulfide sulfur atom, poly~S°, is reduced from
oxidation state 0 to 1, while the cyanide carbon atom is oxidized from
+ + 5
oxidation state 2 to 3.
The rate equation for the change of total free cyanide ([HCN] + [CN~])
per unit time in the presence of polysulfide is as follows:
-d[CNT]/dt = R (CNTJ [Sx - S~2|
where:
(CNjJ » tHCNj * [CN~], moles/L
|SX - S~^J = polysulfide, tnoles/L
K = reaction rate constant (liter/mole/min)
t = tinei minutes
The role of the hydroxyl ion (OH ) in the cyanide-poiysulf ide reaction
system is unknown at this point in time. The hydroxyl ion may initiate or
impede the oxidation process via a free radical chain nechanism. ' The
following rate relations for the cyanide-poiysulf ide reaction system were
derived by researchers at Carnegie-Mellon University to determine tbe effect
of pH.5
pE = 8.2 ~d[CNT]/dt « 1.41 [em1-04 tpoly-S'J0-85
ptt = 10.0 ~d[CNT]/dt = 0.27 (CN~]°-51 [poly-S° j°- 87
pE = 12.0 "dfCNT]/dt = 0.14 ICIT[0-49 (poly-S°j°-78
The initial rate kinetic data shows that the reaction is mixed order and
that both reaction order and reaction rate change are heavily influenced by pH.
Three common forms of polysulfides are sodium .polysulfide, ammonium
polysulfide, and calcium polysulfide { lisaesulfur). Sodium and ammonium
polysulfide are manufactured according to the following stoichiometric
equations:
H2S + (X-l) S° = Na^
2NH40H + H2S + (X-l) S° = (KH4)2 Sx + 2H20
-------
Limesulfur is commercially available as a commonly used pesticide and
fungicide.
Equipment needs are similar to those, described for other chemical
precipitation processes. Storage tmnks, reaction vessels, agitation,
materials handling, and process control equipment are standard process items.
14.4.1.2 INCO Process-
In 1982 Inco Metals Company announced the development of a technology for
the destruction of cyanide in gold mill waste streams. The process involves
the selective oxidation to cyanate of both free and complexed cyanide species
using a mixture of SO and air at controlled pH in the presence of copper as
a catalyst. Metals are precipitated from solution as hydroxides. The process
also removes iron cyanide, not by oxidation, but by precipitation as an
... , ., 10-12
insoluble copper or zinc f errocyanide.
The oxidation cyanide occurs according to e simplified reaction as
follows: ,
CN~ * S02 + 02 •»• «20 = CNO~
Based on the stoichiometry of this reaction, the SO* requirement is 2.47 g
S0,/g CN oxidized.
The SCU/air oxidation process destroys the metal cyanide complexes
typically present in metal finishing and gold mining effluents. Based on
sequential sampling data from batch experiments, the preferential order of
metal cyanide complex removal is:
2n > Fe > Ni > Cu
The SO /air oxidation system has successfully removed iron cyanide complexes
from solution. * During SCL/air treatment, iron remains in the reduced
ferrous state and is not converted to the ferric state as occurs in stronger
oxidizing environments. Ihe iron cyanide complexes are removed from the
solution by precipitation of metal ferrocyanide compounds of the form
Me-FeCCN), (where Me = Cu, Zn and Ni). » Metals liberated fron the
2o
cyanide complexes of copper, zinc, and nickel are removed by precipitation of
metal hydroxides at the reaction pH.
-------
The cyanide oxidation reaction IB catalyzed by the presence of copper in
solution. Copper for catalysis or for precipitation is conveniently added as
a CuSO, solution. Any free CN present is quickly complexed as « Cu(I)
cyanide complex, which apparently is involved as a catalyst in the oxidation
of CN to CNO by SO. and 0_. The effect of copper concentration in
batch treatment of a synthetic cyanide effluent containing 250 mg/L CN_ is
shown in Figure 14.4.1. The effect of treatment pH is Shown in
Figure 14.4.2. The optimum copper concentration is 50 mg/L and the optimum
operating pH is in the range of 9 to 10 which can be achieved by the addition
of lime.
14.4.1.3 Ferrous Sulfate Treatment—
The formation of less toxic cyanide complexes such as ferro and
ferric-cyanides also has been used as a method for detoxifying of cyanide
wastewaters. This process involves the use of iron salts to form complex
compounds with the free cyanide in the wastes. Eventually these cyanide
complexes are precipitated and removed'as a sludge.
The major advantage of this treatment nethod is that it ia relatively
inexpensive in locations where waste ferrous sulfate is available. However,
considerable quantities of sludge may be formed anil the" "treated solutions are
strongly colored. There also is evidence that ferrocyanides may be decomposed
to free cyanide by sunlight. The regeneration of the cyanide under these
conditions would contaminate the receiving stream.
This method has received very little acceptance by industry in this
country, but appears to be used in Europe. The complexing process apparently
does not completely destroy cyanide under practical operating conditions.
Cyanide levels in treated solutions may be as great as 5 to 10 ppm. Thus, the
sludges formed would appear to be toxic and will require substantial
post-treatment prior to final disposal.
14,4.2 Pretreatnent and Post-Treatment Requirements
Very little information exists in the literature concerning pretreatment
and post-treatment requirements for these processes and their feed streams.
Since the processes are aqueous in nature, filtration or some other solids
removal process may be desirable. Adverse effects such as chemical
-------
0.01
Figure 14.4.1, Effect of copper.
100C
Figure 14.4.2, Effect of pH-
Source: Reference 14. " •' •'
-------
interaction, interference with pump operations, abrasion of internal parts,
and fouling of internal surfaces resulting from existing or formed solids are
possible problem areas, but have not been considered in. the literature.
Similarly, post-treatment requirements for sulfur-based cyanide treatment
technologies as reported in literature have been cursory in nature. Residuals
from polysulfide treatment include thiocyanates and other oxidized sulfur
compounds.^»5 Investigators have found chat at elevated temperatures
thlocyanates are corrosive.4 Therefore, additional steps nay be required
in process trains to prevent water containing thiocyanate from reaching
downstream equipment where heat is applied. Residuals from the INCO process
have not been completely determined as of this time. Further fundamental
investigations are required to define the chemical reaction mechanisms and
kinetics, to determine the stability of the precipitated solids, and to
assess the toxicity of treated effluents. Additional process optimization
studies are also recommended.
The ferrous sulfate-cyanide treatment process suffers from the most
serious residual problems• Since cyanides are merely precipitated from
solution without appreciable oxidation in a voluminous sludge product, the
result is a highly toxic sludge.15 Some type of cyanide destruction' and/or
encapsulation process will be necessary prior to'.final disposal.
14.4.3 Process Performance
14.4.3.1 Polysulfide Treatment--
Currently, process performance data for the polysulfide oxidation of
cyanide complexes have been limited to bench- and pilot-scale studies.
Reaction rates and products depend on solution pH the S02 to 02 ratio, and
the catalytic and inhibitory effects of metal ions and organic compounds.
Laboratory tests performed by Luthy, et al. to determine reaction pathways
showed that, jso reaction occurred' between cyanide and sulfide, however sulfur
in the form of polysulfide reacts relatively quickly with the cyanide. The
reaction order was determined to be 1.54 _*_ 0.25 with a rate constant of
approximately 0.24.-'-^ Complex cyanides were evaluated in the presence of
polysulfide at room temperature. It was observed in a survey test that
Fe(CN)~3 produced no or little thiocyanate in the presence of polysulfide.^
6
-------
Subsequent investigations by Trofe, Page and Luthy, at al. sought to
determine the effects of temperature and catalytic/inhibitory compounds. '
The rate constant was found to double for every 12CC increase in temperature.
Certain metals also had an effect on reaction rate. For example, metals ions
+ 2 +2 +2 +2
such as CA , Mg , Ni , and Zn had a catalytic effect at low -
concentration.
In 1985, Ganczarczyk, et.al. investigated the reaction between calcium
polysulfide and concentrated cyanide solutions from electroplating
operations. Previously, Gancearczyk, et al. had conducted a series of
experiments to investigate the cyanide-polysulfide reaction in a 2 -percent
solution (20,000 tng/L CN ) of sodium cyanide. " The reaction proceeded
very rapidly, both at room temperature and at 3°C. It was 95 percent complete
within 1-hour and cyanide concentrations were nondetectable within 2 weeks at
a cyanide-to-polysulfide ratio of 1:2 by weight.
In the later studies, two different wastewater streams from an
electroplating operation were studied. One wastewater was dragout from a
rinse tank in a capper and cadmium plating process. The second wastewater was
a stripper solution for removal of metal plate (Cu/Ni plate stripping
liquor). ^'Tables 14.4,1 and 14.4'.2 show the liquid phase pollutant- •
concentrations following the cyanide-polysulfide treatment of the
electroplating wastewaters. Upon completion of the experiments, the following,
i • j 1-7
conclusions were made:
• The reaction effectively converted CN~ to SCN~ within 2' to
3 days at 3°C, broke down metal-cyanide complexes, and precipitated
metals generally to the levels required by municipal treatment
systems,
* The polysulfide dosages necessary to achieve these goals was only
about 20 percent higher than stoichiometric requirements, but at
very high CN~ concentrations somewhat higher dosages might be
needed.
» It seems that the cyanide-polysulfide reaction was catalyzed by the
presence of metal-cyanide complexes in wastewater and was only
moderately exothermic.
* SCN~ produced by CN~ conversion was partially lost during the
process,
-------
TABLE 14.4.1. LIQUID PHASE POLLUTANT CONCENTRATIONS (mg/L) IN THE TREATMENT
OF THE STRIPPER WASTEWATER FROM COPPER/NICKEL-PLATING, SERIES 2
(INITIAL CYANIDE-TO-POLYSULFIDE RATIO 1:1.5 BY WEIGHT)
Duration of the experiments
Pollutants
CN~
SON"
Fe
Zn
Cu
Cd
Cr
Mi
Initial
56,200
0
110
7.1
29,020
10.5
0.6
5,130
1 day
87.5
20,330
8.9
3.2
6.5
3.0
NDa
94.1
2 days
25.0
16,260
7.4
1.8
9.3
0.5
NT)
105.0
3 days
31.3
20,910
11.5
1.3
4.1
0.8
ND
5S.8
4 days
50.0
19,460
3,4
0.9
4.0
0.5
ND
92,0
5 days
31.3
14,080
5-5
1.5
3.6
1.0
ND
S3. 5
6 days
28. 1
12,490
3.8
2.2
5.2
1.1
ND
69.4
7 days
31.3
10,750
2.6
1.1
2.2
0.9
ND
66,9
fiND - Nondetectable.
'Source: Reference 17,
TABLE 14.4.2'.
LIQUID PHASE POLLUTANT CONCENTRATIONS .(mg/L) IN THE TREATMENT .-•.:
OF THE STRIPPER WASTEWATER FROM COPPER/CADMIUM PLATING, SERIES 3
(INITIAL CYANIDE-TO-POLYSULFIDE RATIO 1:2.0 BY WEIGHT)
Duration of the experiments
Pollutants
Ctf
sctT
Fe
Zn
Cu
Cd
Cr
Mi
Initial
56,200
0
117.8
7.5
29,920
10.4
0.5
5,020
1 day
30
9,000
5.7
2.2
9.3
2.4
m
25,2
2 days
ma
15,220
5.8
1.2
1.5
0.9
ND
28.4
3 days
N2
7,260
3.0
1.6
2.0
1.0
ND
20.0
4 days
ND
2,030
3.2
1.5
3.3
1.0
ND
17.9
5 days
KB
870
3.1
1.2
2.1
0.7
ND
7.6
6 day s
• ND
4,940
2.7
1.0
2,2
0.6
m
22,1
7 days
ro
'3,490
3.7
1.3
2.0
0.7
ND
20.3
aNB = Nondeteccable.
Source: Reference 17.
-------
Despite the large quantity of experimental data available,, industrial
applications of polysulfide treatment of cyanide bearing wastewaters has been
limited to fluid catalytic cracking and coal gasification effluents. The only
other industrial application reported in the literature is a large commercial
waste treatment facility in California. The cyanide treatment process at
this facility is batch in nature and consists of two 18,000 gallon
•storage/treatment tanks into which cyanide wastes (greater than 100 ppra CN_)
are pumped. The treatment reagent used is calcium polysulfide.• This reagent
is stored in an adjacent fiberglass tank. The amount of reagent required to
complete the cyanide to thiocyanate reaction is predetermined by onsite
laboratory analysis of incoming waste for reactive CN. The process typically
handles approximately 40,000 gallons of waste per month at 50 percent of its
capacity.
14.4.3.2 INCO Process--
Performance data for the INCO process relates to industrial applications
rather than laboratory studies of kinetic properties of the reactants.
Tables 14.4.3 and 14.4.4 present typical INCO process results for selected
gold mill b'arre'ii'and t'aiTihg liquors, and plating -rinse" waters.
Table 14.4,3, which contains gold mill effluent data, show that CN was
consistently removed from feeds containing 40 to 2,000 mg/L down to less than
1 mg/L. Eeagent requirements for these waste streams varied with the type of
feed, but were generally in the range .of 3 to 5 g S02/g CN .for barren
14
solutions and 4 to ? g SO./g CN_ for tailing slurries. Table 14.4.4
shows greater than 99 percent of the CN^ was removed from feeds containing
up to 62,000 mg/L of CN_.
14.4,3.3 Ferrous Sulfate Treatment— -
No data were found for the ferrous sulfate treatment process other than
that reported in Reference 15.
14.4.4 Process Costs
Table 14,4.5 contains cost data developed for an S0_/air oxidation
system (single-stage reactor) sized to treat approximately 34,500 gal/day of
-------
TABLE 14.4.3. SELECTED GOLD MILL BARREN AND TAILING LIQUOR RESULTS
Retention
t line
Strnam (min) Reagent
FEED A
STACE I 97 N,n2S03
FEED 1)
STACE 1 26 S02
STACE 2 26 S02
FEED C
STAGE 1 22 S02
STAGE 2 22 S02
-X FEED I) -
,;£'. STACE 1 GO Nn2S205
• iin/
C» *FEED E
(307.
SOLIDS)
STAGE 1 180 Na2S205
*FEEI) F
(247.
SOLIDS)
STAGE L 15 Na2S205
*FEEU G
(35Z
sol. IDS)
STAGE 1 17 "2K03
STAGE 2 17
I'"
_
9.3
12.5
9.0
9.0
9-5
9.0
9.5
11.8
9.0
-
0.65
-
8.0
10.7
8.0
8.5
CNT
1,680
0.13
420
-
0.11
500
3.0
1.2
2,180
0.43
1,480
1,300
40
0.07
200
6.6
0.2
Assay 9
SCN
820
767
1 , 584
-
1,408
270
•: 220
216
1,820
-
1,380
3.0
87
81
•• •
129
91
92
(mg/L)
Cu
210
0.54
137
13
1.4
55
13
0.4
235
4.4
138
0.1
1.3
0.1
111
16
0.3
or (wt
Ni
0.6
0. 1
1.6
0.2
0.2
53
3.2
0.8
2.0
0.2
1.7
O.t
1.6
.1
-
-
. I)
Fe
2.0
0.1
19
5.2
0.2
66
0.2
0.2
325
0. 1
252
0.7
12.5
0.4
7.0
1
1
Reagents added (g/gCNT)
Zn S02 Lime Cu++
758
3.2 3.20 0 0
71
0.4 5.44 8.16 0
0.2 1.36 2.04 0
53
3.40 4.39 0
0.4 0.85 2.19 0
210
3.8 5.00 4.50 0.25
214
4.4 4.4 0.46
.1 - - -
.1 4.8 7.3 0.91
55 -
.1 6.7 1L.8 0.40
.10 0 0
•Tailing s Luc tie a.
Source: Reference 14.
-------
TABLE 14.4.4. TYPICAL PLATING RTNSE WATER RESULTS
Rntent Lon
t ine
Stream (mill)
FF.EO II
STAGE I 1 ,200
FRED 1
STAGE I 500
STAGE 2 500
FI'.F.n .1
STAGE 1 80
FEED K
STACF. 1 7.2
STAGE 2 7.2
S02
Rc.ngcnt ' Cnnnti-C pll
' 13.
Nn2S205 NaOII 9.
11.
N.i2S205 NnOII 9.
9.
11.
Mn2S20c NaOM 9.
_ — _
S02 I.I HE 9.
S02 LIME 9.
1
0
4
0
1
3
°
0
0
CHT
62,400
12.7 -
1.28D
3.1
2.B
540
1.2
142
.4
.4 .,
Cu
3,
2
Reagents
Assays (mg/L) (B/8CI
added
»T>
Fe Ni Zn Cd Sn S02 Caustic Cu**
600
.6
760
5
4
2
47
6
^ 2
.8
.6
_
.2
.3
.0
.6
26
3
6
1
1
0
18
0
.'i 5,400 - - - -
.0 0.9 - - - 3.5 3.3
.0 2.2 - - 1,500
.2 0.2 - - 100 4.4 1.2
.0 .2 - - 30 0 0
.2 90 - - - -
.2 - 8.4 - - - 4.7 3.5
.0 - 14.3 10. 0 - -
.1 - .1 .1 - 4.7 6.4
.2 - .1 .1, - 1.4 3.2
0
_
0
0
.
0.09
_
0
0
i ^} Source: Reference
-------
TABLE 14.4.5. TOTAL ANNUAL COSTS FOR SO2/AIR OXIDATION OF GOLD MILL
BARREN BLEED SOLUTION
Cost item
Cost (S)a
Reaction system
exclusive of solid-
Liquid separation
facilities
Reaction system
inclusive of solid-
liquid separation
facilities
1, Capital cost
U) Major equipment
(b) Installation at 60X
of major capital
Cc3 Total installed cost
2. Operating cost
(a) Lime § $46/ton
(b) S02
-------
gold mill barren bleed solution (cyanide stream from cyanide leaching
process). Hydraulic retention time in the reactor was 50 minutes to allow
for flow equalization. Operating costs, including reegent consumption, were
based on data generated by INCO in continuous flow laboratory experiments,
Costs of the solid-liquid separation system were estimated with and
without the installation of a Lamella flocculator-elarifier. Reactor design
incorporates a sparger for gas transfer and a turbine-type mixer for
gas-liquid contact. Automatic pH control equipment was provided. Redox
potential in the reactor would be continuously recorded, but not utilized as
an automatic process control variable. No provision was made for a dedicated
lime feed system, dewatering equipment, or sludge disposal. Operator time
associated with the treatment process was based on 2 hours/8 hour shift.
Preliminary data from pilot-scale operation indicate the ORP control may be
feasible, thus reducing labor requirements.
Annual unit treatment costs for this system were approximately
$12.5/1,000 gallons for the reaction system, exclusive of solid-liquid
separation facilities, and $14/1,000 gallons for the system incorporating the
Lamella flocculator/clarifier. While the unit costs for this system are much
higher than those of the alkaline chlorination system shown in-Section 14.1,
the greater influent cyanide concentration (1,300 mg/L) in the gold mill
barren bleed contributes substantially to the higher operating costs. If
equal influent cyanide concentrations were present in each cost model, the
S0?/air process would be much more competitive with alkaline chlorination,
In addition, the SCU/air process provides an added performance benefit since
it is capable of removing any ferro or ferric-cyanides present in the feed
13,14
stream. J
Process costs for the polysulfide and ferrous sulfate treatments of
cyanide-bearing waste streams are not included in this section due to the lack
of reliable cose information.
14.4.5 Status of Technology
14.4.5.1 Availability/Applications—
Sulfur-based cyanide treatment technologies, while not fully developed,
have demonstrated potential for the treatment of cyanide wastes. Both reagent
and equipment requirements are straightforward and simple. Application to
-------
industrial wastes is presently limited, but both polysulfide and INCO Process
technologies have demonstrated high efficiencies in treating dilute and
concentrated aqueous cyanide waste..streams.
Licensing of the INCO Process is handled through INCO Tech, a Division of
INCO, Ltd, Licensing fees are claimed to be a modest fraction of the
operating costs.
14,4.5.2 Environmental Impacts—
The environmental impact of the processes discussed here relate to the
unreaeted contaminants and byproducts (thiocyanates) remaining in the waste
stream. Additional treatment to prevent corrosion and minimize thiocyanate
concentrations probably will be required. Air emissions associated with the
use of these technologies will be minimal, although some care must always be
observed in pH adjustments to prevent hydrogen cyanide evolution.
14.4.5.3 Advantages and Limitations—
The advantages of sulfur-based processes discussed here result from ease
and simplicity of operation. Capital investments are low, relative to other
cyanide oxidation processes, and reagent consumption is also low (due to
nonoxidation of SON to carbon dioxide and nitrogen dioxide). Disadvantages
are the result of incomplete destruction and the need for subsequent treatment
of the partially oxidized waste stream.
-------
REFERENCES
1. • U.S. EPA. • Treatability Manual, Volume III. EPA-600/8'-80-Q42c.'•
July 1980.
2. Cusbnie, G. C. Cencec Corporation, Navy Electroplating Pollution Control
Technology Assessment Manual. CR 84-019. February 1984.
3. Freeman, H. Innovative Thermal Hazardous Treatment Processes. U.S. EPA,
Hazardous Waste Engineering Research Laboratory, Cincinnati, OH, 1985.
4. Ehmke, £» F. Tosco Corporation. Polysulfide stops FCCU corrosion.
Hydrocarbon Processing. July 1981.
5. Luthy, R. G., and S. G, Bruce. Carnegie-Mellon University, Kinetics of.
Reaction of Cyanide and Reduced Sulfur Species in Aqueous Solution.
Environmental Science and Technology Vol. 13, No. 12. December 1979.
6. Trode, T. W., G. C. Radian Corporation. Cyanide Removal in Coal "
Gasification Wastewater Using Polysulfide. 40th Industrial Waste
Conference, Purdue University. 1985.
7. Roeck, D., and M. Gollands. Alliance Technologies Corporation.
Hazardous Waste Treatment Facility Site Visit Report. September 1985.
8. Mitre Corporation, Manual of Practice for Wastewater Neutralization and
Precipitation. EPA-&OQ/2-81-148."v::Aueust 1981.' ' ? •* '
9. Cushnie, G. C. Removal of Metals from Wastewater: Neutralization and
Precipitation. Pollution Technology Review. No. 107. Noyes
Publications, Park Ridge, NJ. 1984.
10. Devuyst, E. A., V. A. Ettel, and G. J. Borbely. New Method of Cyanide
Destruction in Gold Mill Effluents and Tailing Slurries, Presented at
the 14th Annual Operators Conference of the Canadian Mineral Processors
Division of the CIM, Ottawa, Ontario. January 1982.
11. Devuyst, E. A., V. A. Ettel, and G. J. Borbely." New Process for
Treatment of Wastewaters Containing Cyanide and Related Species.
Presented at the 1982 AIMS Annual Meeting, Dallas, Texas. February- 1982.-
12. Devuyst, E. A., B. R. Conard, and V. A. Ettel. Pilot Plant Operation of
the Inco S02/Air Cyanide Removal Process", presented at the 29th
Ontario Industrial Waste Conference, Toronto, Ontario. June 1982.
13. Nutt, S. G., and S. A. Zaidi. Treatment of Cyanide-Containing
Wastewaters by the Copper Catalyzed SO£/Air Oxidation Process. 38th
Industrial Waste Conference, Purdue University. 1983.
14. Devuyst, E, A., B. R. Conard, and W. Hudson. Commercial Operation of
INCO's S02/Air Cyanide Removal Process, Conference on Cyanide and the
Environment Tucson, Arizona. December 1984.
-------
15. Schiller, J. £. Removal of Cyanide and Hetale from Mineral Processing
Wastewaters. Bureau of Mines Report of Investigations No, 8836. 1983.
16. Nuthy, R. G., et al. Carnegie-Mellon University, Cyanide and Thiocyanate
in Coal Gasification Wastewater. Journal of the Water Pollution Control
Federation, Vol. 51, No. 9. September 1979.
17. Ganczarczyk, J. J., P. T. Takoaka, and D. A. Ohashi. Application of
polysulfide for pretreacment of spent cyanide liquors. Journal of the
Water Pollution Control Federation, Vol. 57, No. 11. November 1985.
18. Ganczarczyk, J. J. Biological Decomposition of Thiocyanate. Paper
presented at the Symposium Canadian de la Recherche sur la Pollution de
L'eau, St. Foy, Quebec. November 1977.
19. Ganczarczyk, J. J. Fate of Basic Pollutants in Treatment of Coke-Plant
Effluents. Proc. 35th Purdue University. Ind. Waste Conference 325.
1980.
-------
14.5 MISCELLANEOUS CYANIDE DESTRUCTION PROCESSES
\
A variety of noncon.venti.onal or experimental processes are being studied
for the treatment of cyanide-bearing vastes. The cyanide treatment
technologies examined are: the Modar Process, the use of chemical oxidizing
agents with and without catalysts, and catalytic oxidation.
14.5.1 Process Description
14.5.1-1 The Modar Process—
Supercritical fluid oxidation (the Modar Process) is a technology that
has been proposed for the destruction of organic contaminants in wastewater's.'
It is basically an oxidation process conducted in a water medium at
temperatures and pressures that are supercritical for water; i.e., above 374°
(705°F) and 218 atmospheres. In the supercritical region, water exhibits
properties that are far different from liquid water under normal conditions;
oxygen and organic compounds become totally miscible with the supercritical
water (SOW) and inorganic compounds, such as salts, become very sparingly
•soluble,. When these materials are combined in'the SCW process,- organ i'cs are1'
oxidized and inorganic salts present in the feed or formed during the
oxidation are precipitated from the SCW.
The oxidation reactions proceed rapidly and completely. Eeacti'on times
are less than 1 minute, as comapred to reaction times o£ about 60 minutes used
in the subcritical wet air oxidation (WAG) process. Moreover, the reaction is
essentially complete. Carbon, nitrogen and hydrogen atoms within the organic
contaminants are reacted to form C02, NC>2 and H?*-1 (residuals such as the •
low molecular weight organic acids and alcohols found in the treated MAO
effluent are not found in the SCW process effluent). Heteroatoms- - '
(e.g., chlorine and sulfur) are oxidized to their corresponding'acidic auion
groupings. These anions, and those occurring naturally in the feed, can be
neutralized by cation addition to the feed, and the total inorganic content o£
the waste, save that soluble in the SCW, can be precipitated and recovered by
mechanical separators operating SCW conditions.
-------
Presently, the Modar process has not been dedicated .to cyanide
destruction. However, the high oxidation efficiency and rapid reaction rate
of the SCW process .in treating other organic compounds warrants further
investigation. For s further discussion of SCW technology see References 1-7.
14.5.1.2 Other Chemical Oxidizing Agents—
As shown previously in Table 14.1.1, hydrogen peroxide, H_0_, and
potassium permanganate, KMnO,, are both relatively strong oxidizing agents.
Hydrogen peroxide has been used to treat phenols, cyanides, sulfur compounds,
and metal ions in dilute waste streams. Potassium permanganate is primarily
used for the treatment of phenols. The choice of these and other oxidants is
dependent upon such factors as toxicity, reaction rate, ease of removal of
secondary products, simplicity and cost.
Oxidation with H_0. is generally performed in the presence of a metal
catalyst. Typical catalysts include ferrous sulfate, nickel salts, and
aluminum salts. The waste is heated and then treated with H_0,, while
being agitated,. The H^CK oxidation tends to proceed quickly under basic
10
conditions. The feasibility of ultraviolet catalyzed H^CL oxidation
has been studied, but it does not appear to be used on an industrial
scale. Potassium permanganate oxidation is favored under basic •
conditions. Raising the pH to the optimum level is accomplished by the
addition of lime, soda ash, or caustic soda.
The equipment required for chemical oxidation is very simple. This
includes storage vessels for the oxidizing agents and perhaps for the waste,
metering equipment for both streams, and vessels with agitators to provide
contact between the oxidant and the waste. Some instrumentation is required
to determine the concentrations of pollutants, pH., and the degree of
completion of the oxidation reaction. The process is usually monitored by an
12
oxidation reduction (ORP) potential electrode.
For the treatment of sodium, potassium, zinc, and cadmium cyanide, a.
hydrogen peroxide solution with formalin may be used to reduce the cyanide
(Kastone Process). This process is usually operated at ambient
temperature and a pH between 10 and 11.5. The effluent from this process has
a high biochemical oxygen demand and requires biological treatment before
direct discharge to sewers.
.",14-M.
-------
The treatment o£ cyanide waste streams with alternate oxidizing agents
has been limited to batch processes or low effluent flows. The treatment of
large effluent flows is generally not practicable because of a lack of
suitable means of determining the correct dosage quickly and accurately enough
to allow efficient use of the reagent.14 Other limitations include chemical
interference from other oxidizable species, limited shelf life (H^o^) ;
inability to effectively oxidize cyanide beyond the cyanate level, and the
need for catalysts. Therefore, the use of this technology is restricted to
process situations where alkaline chlorination would not be feasible, i*e.,
waste streams containing phenols or aliphatic hydrocarbons,
14,5.1.3 Catalytic Oxidation—
One of 'the earliest investigations of catalytic oxidation was conducted
by Battelle Laboratories in 1971 to study the adsorption of free and cotnplexed
cyanide onto activated carbon in the presence of copper.*5 Subsequent
efforts were undertaken by the Calgon Corporation to also develop a cyanide
detoxification method utilizing catalytic oxidation on granular activated
carbon.*°, Cupric ions are added to the wastewater along with oxygen prior
to passing the cyanide-bearing waste through a granular.activated carbon . .
column. According to Calgon, "cupric ions are added to the water to
accelerate and increase the efficiency of'the catalytic oxidation of cyanide
by granular activated carbon." In addition to improving the catalytic
oxidation of the cyanide, "the presence of cupric ions results in the
formation of copper cyanides, which have a greater adsorbability capacity than
copper or cyanide alone."16
14.5.2 Proc_egs_Per f qrmance_
According to a 1979 survey of 216 metal finishing plants'practicing
cyanide oxidation, three plants were found to be using- hydrogen peroxide as an
oxidizing agent. The process used by these plants is a proprietary treatment
called the Kastone process. J
The Kastone hydrogen peroxide oxidation treatment process treats both the
cyanide and metals in cyanide wastewaters containing zinc or cadmium. In this
process, cyanide rinse waters are heated to 49-54°C (120-130°) to break the
cyanide complex, and the pH is adjusted to 10.5-11.8. Formalin (37 percent
-------
formaldehyde) is added, while the tank is vigorously agitated. After 2 to
5 minutes, a proprietary formulation (41 percent hydrogen peroxide with a
catalyst and additives) is likewise added. After an hour of mixing,1 the
reaction is complete. The cyanide is converted to cyanate and the metals are
precipitated as oxides or hydroxides. The metals are then removed from
solution by either settling or filtration.
In terms of waste reduction performance, the Kastone process was found to
be capable of reducing the cyanide level to less than 0.1 mg/L and the zinc or
cadmium to less than 1.0 mg/L. Table 14.5.1 presents performance data for a
treatment process using hydrogen peroxide to treat, gold mill tailings. The
process uses an excess of HjO- to achieve rapid 'oxidation of cyanide ions
to cyanate. A slight reduction in pB (0.2 to 0.3 units) was found to take
place during the reaction. The concentration of available cyanide was reduced
to less than 0.5"mg/L, but could be reduced to 0.1 mg/L at the cost of
increased H (}„ consumption (present consumption is 0.4 to 1.0 L H_02
70 percent/cubic meter of tailings). Catalytic oxidation effectiveness is
shown .in Table 14.5-..2, • While the results demonstrated that cyanide could be
effectively adsorbed (80 to 99 percent) by activated carbon, regeneration
•efficiencies were poor (1.2 to 28 percent), and residuals; remained in the toxic
cyanide state.
Tables 14.5.3 through 14,5.5 show the catalytic oxidation adsorption
data for the treatment of copper, zinc, and cadmium wastes determined by
Calgon. While Calgon has not pursued the implementations of this technology
on a commercial scale, research into possible applications has continued.
14.5.3 Process Costs
Due to the current level of development of these technologies, limited
cost data are not available. Major costs would be associated with process
equipment and the cost of chemical reagents. Pretreattnent, operating, and
post-treatment costs are unknown, but are expected to be similar to alkaline
chlorination (Section 14.1) in the case of hydrogen peroxide and carbon
adsorption (Section 8.1) for catalyzed oxidation.
-------
TABLE 14.5.1. HYDROGEN PEROXIDE TREATMENT OF GOLD HILL TAILINGS
Before ^02 After H202
treatment treatment •' -..^
Tailings flow m /h (nominal)
Solids content % (nominal)
PH
Free cyanide tng/L
Basily-liberatable cyanide mg/L
Total cyanide rog/L
Dissolved Cu mg/L
Zn uig/L
Fe mg/L
1,100 1,100 .
• AS 45 . / ".
10.5 - 11.0 10.2 - 10.8 • -
. 50 - 100 ' undetectable
90 - 200 .
-------
TABLE 14.5.2. RESULTS OF MULTIPLE CYCLE ADSORPTION AND REGENERATION RUNS ON CONCENTRATED
ZINC CYANIDE WATERS :
Typical concentrations, ppm
Cycle J , 2,460 gal
treated
Cycle 2, 1,770 fj.-il
treai ed
Cycle .1, 1 ,800 gal
t roal.ed
_,-, t
! i'A Cycle 4, 3,540 r,"l
,g: t reared
Cycle r>, 4,100 pal
trcal.ed
Component
Cyanide, CN
Copper, Cu
Iron, Fe
Cyanide, CN
Copper, Gi
Iron, Fe
Cyanide, CN
Copper, Cu
Iron, Fe
Cyanide, CN
Copper, Cu
Irnn, Fe
Cyanide, CN
Copper, Cu
Irnn, Fe
Tn feed
220-340
240-317
6-17
239-364
145-252
0.1-16.7
203-270
164-214
1.3-8.5
234-436
170-365
0.3-33
333-468
224-330
0.3-21
Effluent
average
1
1
2
30
1
2
50
1
"*—
2
1
0.3
2-3
I
0.3
"Effluent at
breakthrough
61
30
12
94
43
0.2
60
32
0.2
31
5.8
—
—
Weight,
adsorbed
5.88
5.47
0.33
(4.93)
4.10
~ ~
3.73
3.06
7.36
6.72
0.26
10.15
9.57
0.33
pounds
stripped
1.67
5.06
0.03
1.05
3.09
0.01
0.97
2.B7
0.18
0.89
5.59
0.13
—
— T
Efficiency, percent
adsorpt ion
99
99
90
99
—
80
99
~~"~
98-99
99
—
98-99
99
"
regeneration
28
92
10
21
75
• ~—
26
94
"
12
83
50
—
"
Source: Reference 15.
-------
TABLE 14.5.3. COPPER CYANIDE WASTE TREATMENT USING CATALYTIC OXIDATION
Days on
stream
1
2
3
4
5
6 '
7
'8 "
9
-- 10
11
12
13
14
15
Inf luen
f'Kl"™"
tn
(mg/L)
32
32
32
32
32
32
28'
""28 ' ''
28
28
28
30
30
30
. 30
t
CN~
0.01
0.01
0.01
0.01
0.01
0.02
0.04
'"***' ''o'.dff • ' '
0.02
0.77
0.80
0.32
0.25
0.10
0.10
Effluent (mg/L)
Cu
0.05
0.05
0.05
0.05
0.05
0.05
0.05
'• 0.05"
0.05
0.05
^0.05
0.05
0.05
0.05
0.05
• Fe
0.05
0.05
. 0.05
0.01
0.01
0.05
0.05
0.05"
0.05
0.45
0.35
, 0.15
0.10
0.05
0.10.
-------
TABLE 14.5.4. 21NC CYANIDE WASTE TREATMENT USING CATALYTIC OXIDATION
Days 00
stream
1
2
3
4
5
6
7
8 ......
9
10
11
12
13
14
15
Influent
Img/L)
22.0
22.0
22.0
22.0
22.0
22.0
30.0
30.0,
30.0
30.0
25.6
25.6
25.6
25.6
25.6
CN~
0.01
0.02
0.03
0.01
0.04
0.04
0.08
0.05
0.14
0.04
0.04
0.07
0.03
Effluent
Cu
0.05
0.05
0.05
0.05
0.05
0.05
0.05
0.05
0.05
0.05
0.05
0.05
0.05
0.05
0.05
Zn
0.05
0.05
0.05
0.05
0.05
0.05
0.05
0.05
0.05
0.05
0.05
0.05
0.05
0.05
0.05
Fe
0.05
0.05
0.05
0.05
0.05
0.05
0.05
0.05.
0.05
0.05
0.05
0.05
0.05
0.'05
0.05
-------
TABLE 14.5.5. CADMIUM"CYANIDE WASTE TREATMENT USING CATALYTIC OXIDATION
Ij3 ys on
scream
1
2
3
4
5
6
7
'8 -• '
9
10
11
12
13
14 ;
15
• Influent
(mg/L)
30 •
30
• 30
30
30
30
29
•'29 ''
29
29
• 22 ,
' ' 22
• - 22
• 22
22
CN~
0.01
0.04
0.01
0.01
0.01
0.08
O."04'v'-
0.06
0.03
0.06
0.11
• 0.14
0.06
0.03
Effluent
Cu
0.05
0.05
0.05
0.05
0.05
0.05
0.05
0.05
0.05
0.05
0.05
0.05
0.05
0.05
0.05
(mft/L)
Zn
0.05
0.10
0.20
0.05
0.20
0.05
Q.2Q£
0.20
0.10
0.05
0.10
0.10
0.05
0.05
Fe
0.05
0.05
0.05
0.05
0.05
0.05
0.05
0.05
0.05
0.05
0.05
0.05
0.05
0.05
-'0.05
-------
14.5.4 Overall Process Status
Other Chan hydrogen peroxide oxidation, the commercial 'application of •
these processes to free and total cyanides has yet to be established. Some
level of destruction can be expected, but economic considerations have limited
application. Hydrogen peroxide oxidation has been commercially applied, but
typically to waste streams in which chlorine or hypochlorite oxidation would
not be feasible. Hydrogen peroxide oxidation has limited application to
slurries, tars, and sludges. This is due to the presence of other oxidizable
components in the sludge which may be attacked indiscriminately by the
oxidizing agent, thus increasing reagent consumption.
The environmental impact of the processes discussed here relate to the
unreacted contaminants and byproducts remaining in the waste stream.
Additional treatment usually will be required. Air emissions associated with
the use of hydrogen peroxide and permanganate oxidant will be minimal,
although some care must always be observed when the contaminants are high
vapor pressure solvents and ignitables.
The advantages of the oxidation processes discussed here result from ease
and simplicity of operation. Disadvantages are- the result of incomplete
destruction and the need for subsequent treatment of the oxidized waste stream.
-------
REFERENCES
1. Josephson, J. Supercritical Fluids. Environmental Science and
Technology. Volume 16, No. 10. October 1982.
2. Thomason, T. B. and M. Model!. Supercritical Water Destruction of
Aqueous Wastes, Hazardous Waste. Volume 1, No. 4. 1984.
3. Franck, E. U. Properties of Water in High Temperate, High Pressure
Electrochemistry in Aqueous Solutions (NACE-4). p. 109. 1976.
4. Sieber, F. MODAR Inc. Review of Draft Section, .Supercritical Water
Oxidation. May 16, 1986.
5. National Academy of Sciences, Medical and Biological Effects of
Environmental Pollution." Nitrogen Oxides. 1977.
6. Modell, M., G. Gaudet, M. Simeon, G. T. Hong, and K. Bieiaann.
Supercritical Water Testing Reveals New Process Holds Promise. Solid
Wastes Management, August 19fl2.
7. Freeman, H. Innovative Thermal Hazardous Waste Treatment Processes.
U.S. EPA, HWERL Cincinnati, OH. 1985.
8. Prengle, H. W., Jr. Evolution of the Ozone/UV Process for Wastewater
Treatment.,,.. Paper presented at Seminar on Wastewater Treatment and
Disinfection with Ozone., Cincinnati, OH. September 15, 1977.
International Ozone Association, Vienna, VA.
9. Harris, J. C. Ozonation. In: Unit Operations for Treatment of
Hazardous Industrial Wastes. Noyes Data Corporation, Park Ridge, N.J,
1978.
10, Hackman, E. Ellsworth. Toxic Organic Chemicals-Destruction and Waste,
Treatment. Park Ridge, NJ, Noyes Data Corporation. 1978.
11. Sundstrom, D. W., et al. Destruction of Halogenated Aliphatics by
Ultraviolet Catalyzed Oxidation with Hydrogen Peroxide. Department of
Chemical Engineering, The University of Connecticut. Hazardous Waste and
Hazardous Materials, 3(1). 1986.
12. Chillingworth, M. A., et al. Industrial Waste Management Alternatives
for the State of Illinois, Volume IV - Industrial Waste Management
Alternatives and their Associated Technologies/Processes, prepared by
Alliance Technologies Corporation. February 1981.
13. Raditasky, J., et al. California Department of Health Services.
Recycling and/or Treatment Capacity for Hazardous Waste Containing
-------
14. Knorre, M., and A. Griffiths. Cyanide Detoxification with Hydrogen
Peroxide Using the Degussa Process. Cyanide and the Environment
Conference Proceedings, Tucson, Arizona. December 1984.
15. Battelle Laboratories. An Investigation of Techniques for Removal of
Cyanide from Electroplating Wastes. Water Pollution Control Research
Seriei. 12010 EIE, November 1971.
16. Bernardin, F. E. Cyanide Detoxification Using Adsorption and Catalytic
Oxidation on Granular Activated Carbon. Journal of the Water Pollution
Control Federation. 45(2). 1973.
17. Huff, J. E., and J. Bigger. Cyanide Removal from Refinery Wastewater
Using Powdered Activated Carbon. U.S. EPA Draft Report on Grant
No. R804029-01. 1976.
-------
-------
SECTION 15
MISCELLANEOUS CYANIDE DESTRUCTION PROCESSES
The main miscellaneous cyanide destruction processes are biodegradation
and thermal treatment, Biodegradation as a process for treating wastes
containing cyanide is still in the developmental stage. Certain types of
microorganisms have shown the ability to completely degrade low concentrations
of simple cyanides. The major obstacle to implementation has been the
inability of most conventional biosystetns even when acclimated, to degrade
fixed cyanides or simple cyanides in high concentrations. However, since the
end products of complete biodegradation are nontoxic, continued research is
advisable. In addition, many of the new bioaugtnentation processes which can
degrade fixed and/or concentrated cyanide wastes, may render biological
treatment as a feasible alternative to conventional chemical or''thermal
destruction technologies.
Thermal treatment technologies which may be applied to cyanide-bearing
hazardous wastes include incineration, evaporation, and crystallization. The
processing systems involved in each of these technologies are similar to those
described for management of metal~bearing hazardous wastes. Test studies have
indicated high potential levels of waste destruction (i.e., in excess of
99.99 percent) for the incineration of cyanide wastes. Incineration is most
typically used to destroy cyanide wastes generated in organic chemical
manufacturing; e.g., acrylonitrile production. Other cyanide waste candidates
for incineration are waste organic cyanide compounds such as cyanogen.
Cyanide waste component recovery by evaporation or crystallization has been
demonstrated to achieve yields in excess of 90 percent in certain cases.
15-1
-------
15.1 BIOLOGICAL DESTRUCTION OF CYANIDES
15.1.1 Process Description
Microbiological degradation of cyanide is a developing technology,
capable of oxidizing low concentrations of simple cyanides into carbon dioxide
123
and ammonia. ' ' The process offers several advantages over ocher methods
of degrading or detoxifying cyanide-bearing waste streams. For exatnple5
hazardous chemicals such as caustic soda, gaseous chlorine, and hypochlorite
4
salts are not required, thereby reducing exposure risks. In addition,
toxic by-products and/or sludges are not generated during processing,
eliminating secondary treatment for cyanides; howeyer, removal of the
non-toxic by-products and/or sludge is still required.
However, most conventional cyanide biodegradacion systems are only able
to treat total cyanide concentrations of approximately 10 mg/L or less without
noticeable performance impairment. Furthermore, only free cyanide is
biodegradable, with waste streams containing fixed cyanides experiencing the
lowest removal rates". These drawbacks have limited the application of
biological treatment for'1 cyanide-containing wastes.
The principal factors which control the microbial degradation process are
moisture level, organic content, oxygen level, temperature, pH, and nutrient
source. Typical design factors include BOD and toxic constituent removal
rate, detention time, reactor surface area and type, nutrients required to
sustain biological activity, and sludge production. Operating.parameters,
pretreatment and post-treatment requirements, and process costs have been
presented in Section 11.0. Therefore, the remainder of this section will
focus on performance and technological status of biological treatment of
cyanides.
15.1.2 Process Performance
In 1982, wastewaters from a benzol plant were biologically treated in an
upflow biotower (UBT). The results were compared to 12 previous studies
performed with similar wastewater in activated sludge and other types o£
fixed-filter reactors. Figure 15.L.I summarizes percent cyanide removal and
' 15-2
-------
-J
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o
5
LU
cc
2
U
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5-
_J •
•v
5
z"
o
H
2
U
3
_J
LL.
U.
LU
IMU
90 -
80 -
70 -
60 -
50 -
40 -
30 -
10 -
6 -
'5 -
4 -
3 -
2 -
1 -
0.8 -
0.6 -
0.4 -
0.2 -
0 -
/~N^>l(5) ' REFERENCES
l^A^ /gN • UBT - A
_ ^. AS'4 1,2,35,7,9,11,125,13
XChr ©^ B8C"3>12 ' -:
v!y ^ UJ FLUIDIZEO - IB
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/— s
®
/"~\
(li)
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^0 ^s)
/~\ /~N
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rn ©
A-^« ' UBT - Up f Low Biological Touer
\l) VX /->. AS - Activated Sludge
\sli/ RBC - Rotating Biological Contactor
©Fluidized - Fluidized Bed
_^
1111
0 100 200 . 300 400 50
COD LOADING
Ib/ 1000 ft3/DAY
Figure 15.1,1. 'Cyanide removal vs. organic load
for various biological reactors,
Source: Reference 5.
15-3
-------
effluent concentrations as a function of organic loading. The data indicate
that cyanide removal was only partial for the UBT, and show wide variation
between the 13 studies. The inconsistency is thought to be due, in part, to
the fact that investigators sometimes recorded total cyanide concentrations
rather than free cyanide concentrations.
However, in the overwhelming majority of cases studied, the influent
cyanide concentration was less than 10 mg/L and less.than 90 percent cyanide
removal was achieved, These results suggest that conventional biological
systems are not capable of reducing cyanide concentrations to acceptable
discharge levels.
To remedy this problem, some research has been conducted to develop
organisms specifically designed to degrade cyanide at levels which would
normally be toxic to conventional systems. For example, Imperial Chemical
Industries has recently tested and marketed the enzyme foraamide hydrolase for
this purpose. Commercial trials were initiated in the summer of' 1985 on a
continuous system that provides a 2-hour residence time, 95 percent enzyme
recycle, and an enzyme concentration of 10 g/L. Temperatures were
approximately 30 to '35°C and the waste was maintained at a pH of 8 to 8.5. -
The trial results showed that organic cyanides (nitriles) could be
effectively treated after pretreatment with alkali. Some metal cyanide
complexes are treatable, but the stronger ones (e.g., iron and copper) proved
to be more resistant. The maximum concentration which could be handled was'
10,000 ppm. The optimal feed concentration was found to be 5,000 ppm with
reduction to 10 ppm achieved in 6 hours. Further reduction was reportedly
possible (e.g., 1 ppm) but no data were presented to document this claim.
Cyanide is degraded to formamide and eventually to ammonia and carbon dioxide.
Researchers at Homestake Mining Company in South Dakota have developed a
strain of Pseudomonas paucimobi-lis that oxidizes the free-,and complexed
cyanides and thiocyanates from the mine's wastewaters. After 7 years of
bench- and pilot-scale evaluations, the process was commercialized at a
5.5 million gallon per day plant in the summer of 1984. The final design uses
the strain and 48 rotating biological contactors. Among the alternative
i
processes investigated by Homestake prior to commercialization were
acidification/volatilisationj ozonation, ion exchange, Prussian blue
15-4
-------
oxidation/precipitation, carbon adsorption, alkaline chiorination, and
copper-catalyzed hydrogen peroxide (DuPont Kaetone process). Biological
processes investigated included activated sludge, suspended growth, and
several attached growth processes, including trickling filters, biological
towers, aerated biological filters, and rotating biological contactors.
The only chemical requirements for the process are soda ash and
phosphoric acid. Products of the biological degradation are relatively
harmless anions such as sulfates, nitrates, and carbonates.. Reportedly,
ammonia is not released as a by-product. Kinetics are first order until low
levels of the pollutants are reached.
15.1.3 Procesa_S>ca_Eu_s
Currently, biodegradation of wastes containing cyanide is still in the
developmental stage. Certain types of microorganisms have shown the ability
to completely degrade low concentrations of simple cyanides. The major
obstacle to implementation has been the inability of most conventional
biosystems, even when acclimated, to degrade fixed cyanides or simple cyanides
in high concentrations. However, since, the end products of compl'ete*
biodegradation are nontoxic, continued research is advisable. In addition,
new bioaugtnentation processes which degrade fixed and/or concentrated cyanide
wastes appear to have substantial potential as an alternative to conventional
chemical or thermal destruction technologies.
15-5
-------
REFERENCES
1. Howe, R.H.L., "Bio-Destruction of Cyanide Wastes-Advantages and
Disadvantages," Jour. Air Wacer Pollution, Pergaraon Press, 9(1965).
2. Key, Arthur, "Gas Works Effluents and Ammonia,"' Institute of Gas
Engineers, Britain (1938).
3. Painter, H.A., and Ware, G.C., Nature, London, 175,900 (1955),
4, Alliance Technologies Corporation. Treatment Technologies for Dioxirt-
Containing Wastes. Contract No. 68-03-3243. August 1986.
5. Olthof, M., Oleszkiewicz, S. Benzol Plant Wastewater Treatment in a
Packed-Bed Reactor. 3?th Industrial Waste Conference. Purdue
University. 1982.'
6. Technical .Insights. New Methods for Degrading/Detoxifying Chemical
Wastes. Englewood, N.J. 1986-. . .. ,, . _.. .• .-,.
7, Shahalotn, A.M. Mixed Culture Biological Activity in Water Containing
Variable Low Concentrations of Cyanide, Phenol, and BOD. 38th Industrial
Waste Conference, Purdue University. 1983.
15-6
-------
15,2 THERMAL PROCESSING OF CYANIDE-BEARING' WASTES
Several of the thermal processes outlined in Section 12 for treatment of
octal-bearing hazardous wastes may also be considered as alternatives for
treatment of hazardous wastes containing cyanides. As discussed in
Section 12, many of the cyanide-bearing hazardous wastes are generated by
essentially four different industries. Wastes from electroplating and metal
finishing operations comprise by far the greatest percentage of the overall
volume of cyanide-bearing waste. Much, if not all of those wastes also
contain heavy metals swch as chromium, nickel, and-lead. The other important
sources of cyanide-bearing waste include the metallurgical coke industry
(whose wastes also contain metal constituents), aerylonitrile or
cyanide-related organic compound manufacturing, end the manufacturing of
cyanide salts such as sodium cyanide or potassium cyanide- The
characteristics of cyanide—containing wastes from the organic chemical
industry would appear to be most suitable for treatment by thermal processes
such as incineration.
Economic and environmental factors constitute the most significant
barriers to selection of inciner'atiori for treatment of cyanide-bearing
wastes. Further, many commercial incinerators surveyed do not handle cyanide
wastes at all, citing emissions of deadly cyanide gases (e.g., HCN) as a major
concern. Such gases would require extensive environmental control and safety
precautions, including secondary incineration. Given the expense of such
procedures relative to available chemical treatment processes, most of the
commercial waste processors surveyed could not recommend incineration as an
Option for cyanide-bearing wastes. Finally, such systems may also'generate
high levels of NO emissions, and solid and liquid waste streams requiring
additional control.
15.2.1 Process Descriptions
The incineration technologies which may be employed for the disposal of
cyanide-bearing hazardous wastes are similar to those identified in Section 12
of this report, and described in detail in Reference 2. Incineration and
15-7
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pyrometallurgical processes would lead to Che destruction of the cyanide
group; evaporation would result in increased concentration levels. No
application of crystallization for recovery of cyanides was identified in
the literature.
15.2-.,2 Performance Data
As detailed in Section 12, the relative ease with which hazardous wastes
containing cyanides may be incinerated has been studied within the context of
a general study of hazardous waste incinerability conducted by EPA.3 A
summary of the "incinerability" ratings developed by EPA for such wastes is
presented in Table 15.2.1. As shown, many of the organic cyanide-bearing
hazardous wastes are considered to be at least "low" potential candidates for
incineration. In addition, it may be noted that most of the "incinerable"
wastes may be burned in either two or three of the most widely used
incineration systems. The incinerability ratings are somewhat consistent with
data compiled in a 1981 study of incineration risk analysis prepared for EPA
fey IE Inc.,* a similar study conducted in 1984 by IGF,5 and a 1982 study
prepared for EPA by MITRE, Inc. ,6 in which the quantities of waste currently
incinerated we're estimated. These da'ta indicate that high volumes of waste
containing cyanides from certain industries are incinerated, including
acrylonitrile manufacturing, and paint production. A summary of these data is
shown in Table 15-2.2. Clearly, many cyanide-bearing hazardous wastes are
incinerable and are currently being incinerated. Relative to the overall
volume of cyanide-bearing waste, however, the amount incinerated .is very small.
The "incinerability" of cyanide-bearing wastes may be evaluated through
assessment of a variety of key waste characteristics. These include:
o Physical form;
o Heat content/heat of combustion;
o Autoignition temperature/thermal stability;
o Moisture content;
o Organic content;
15-8
-------
TABLE 15,2.1,
RANKING OF INCINERABILITY OP
CYANIDE-BEARING WASTES
Waste code
description*
Ranking*1
Applicable technologyc
LI ' RK FB
D003 Reactive wastes
F006
F007
F008
F009
F010
Electroplating
sludges and
spent solutions
F011 Heat treating
F012 operations
F014 Tailing pond sediment
F019 Wastewater sludge
KOll
KOI2 Acrylonitrile
K013 production
KG 14
K027 Diisocyanate product
Not Listed
Poor
Poor
Poor
Poor
Poor
Poor
Poor
Poor
Not Listed
Low
High
Low
Low
Low
K060
K087
P013
P021
P027
P029
P030
P031
P032
P033
P052
P055
P063
P064
P069
P074
P098
P099
Coking
Barium cyanide
Calcitna cyanide
3 Chloropropionitrile
Copper cyanide
Cyanide salts
Cyanogen
—
Cyanogen chloride
—
—
Hydrogen cyanide
Isocyanic ester
2 Methylacrylonitrile
Nickel cyanide
Potassium cyanide
Pot, /silver cyanide
Poor
Not Listed
Poor
Poor
Low
Poor
Poor
High
Low
Low
High
Poor
High
High
Low
Poor
Poor
Poor
X X
X
X
X
X X
X X
X X
X X
X
X
X
X
X
X
X
(Continued)
15-9
-------
TABLE 15.2.L (continued)
Waste code
description1*
P104 Silver cyanide
P106 Sodium cyanide
P121 Zinc cyanide
U003 Acetonitrile
0009 Acrylonitrile
U152 Methaerylonitrile
0223 Toluene Biisocyanate
U246
Applicable technology*"
Ranking13
Poor
Poor
Poor
Low
High
Low
Low
Not Listed
LI
X
X
X
X
RK
X
X
X
X
FB
X
X
' X
X
SIC codes:
2869 - Solid waste from ion Low
exchange column,
acrylonitrile production
2869 - Dottoo stream from quench High
column, acrylonitrile
production
2869 - Still bottoms, aniline . High
production
aEPA and SIC waste codes determined to contain cyanides (see Section 2).
General rationale for ranking as "poor", "low", or "high" is based upon
heat of combustion, moisture content, solids content, and several other
key waste characteristics. For a detailed explanation reference should be
made to Reference 3. !
CLI = liquid injection incinerators; RK. = rotary kiln incinerators;
F3 = fluidized bed incinerators.
Source: Adapted from Reference 3.
15-10
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TABLE 15.2.2,
SUMMARY OF WASTES CONTAINING CYANIDES CURRENTLY
INCINERATED OR POTENTIALLY INCINERABLE (1981)
Waste
code or
SIC code
Quantity incinerated
or incinerable Data
Description of waste stream (metric ton/yr) source
Currently incinerated:
D003
P063
P074
P106
U003
U223
Potentially
K011
K012
K013
K014
K027
K048
' K049
K050
K051
K052
K086
2851
2851
34XX
35XX
36XX
37XX
2911
2834
2869
2869
F007,
F009a
3471
3471
R060
F010
K011.K013,
K014a,K027
Non-listed reactive wastes
Hydrogen cyanide
Nickel cyanide
Sodium cyanide
Acetonitrile
Toluene Diisocyanate
incinerable;
Acrylonitrile Stripper Bottoms
Acrylonitrile Bottoms
Crude Acrylonitrile
Acrylonitrile Purification Wastes
TDI Sludge
Petroleum Refining, Dissolved
Air Flotation Wastes
Slot Oil Solids'
Petroleum Industry, Heat
Exchanger Sludge
API Separator Sludge
Leaded Tank Bottoms
Printing Ink Sludges
Paint Prod. Trim Sludge
Paint Prod. Paint Waste
Fab. Metal Prod. Paint Waste
Machine Man. Paint Waste
Electric Eq. Man. Paint Waste
Transportation Eq. Paint Waste
Crude Tank Bottoms
photochemical Wastes
Acrylonitrila Sludges
Acrylonitrile Acid Wastes
Spent Cleaning and
Electroplating Solutions
Spent Electroless Nickel
Plating Solutions
Electroplating Rinse Water
Ammonia Still Lime Sludge
from Coking
Heat Treatment Wastes
Acrylonitrile Bottoms
TDI Residues
12,973.7
230.6
874.1
3.7
9,068.2
178.5
2,700,000
5,900
6,615
47,628
6,284
41,600.
46', 800 "'•
1,700
312,500
1,200
25,200
33,500
11,898
399,237
62,962
117,264 ' '
248,029
1,000
661
1,852 "
33,075
1,990
30.7
17,300
72.0
6.0
3,181
107.9
MITRE
MITRE
MITRE
MITRE
MITRE
MITRE
IEI
IEI
IEI
IEI
IEI
IEI
". ' IE I- •
IEI
IEI
IEI
IEI ,
IEI
III
IEI
IEI
. IEI
IEI
IEI
IEI
IEI
IEI
ICF
ICF
ICF
ICF
ICF
ICF
ICF
aICF lists as combined waste stream.
Sources: Reference 4, 5, and 6.
15-11
-------
* Solids content;
• Chlorine content;
• Viscosity; and
* Corrosivity.
A detailed discussion of each of these parameters is presented in Reference 2.
Relative to the characteristics of cyanide-bearing wastes, it is clear
that, in general, several key factors are regarded as highly restrictive to
incineration. First, as indicated by several officials at commercial
incineration facilities contacted by Alliance in an industry survey, the
potential for formation of highly toxic cyanide pases, such as HCN, -present a
significant restriction on the application of incineration by these commercial
facilities. Such gases would require stringent control and safety
precautions, which would in turn significantly affect overall treatment
costs. Second, many cyanide—bearing wastes, particularly the inorganic
wastes, are aqueous waste streams. Such wastes would require blending with
auxiliary fuel prior to incineration, which could constitute a significant
;- ?' "i I. - '
cost increase. Third, such wastes may often exhibit highly corrosive
properties, which would necessitate the usage of thermal systems which are
resistant to corrosion. While such systems 'are common, they involve higher
costs due to increased air pollution requirements and solid and liquid waste
effluent handling requirements.
15.2.3 Process Costs
Process costs constitute the primary constraint to the usage of
incineration (and other.thermal treatment processes) for management of "•*>
cyanide-bearing wastes. Overall, both capital costs and operating costs are •
high, due to the size of such systems and their requirements. Costs for
incineration systems are detailed more fully in Reference 2,
15-12
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15.2.4 Process Status
Incineration of cyanide-bearing wastes is a common method of disposal of
several by-product streams in the chemical manufacturing industry, most
notably, acrylonitrile manufacturing. Cyanide-bearing wastes arc not commonly
accepted, however, by operators of commercial incinerators, who cite economic
and environmental constraints - in particular those related to the potential
for generation of highly toxic cyanide gases in air emissions - as the
principal deterrent.
15-13
-------
REFERENCES
1, Versar, Inc. Technical Assessment of Treatment Alternatives for Wastes
Containing Metals and/or Cyanides, Draft Final Report. Versar, Inc.,
Springfield, VA. Prepared for U.S. Environmental Protection Agency,
Office .of Solid Waste, Washington, D.C, EPA Contract No. 68-03-3149,
October 31, 1984.
2, Breton, M. et al. Technical Resource Document: Treatment Technologies,
for Solvent-Containing Wastes, Final Report. CCA Technology Division',
Inc., Bedford, MA. Prepared for U.S. Environmental Protection Agency,
Hazardous Waste Engineering Research Laboratory,, Cinneinati, OH. EPA
Contract No. 68-03-3243. August, 1986
3. Advanced Environmental Control Technology Research Center. Research
Planning Task Group Study - Thermal Destruction. EPA-6GO-2-84-Q25.
Prepared for U.S. Environmental Protection Agency, Industrial
Environmental Research Laboratory, Cincinnati, OH. January 1984.
4. Industrial Economics, Inc. Interim Report on Hazardous Waste
Incineration Risk Analysis (Draft). U.S. EPA-WH565. Industrial
Economics, Inc., Cambridge, MA. Prepared for U.S Environmental
Protection Agency, Office of Solid Haste, Washington, D.C. August 1982.
5. ICF Incorporated. RCRA Risk Cost Policy Model - Phase III Report.
Prepared for U.S. Environmental protection Agency, Office of Solid Waste,
Washington', D.C. 1984.' " '•••'• '•• • •
6. MITRE Corporation. A Profile of Existing Hazardous Waste Incineration
Facilities and Manufacturers in the United States. PB-84-157Q72 MITRE
Corporation, McLean, VA. Prepared for U.S. Environmental Protection
Agency, Office of Solid Waste, Washington, D.C. 1984.
7. Herrel, A. GSX Services Corp. Telephone Conversation with M. Kravett,
Alliance Technologies Corp. March 1987.
8. Young, C, • Waste-Tech Services, Inc. Telephone Conversation with M.
Kravett, Alliance Technologies Corp. March 1987.
9. Mullen, D. SCA Chemical Services. Telephone Conversation with M.
Kravett, Alliance Technologies Corp. March 1987.
10. Garcia, G. TWI, Inc. Telephone Conversation with M. Kravett, Alliance .
Technologies Corp- March 1987.
11. Frost, D. Rollins Environmental Services, Inc. Telephone Conversation
with M. Kravett, Alliance Technologies Corp, March 1987.
12. Powers, P.M. How to Dispose of Toxic Substances and Industrial Wastes.
Noyes Cats Corporation, Park Ridge, N.J. 1978.
15-14
-------
13. Anguin, M.T., and S. Anderson. Acrylotvitrile Plant-Air Pollution
Control. Acurex Corporation, Mountain View, CA. Prepared for U.S.
Environmental Protection Agency, Industrial Environmental Research
Laboratory, Research Triangle Park, N.C. EPA-600/2-79-048.
February 1979.
14. Ottinger, R.S., et al. Recommended Methods of Reduction, Neutralization,
Recovery, or Disposal of Hazardous Waste. Volumes 1—16. TRW Systems
Group, Redondo Beach, CA. Prepared for U.S. Environmental Protection
Agency, Office of Research and Development, Washington, B.C. •
EPA-670/2-73-053. August 1973.
15-15
-------
SECTION 16.0
CONSIDERATIONS FOR SYSTEM SELECTION
16.1 GENERAL CONSIDERATIONS
Waste management options consist of four basic alternatives: source
reduction, waste exchange, recycling/reuse, use of a treatment
!
(e.g., precipitation) or disposal processing system or some combination of
these waste handling practices. Recovery, treatment, and disposal may be
performed onsite in new or existing processes or through contract with a
licensed offsite firm which is responsible for the final disposition of the
waste. Selection of the optimal waste management alternative will ultimately
be a function of regulatory compliance and economics, with additional
T ' ' '~'S , ' ;
consideration given to factors such as safety, public and employee'acceptance,
liability, and uncertainties in meeting cost and treatment objectives.
Many of the technologies discussed in previous sections can be utilized
to achieve waste reduction or to meet land disposal ban requirements.
However, practicality will limit applications to waste streams possessing
specific characteristics. Since many processes yield large economies of
scale, waste volume will be a primary determinant in system selection. The
physical and chemical nature of the waste stream and pertinent properties of
its constituents will also determine the applicability of waste treatment
processes. Economical treatment will often involve waste segregation followed
by chemical reduction (e.g., chromium), precipitation (e.g., other metals),-
and/or oxidation (e.g., cyanides) and the use of other technologies in a
system designed to progressively recover/destroy hazardous constituents.
Incremental costs of contaminant removal will increase rapidly as lower
concentrations are attained.
16-1
-------
16.2 WASTE MANAGEMENT PROCESS SELECTION
All generators of hazardous metal/cyanide wastes will be required to
undertake certain basic steps to characterize regulated waste streams and to
identify potential treatment options. Treatment process selection should
involve the following fundamental steps:
1. Characterize Che source, flow, and physical/chemical properties of
the waste.
2. Evaluate the potential for source reduction.
3. Evaluate the potential for waste exchange.
4, Evaluate the potential for reuse or sale of recycled streams and
valuable waste stream constituents; e.g., recovered metals.
5. Identify potential treatment and disposal options based on technical
feasibility of meeting the land disposal restrictions. Give
consideration to waste stream residuals and fugitive emissions to
air.
6. Determine the availability of potential options. This includes the
use of offsite services, access to markets for recovered products
>•, or' waste- exchange-,-? and availability of commercial' equi-pment'"~and .. .
existing onsite systems.
7. Estimate total system cost for various options, including costs of
residual treatment and/or disposal and value of recovered products.
Cost will be a function of Items 1 through 5.
8. Screen candidate management options based on preliminary- cost
estimates.
9. Use mathematical process modeling techniques and laboratory/
pilot-scale testing as needed to determine detailed waste management
system design characteristics and process performance capabilities.
The latter,will define product and residual properties and identify
need for subsequent processing.
10. Perform process trials of recovered products and wastes available
for exchange in their anticipated end use applications.
Alternatively, determine marketability based on stream
characteristics.
11. Generate detailed cost analysis based on modeling and performance
results.
16-2
-------
12. Perform final system selection based on relative cost and other
considerations; e.g., safety, acceptance, liability, and risks
associated with data uncertainties.
Key system selection steps are discussed in more detail in the remainder of
this section,
6.2.1 Waste Characterization
The first step in identifying appropriate waste management alternatives
to land disposal involves characterizing the origin, flow, and quality of
generated wastes. An understanding of the processing or operational practices
which result in generation of the waste forms the basis for evaluating waste
minimization options. Waste flow characteristics include quantity and rate.
Waste quantity has a direct irapact on unit waste management costs due to
economies of scale in processing costs and marketability of recovered
products. Flow can be'Continuous, periodic, or incidental (e.g., spills) and
can be at a relatively constant or variable rate. This will have a direct
impact on stora*g"e requirements and waste management process design; e.g.,
continuous or batch flow.
Waste physical and chemical characteristics are generally the primary
determinant of waste management process selection for significant volume
wastes. Of particular concern is whether the waste is pumpable, inorganic or
organic, and whether it contains recoverable materials or constituents which
may interfere with processing equipment or process performance. Waste
properties such as physical form., degree of 'Corrosivity, reactivity,
compatibility with other wastes and reagents, heating value, viscosity,
concentrations of metal/cyanide chemical constituents, biological and'-chemical
oxygen demand, and solids, oil, grease, total organic, and ash content need*to
be determined to evaluate applicability of certain waste management
processes.' Individual constituent properties such as solubility (affected by
the presence of chelating compounds),- vapor pressure, partition coefficients,
reactivity, reaction products generated with various biological and chemical
(e.g., neutralising, oxidizing, and reducing) reagants, and adscrprion
coefficients are similarly reouired to 'assess "Testability.
16-3
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CheLators and eomplexants will enhance metal solubility, requiring over-
neutralization to alkaline pH to effect metal precipitation. The presence or
absence of buffers will affect neutralization reagent .and •. pK • control system
requirements. Cyanides and chromium will require treatment through oxidation
and reduction, respectively, prior to being combined with other metal-
containing wastes. Finally, wastes with high concentrations or organics may
require subsequent treatment (e»g«» thermal destruction £or sludges,
biological destruction for wastewaters) before wastes can be land disposed.
Finally, variability in waste stream characteristics will necessitate
overly conservative process design and additional process controls, thereby
• increasing costs. Marketability oE recovered products or materials offered
for waste exchange will also be adversely affected by variability in waste
characteristics,
16.2.2 Source Reduction Potential
Source reduction potential is highly site specific, reflecting the
variability of industrial waste-generating processes and product
requirements. Source reduction alternatives which should be investigated
* -inc iude raw material subs titution,.'~--p'roduct; reformulation", ;• process-.-rede-sign-" 'and
waste segregation. The latter may result in additional handling and storage
requirements, while viability of other waste reduction alternatives nay be
more dependent on differential processing costs and impact on product quality.
• Many opportunities exist for firms to achieve waste minimization through
implementation of simple, low-cost methodologies currently proven in
successful programs.^ Lack of available techniques has been less of an
impediment to increased implementation than perception that these methods are
not available.2 Historically, management has favored end-of-pipe treatment
and has bean reluctant to institute waste reduction and reuse -practices. This
reluctance is primarily due to potential for process upsets or adverse impacts
on product quality. Other risks of installing waste reduction methods include
uncertain investment returns and production downtime required for
installation. However, in the wake of increasing waste disposal and liability
costs, source reduction has repeatedly proven to be cost effective, while at
16-4
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Che same time providing for minimal adverse health and environment impact.
Thus, source reduction should be considered a highly desireable waste
management alternative.
16.2.3 jJas_te Exchange Potential
As discussed in Section 5.0, metal bearing wastes have significant
potential for being managed through waste exchange whereas cyanides have
limited potential. Metal bearing wastes will be good candidates for exchange
if: (1) metal concentrations are high; (2) contaminant concentrations are low,
consistent, and at levels which are compatible with user processes;
(3) processing reauirements are minimal; and (4) the waste is available in
sufficient volumes on a regular basis. Waste rinses and solutions
recovered from processes with high purity requirements may be used directly in
processes with lower specifications. An offsite reuse method with high
exchange potential is metal sludge recovery through thermal processing
(Section 12.0). Economics are particularly favorable when these individual
wastes would, .have required-..s,epara.te} .treatment .or costly post-treatment for
organic removal. Finally, waste exchange may prove to be the least cost
management option for firms with wastes that have high recovery potential, but
lack the waste volume or capital to make onsite recovery viable.
Potential for waste exchange is reduced when industries are faced with
liability or confidentiality concerns, and stringent quality requirements.
Additionally, transportation costs are frequently a limiting factor in the
exchange of high volume, low concentration wastes.
16.2.4 Recovery Potential
As part of the waste characterization step, the presence of potentially
valuable metal waste constituents should be determined. Alternatively, in the
case of concentrated acid or alkali solutions, the bulk of the waste may have
recycle potential. Economic benefits can result from recovery of toxic metals
from these materials if the purified solution can then be either reused in
onsite applications or marketed as & saleable product. In the former case,
economic benefits result from decreased consumption of virgin raw materials.
16-5 . .
-------
This must be balanced against possible adverse effects on process equipment or
product quality resulting from buildup or presence of undesirable
contaminants. Market potential is limited by the lower value of available
quantity or demand. Market potential will be enhanced with improved product
purity, availability, quantity, and consistency.
Onsite reuse has several advantages relative to marketing for offsite use
including reduced liability and more favorable economics. Offsite sale is
less profitable due to transportation costs and the reduced purchase price
which offsite users can typically be charged as a result of uncertainties in
product quality. Thus, economics and liability combine with factors such as
concerns about confidentiality to encourage onsite reuse whenever possible.
In practice, recyclinE of metal/cyanide wastes has been limited to
recovery of metals from concentrated solutions, such as plating and etching
baths, thermal recovery of highly concentrated sludges and solids, and removal
from rinses through use of membrane separation and electrolytic recovery
techniques. Cyanide solutions (e.g., plating baths) are sometimes recycled
using metal removal processes but are not more frequently recovered due to the
low purchase price of cyanides. Recycling options have been sutntaarized in
Section 5.0 and discussed in detail in Sections 6.0 through 13,0, These
technologies are summarized in Table 16.2.1 with information provided on
current applications, residuals generated, and availability.
16.2.5 Identifying Potential Treatment and Disposal Options
Following an assessment of the potential for source reduction and
recycling, the generator should evaluate treatment systems which are
technically capable of meeting the necessary degree of hazardous constituent
removal or destruction. Guideline considerations for the investigation of
treatment technologies are summarized in Table 16.2.2 and discussed below.
Waste characterization steps outlined previously define inputs to the
treatment process. Similarly, discharge and residual disposal requirements
(e.g., land disposal restrictions on leachate concentrations) define the
extent to which processing is required. Thus, restrictive waste
characteristics (e.g.", concentration range, flow, interfering compounds) and
technological limitations of candidate treatment processes will define the
field of potential technologies for a specific waste.
16-6
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TABLE 16.2.1. SUMMARY OF RECYCLING TECHNOLOGIES FOR METAL WASTES
Process Applicable waste streams
Evaporation/ Metal plating rinses;
distillation acid pickling liquors
Crystallization h^SO* pickl ing liquors;
HN03/I1F pickling liquors;
solutions.
pickling baths ; aluminum
etching solutions;
h^SO/, anodizing
solutions; rack-stripping
solutions (HF/HNOj).
Electrodialysis Recovery of chromic/
sulfurtc acid etching
solutions-
Recovery of plating rinses
(particularly chromic acid
rinae water) .
Recovery of HNO3/HF
pickling I iquora.
1
\
Stage of development
Well-established for
treating plating
rinses.
20 to 25 systems cur-
rently in operation
(fewer applications
Severs 1 RFIE units in
operation for treat-
Unita for direct
bath only available
from ECO-TEC, Ltd.
Units currently being
sold, but limited
' area of application.
5 in operation.
Severs! in ope rat ion.
Marketed, none in
operation to date.
Performance
Plating solution recovered
for reuse in plating bath.
97-98Z recovery for HoSO^
(80-851 metal removal).
991 HN03 and 501 HF
recovered.
BOX recovery of NaOH.
Cocurrent ayetems not tech-
nically feasible for direct
be used in conjunction with
RFIE units show good results.
Conventional RFIE performs
AFU performs best with high
(30 to 100 g/L).
solution.
45Z copper removal;
301 cine removal.
Works best when copper con-
6 oz/gal usage.
3 H HF/HN03 recorded.
Impurities uill be
concentrated , therefore,
17 °h '
Ferrous sulfste heptnhydeate
sold).
Metal fluoride crystals (can
recover additional HF by
thermal decomposition).
Aluminum hydroxide crystals
Cocurrent process generates
spent re gene rant , which ia also
Recovered metals which can be
reused , t reated , disposed , or
Metals which can be treated,
reuse.
Chromic acid can be returned to
be reused.
2 M KOH Soln which can be
Coat
1 G<
plating solutions from
Cost-effective if treat-
waste.
RFIE and APU are
cost-effect ive .
Cost-effective for
specific applications
( chromic /BU If ate acid
etchants).
Low capital investment;
specific application
Cost-effective for large
meat step for this ED appl ication.
(continued)
-------
-------
TABLE 16.2.1 (continued)
Applicable
streams Stage of development
Performance
Residuals generated
Cost
Reverse osmosis Plating rinses.
Donnan dialysis/
coupled
transport
Plating rinses; poten-
tially applicable to
acid baths.
Solvent
extraction
Thermal
decomposition
HN03/HF pickling
liquors.
Acid wastes.
Corrosive waste mem-
branes marketed by
four companies.
RD module Bysterna
applicable to corro-
flivefl available from
two companies*
Donnan dialysis only
lab-scale tested.
Coupled transport
lab and field tested.
Coupled transport
system ie currently
being marketed.
Commercial-scale
systems installed for
development purposes
in Europe and Japan.
No commercial-scale
inatallationa in U.S.
Well-established for
recovering spent
pickle liquors gen-
erated by steel
industry. Pilot-
scale stage for
organic wastes.
90T conversion achieved
Recovered plating solution
Cost-effective for
with cyanide plating rinses, returned to plating bath (after limited applications.
Data not available for
Donnan dialysis (further
teat ing requi red),
Coupled transport has dem-
ons t rated 99It recovery of
chromate from plating rinaea.
Other plating rinses should
being concentrated by an
evaporator)- Rinsewater
reused.
Data not available for Donnan
dialysis.
For chromate plating rinse
applications, sodium chromate
ie generated; can be used else-
where in plant or subjected
be applicablep but not fully to ion exchange to recover
teated. chromic acid for recycLe to
plating solution.
951 recovery of HNOj;
70X recovery of HF.
991 regeneration efficiency
for pickling liquora.
Metal sludge (952 Iron can be
recovered by thermal
decomposition).
98-991 purity iron oxide which
can be reused, traded, or
marketed•
Development of a more
chemically resistant
membrane would make it
very cost-effective for
a wider area of
application.
No cost data available
for Donnan dialyaia.
Average capital cost
for plating shop LQ
420,000. Can be coat-
effect ive for specific
applications.
Mot available.
Expensive capital
investment. Only coat-
effect ive for Large
quantity waste acid
generators.
-------
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TABLE 16.2.2. GUIDELINE CONSIDERATIONS FOR THE INVESTIGATION OF WASTE
TREATMENT, RECOVERY, AND DISPOSAL TECHNOLOGIES
A. Qbj ec tives of treatment:
- Primary function (pretreatment, treatment, mutual neutralization,
residuals treatment)
Primary mechanisms (destruction, removal, conversion, separation)
Recover waste for reuse
Recovery of specific 'chemicals, group of chemicals (acids, alkalis,
metals, solvents, other organics)
Polishing for effluent discharge (NPDES, PGTOj
Immobilization or encapsulation to reduce migration (inorganic sludge)
- Overall volume reduction of waste
- Selective concentration of constituents (acids, alkalis, metals,
solvents, other organics)
- Detoxification of hazardous constituents
B. Waste applicability and restrictive waste characteristics:
- Acceptable concentration range of primary and restrictive waste
constituents
Acceptable range in flow parameters
Chemical and physical interferences (compatibility with reagents)
C. Process operation and design:
- Batch versus continuous process design
Fixed versus mobile process design
Equipment design and process__cpntrol complexity (pH, flow, Reddx
potential, conductivity, temperature, pressure, level indicators)
variability in system designs and applicability
Spatial requirements or restrictions
(continued)
16-9
-------
-------
. TABLE 16.2.2 (continued)
Estimated operation time {equipment down-time)
Feed mechanisms (wastes and reagents; solids, liquids, sludges,
slurries)
Specific operating temperature, flow, and pressure
Sensitivity to fluctuations in feed characteristics
Residuals removal mechanisms
Reagent selection and requirements
Ancillary equipment requirements (tanks, pumps, piping, heat transfer
equipment)
~ Utility requirements (electricity, fuel and cooling, process and
make—up water)
D. Reactions and theoretical considerations:
- Waste/reagent reaction (neutralization, destruction, conversion,
oxidation, reduction)
... - Competition, or suppressive reactions (consplexants, chelators, buffers)
Enhancing conditions (specify chemicals)
Fluid mechanics limitations (mass and heat transfer)
— Reaction kinetics (temperature and concentration effects)
- Reactions thermodynamics (endothernjic/exothermic/catalytie)
E. Processefficiency:
Anticipated overall process efficiency
- Sensitivity of process efficiency to:
o feed concentration fluctuations
o reagent concentration fluctuations
o process temperature fluctuations
o toxic constituent concentrations (biosystems)
o physical form of the waste
o other waste characteristics
(continued)
16-10
-------
.TABLE 16.2.2 (continued)
F. Emissions and residuals management:
- Extent of fugitive and process emissions and potential sources
(processing equipment, storage, handling)
Ability (and frequency) of equipment to be "enclosed"
Availability of emissions and residuals data/risk calculations
Products of incomplete reaction
- Relationship of process efficiency to emissions or residuals generation
- Air pollution control device requirements
Process residuals (fugitive/residual reagents, recovered products,
filter cakes, sludges, incinerator scrubber water and ash)
- Residual constituent concentrations and leachability
Delisting potential
G. Safety considerations:
.., . ,...., f , 4. , •,
Safety of storing and handling reactive or corrosive wastes, reagents,
products and residuals
- Special materials of construction for storage and process equipment
- Frequency and Deed for use, of personnel protection equipment
Requirements for extensive operator training
f Hazardous emissions (e.g., HCN) of wastes or reagents
Minimization of operator contact with wastes or reagents
Frequency of maintenance of equipment containing hazardous materials
High operating temperatures and pressures
- Difficult to control temperatures
Resistance to flows or residuals buildup
Dangerously reactive wastes/reagents
- Dangerously volatile wastes/reagents
16-11
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Estimating the appropriateness of waste treatment options requires an
in-depth understanding of theoretical considerations. All unit operations
have inherent limitations based on technical constraints (e.g., mass transfer
limitations, reaction kinetics) aod economic feasibility (e.g., restricted
range of temperature, pressure, and other operating conditions; limits on
materials of construction). Estimation of system performance capabilities
will involve a systematic analysis of several interdependant considerations:
(1) expected equilibrium products for chemical, biological, thermal, or
physical processes; (2) reaction kinetics; (3) heat transfer and mass
transport phenomena; and (4) process control requirements.
A key consideration in the choice of chemical treatment systems for
metal/cyanide wastewaters is reagent selection (Section 10.0 for metals and
Section 14.0 for cyanides). Reagents may require special handling
characteristics or form hazardous or difficult to manage reaction products.
Potential reagents and their associated advantages and disadvantage with
respect to costs, handling, processing and- sludge generation, are summarized
in Table 16.2.3.
Residual characteristics will have a significant impact on ultimate
"reagent selection since treatment or'disposal of thes'eTmaterials- constitutes a"
large percentage of total waste management costs. Depending on the reagent
selected and original waste characteristics, sludges will have different
settling, dewatering, and compactability characteristics, as well as varying
tendencies for their heavy metals to resolubilize. For wastewaters, the
presence of toxic organics will also significantly add to post-treatment
costs. Costs will increase with organic concentration and required removal
efficiency and decrease with reactivity, volatility, adsorbability and
. . 4
biodegradability.
Ultimately, the selection of a specific treatment system from a list of
technologically feasible alternatives will depend on cost, availability, and
site specific factors. These considerations are discussed below.
16.2.6 Availability of Potential Management Options
The availability of each component of a waste management system may
restrict its overall applicability. Existing available capacity of onsite
treatment processes (e.g., wastewater treatment systems), ancillary equipment,
16-12
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TABLE 16.2.3. METALS/CYANIDE TREATMENT AND CHARACTERIZATION
Convnnn form nnd
Commercial Appro* im.itc '
I'c.Tf.^nt Clicmicnl nnmc atrcngh coat/ton(t)
Met a In P re c i |' it.it ion
Hi nh
-------
TABLE 16.2.3 (continued)
Reagent
Metals Reduction
Sulfur Dioxide
Sodium Metabisulfite
Ferroua Sulfate
Sodium Dorohyilride
Cyan ide Ox irlat ion
Qilorine
Sodium llypochlorite
Chemical name
Sul fur dioxide
Sad ium
py rosul f ite
heptahydrate
Sodium
borohydride
Chlorine
Sodium hypochlo-
rite pentohydrate
Common form and
commerc ial Approximate
etrengh cost/ton ( fc)
Gas 230
99. 9X S02
Flake . 64
70-72Z
Powder 6,000;
97Z NaBII4
Cos 195
99.51 C12
Solution 304
29Z NaOCl
Handling properties
from cylinders potential
explosion hazard.
or bulk liquid applica-
tions .
Good- avail able in flake
or solution form.
Good-Bui table for dry
or liquid feeds.
from cylinders or bulk
potent ial explosion
hazard.
Good-available In flake
or solution form.
Fast-requires IS to
30 minutes for complete
reaction.
chrome redact ion.
Faet-a imilar to other
giea.
Fast-requires 5 to
60 minutes depending
process stable com-
pleMea.
Foflt-slmilar to chlo-
rine.
Sludge generation
Lou-al 1 endproducta
are aolub le.
Lou- all e nd p roduc t a
High-will result in
times higher than
busulfite reduction.
Low-high dens ity high
metal content sludge.
Low—no inly chlorine
byproducts.
Low— u ill no t re ad i 1 y
-------
labor, physical space, and utilities will have a significant impact on the
economic viability of a treatment system. Purchased equipment must be
available in sizes and processing capabilities which meet the specific needs
of the facility. Offsite disposal, recovery, and treatment facilities, and
companies using eKchahged materials or purchasing saleable products, must be
located within a reasonable distance of the waste generator. In addition,
they must have available capacity for the waste type and volume generated,
Finally, time constraints may eliminate certain treatment processes from
consideration as a result of anticipated delays in procurement, permitting,
installation, or start-up.
In general, precipitation and chemical oxidation systems are widely
applied and readily available- However, several recovery systems (e.g.,
Devoe-Holbein extraction ,for metals, 1HCO S0~/air process for cyanide
oxidation) and post-treatment systems for organic wastes (e.g., chemical
stabilization) have only recently been applied in metal /cyanide waste
treatment. Availability and uncertainty in expected cost-effectiveness will
play a significant role in the decision to implement these technologies.
' • '.'•'' *!"*• - ; .-- '- .. ' . .—-.-;- . ...*.• $ c'-f i ... . - ••-- •-•:•..• •• ~ts •-:-•- - -
^anagg?"ent System Cost Estimation
The relative economic viability of waste management systems will be the
primary determinant of system selection for processes which are capable of •
achieving comparable performance. Economic viability must be evaluated on the
basis of total system costs. This includes operating and capital purchase-
costs as well as the availability of onsite equipment, labor and utilities,
net value of recovered products, and residuals disposal costs. High capital '
equipment expenditures and financing constraints are frequently a limiting
factor in system selection, particularly for firms, with marginal profitability
or high debt/equity ratios, and for processes which hsve higher uncertainties
of success. • '
Costs for a given management system will be highly dependent on waste
physical, chemical, and flow characteristics. Thus, real costs are very
sice-specific and limit the usefulness of generalizations. The reader is
referred to tne sections on specific tecnnoiogies ^Sections 6.0 -through 15.0}
for data on costs and tneir variability uitn respect co flow and waste
16-15
-------
characteristics. Costing methodologies have also been described in the
c _ T O
literature and are available in software packages for select
processes. Major cost centers which should be considered are summarized in
Table 16.2,4.
16,2.8 Modeling System Performance and _Pilot~Scale Testing
Following a preliminary cost evaluation, which will enable the generator
to narrow the field of candidate waste management options, steps must be taken
to finalize the selection process. Initially, these could involve the
application of theoretical models to predict design and operating
requirements. However, models generally sacrifice accuracy for convenience
and are often not sufficiently accurate to describe complex waste streams.
Laboratory or pilot-scale data are often needed as model inputs and, in most
cases, are ultimately required to confirm predicted performance prior to final
system selection.
Nevertheless, in many cases, modeling can minimize costly laboratory
testing. Models are particularly useful in assessing relative
'cost-effectiveness1 -with respect to changes .in process variables and the
incremental costs of achieving increasingly stringent treatment concentration
levels. Thus, process evaluations should begin with the formulation of a
model which incorporates the conceptual process train and the primary
variables which affect process performance and design. These variables can
then be assigned a range of values to reflect the previously defined source
conditions. The results of computer simulations or paper studies can then be
used to project anticipated full~scale results and define areas for
bench-scale testing.
Bench scale studies must be designed to provide maximum accuracy, and to
facilitate subsequent scale-up. Equipment design parameters and operating
conditions must preserve geometric, kinematic, dynamic, and thermal
similarity. When possible, input parameters should therefore be arranged in
the form of dimensionless variables (e.g., Reynolds number). Chemical
similarity should also be maintained by using representative samples from the
waste generating process. Factorial experiment design and response surface
methodology techniques can be applied in bench and process trials to ensure
' 16-16
-------
TABLE 16.2.4. MAJOR, COST CENTERS FOR WASTE MANAGEMENT ALTERNATIVES
A. Credits: - . •' .•.•'•'.
- Material/energy recovery resulting in decreased consumption of •. • .
purchased raw materials
- Sales of waste produces - ' '•
B. Capital costs;*
•" Processing equipment (reagent addition reaction vessel, recovery
apparatus, sludge end other residual handling equipment)
Ancillary equipment (storage tanks, pumps, piping)
Pollution control equipment
Vehicles ' ' .
Buildings, lend
- Site preparation, installation, start—up
C. Operating and maintenance costs:
Overhead, .operating, and maintenance labor
~ Maintenance materials
- -( Utilities .(electricity, fuel, , Duster-): , , , _m , . ..•<•.
~ Reagent materials
Disposal, offsite recovery, and waste brokering fees
Transportation
- Taxes, insurance, regulatory compliance, and administration
3, Indirect costs and benefits:
- " Inpacts on other facility operations; e.g., changes in product quality
as £ result of source reduction or use of recycled materials
Use of processing equipment for management of other wastes
aAnnuaI costs derived by using a capital recovery factor:
Where: i = interest rate and n = life of the investment. A CRF of
0.177 was used to prepare treatment cost estimates in this
document. This corresponds to an annual interest rate of
12 percent and an equipment life of 10 years.
16-1?
-------
that optimal performance results are obtained in the most cost-effective
9
manner. Quality control procedures should be implemented to ensure
consistency and accuracy of results. Finally, precautions should be taken to •
ensure that measurement and control equipment employed in the process
evaluation is sufficiently sensitive and versatile to assess the effect o£
process and feed variations on overall treatment effectiveness.
The final step in the technical approach may involve design,
installation, and testing of treatment systems which have been identified as
the most promising candidates for specific applications. Standard chemical
engineering techniques should be utilized to scale-up control equipment from
bench scale results, ' Integration of a treatment technology into an
industrial process will require development of energy and material balances
and a detailed economic analysis. Potential process variations and upsets,
impact on existing operations, ease of operation and control, safety factors,
and other considerations will be incorporated into the final design. These
factors will be evaluated on a case by case basis taking into consideration
input data uncertainties, institutional and regulatory constraints, and the
probability and consequences of failure to meet control objectives.
. .-. : , Many suppliers of-treatment and recovery1 equipment1 use models to ;optiinize
design and operations parameters and to scale-up processes. Some equipment
manufacturers are also able to provide experimental equipment and models to
establish process parameters and cost.
16-18
-------
SECTION 16 REFERENCES
1. Allen, C.C., and B. L. Blaney.' Research Triangle Institute, Techniaues
for Treating Hazardous Waste to Remove Volatile Organics Constituents,
Performed for U.S. EPA HWERL, EPA-600/2-85- 127. March 1985.
2. Committee on Institutional Considerations in Reducing the Generation of
Hazardous Industrial Wastes. Environmental Studies Board, National1
Research Council. Reducing Hazardous Waste Generation: An Evaluation
and a Call for Action. National Academy Press, Washington, D.C. 1985.
3, GCA Technology Division, Inc. Industrial Waste Management Alternatives
Assessment for the State of Illinois. Volume IV: Industrial Waste
Management Alternatives and Their Associated Technologies/Processes.
Final Report prepared for the Illinois Environmental Protection Agency,
Division of Land Pollution Control. GCA-TR-8Q-8Q-G. February 1981.
<3
4. Breton, M. et al. Alliance Technologies Corporation. Technical Resource
Document: Treatment Technologies for Solvent Containing Wastes..
Prepared for U.S. EPA HWERL under Contract No. 68-03-3243, August 1986.
5. Peters, M. S., and K. D. Tinanerhaus. Plant Design and Economics for
Chemical Engineers. 3rd Edition. McGraw Hill Book Company, New York,
NY. 1980.
6. U.S. EPA Design Manual: Dewatering Municipal Kasteuater Sludges. U.S.
EP_A Municipal Environmental Research Laboratory,,,. ,Cincinr;ar.i.,«.,,OH.,. ,. ..
•"-• ' EPA-625yi-8Z-01^r .October 1982. " "
7. MITRE Corp. Manual of Practice for Wastewater Neutralization and
Precipitation. • EPA-6QO/2-81-148, August 1981.
8. Cunningham, V. L. et al. Smith, Kline & French Laboratories.
Environmental Cost Analysis System. 1986.
9. Box, G.E.P., W. G. Hunter, and J. S. Hunter. Statistics for
Experimenters - An Introduction to Design Data Analysis and Model
Building. John Wiley and Sons, N.Y., N.Y. 1978.
10. Johnson and Thring. Pilot Plants, Models, and Scale-Op Methods in
Chemical Engineering. McGraw Hill Book Company. N.Y., N.Y. 1980.
16-19
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