Note: Descriptions are shown in the official language in which they were submitted.
~ 11288
BACRGROUND OF THE INVENTION
1. Field Of The Inventi~on
This invention relates to novel processes for
removing sulfur o~ides from gas mixtures containing same
as in the removal o~ sulfur oxides from combustion waste
yases, or stack gases of electric power plants, pyrites
roasting processes, smelters, sulfuric acid plants, off-
gases from incinerators, and off-gases from other chemical
procedures. More generally, the invention relates to the
removal of heat stable salts containing heat stable,
divalent, sulfur oxyanions from systems in which they
accumulate.
2. Description Of The Prior Art
The United States and the world at large are
currently faced with conflicting crises of the shortage
of energy or clean-burning fuels on the one hand and e~er-
increasing environmental pollution on the other. ~he energy
crisis could be largely eliminated if it were possible to
burn the abundance of high sulfur coal or other high
sulfur uels available in this country. Many attempts have
been made to develop power plant stac~ gas clean-up processes
so that this could be done but with little technical and
economic success.
-Many proposed processes react the sulfur dioxide
~-(52) with limestone or other inorganics. Because o~ the
inefficiencies of these reactions, the reagents are used
in greater than stoichiometric amounts, and greater than
stoichiometric quantities of waste solids or liquids are
produced, presenting an additional disposal and secondary
pollution probIem. In addition, the reaction ineEficiencies
11288
S~
permit large quantities of the So2 to go out the stack
anyway.
Attempts have been made to absorb the So2 in
regenerable type solvents (see~ for example, U.S.P.
1,908,731). These processes, too, have been plagued with
inefficiencies, typified by poor S02 pick-up, extremely
high stripping steam requirements, and side reactions such
as S~2 oxidation and disproportionation. The latter lead
to non-volatile or heat stable sulfur oxygen anions, tying
up the solvent and diminishing its capacity so that a large
purge stream must be taken for discard or chemical reclaim-
ing.
Among the solutions proposed for this problem are
those which involve absorbing and/or reacting the sulfur
oxides with inorganic reagents, e.g., sodium carbonate
sodium hydroxide, ammonium hydroxide, aqueous ammonia,
other alkali metal or alkaline earth metal hydroxides or
carbonates and the like, in solution, slurry or powder
form to yield the corresponding sulfate and sulfite salts.
In many of these processes, the absorbing
solutions are regenerated by heating, in a separate vessel,
thus liberating concentrated S02. This desorption step
does not, however, remove any sulfate, thiosulfates or
polythionates that result from absorption of th~ sulfur
trioxide and thermal disproportionation of sulfite and
bisulfite and which eventually build up in the system.
In many of these previously proposed solutions,
the reagent cannot be readily regenerated without the
~ 65~ 2~8
expenditure of considerable amounts of energy or consider-
able amounts of other reagents. In those instances where
a regenerated absorbent can be usea, the sulfate concentra-
tion in the absorbent buildsup both by absorption of
sulfur trioxide or sulfuric acid mist which might be and
usually are present in the stack gas and by oxidation of
dissolved sulfur dioxide by the reaction of oxygen which
is also sometimes present in the stack gas. A further
source of the build up of sulfates or other sulfux
oxyanions of heat stable salts is by disproportionation of
dissolved sulfites and bisulfites. Such heat stable salts
=
include, in addition to the sulfates, SO4 ; the thiosul-
fates, S203 ; the dithionates, S~O6 ; the trithionates,
S36 ; other higher polythionates, SxO6 , and other
divalent sulfur oxyanion-containing heat stable salts. The
sulfates usually can be removed essentially quantitatively
through the use of an alkali metal hydroxide equival~nt
to twice the molar concentration of the sulfate resulting
in substantial quantitative precipitation of the sulfate as
the di-alkali metal salt without precipitation of sulfite
or bisulfite ions. However, the other divalent sulfur
oxyanions of strong acids such as the thiosulfates, dithio-
na~es and higher polythionates also build up in the system
and cannot be quantitatively removed by means of alkali
metal hydroxide precipitation. ~urthermore, the presence
of such other divalent sulfur oxyanions of heat stable salts
ac~ively interfere with -the quantitative removal of the
sulfates.
; In some instances, as in U. S. Patent No. 3,503,185,
the combustion waste gas was prewashbd to remove sulates
11288
which were then purged from the system. Such prewashes
were not capable of removing all sulfur trioxide as sulfate
and, of course, would not re~ove sulfates formed in other
parts of the system. This patent, furthermore, does not
disclose any means for eliminating the thiosulfates,
dithionates and higher polythionates. U. S. Patent No.
3,790,660 is similar in showing a water prewash to remove
sulfur trioxide and fly ash. It specifies a sulfate purge
stream to remove the sulfate; unfortunately, a considerable
amount of the alkali metal sulfite and bisulfite also
accompany the sulfate. This requires a considerable addition
of alkali metal hydroxide to make up for the loss. Further-
more, there is no system disclosed for removing the thio-
sulfates, dithionates or other polythionates except by
purging them with the sulfate in a waste stream. The waste
stream itself is relatively dilute and poses a problem in
disposing of the waste stream which is difficult and
expensive to handle.
There are prior processes which utilize H2S, itself
a noxious gas, to react with sulfur oxides which are
dissolved in solvents, such as alkali metal bisulfites,
ammonium bisulfite, aqueous ammonia or ammonium sulfite.
These prior processes are disclosed in U. S. Patents Nos.
3,561,925; 3,598,529; 3,719,742; 3,833,710; 3,839,549, and
3,883,638. All but the last of these patents fail to
specifically address the problem of removing sulfates and
other heat stable salts which build up or accumulate during
removal of sulfur dioxide. Furthermore, H2S in some cases
is not readily available and can be difficult to store and
handle and can itself possibIy lead to pollution problems.
11288
The use of alkanolamines, such as trialkanol-
amines, has been found to be a highly efficient way o
absorbing sulfur dioxide from waste gases in a cycle in
which the alkanolamine solvent contacts the waste gas to
absorb the sulfur oxides and is thereafter stripped by heat
to release the sulfur dioxide as a gas whereupon it is
collec~ed for safe disposal. The stripped alkanolamine
is then recycled back to the absorber for ~urther contact
with incoming waste gases and further absorption of sulfur
oxide. This type of system is disclosed in U. S. Patent
Nos. 3,620,674 and 3,904,735. Heat stable salts, such as
those mentioned hereinabove, accumulate in the recycling
absorbent to a troublesome extent and must be removed in
order to maintain the absorbing capability of the absorbent.
The latter patent does disclose a sulfate purge cycle in
which a portion of the lean absorbent is treated with
potassium hydroxide or potassium carbonate to precipitate
out ~he sulfate as potassium sulfate. While this type of
purge system is quite effective in removing sulfates, it
is severely limited in removing other heat stable salts
or their divalent sulfur oxyanions, which also seem to
interfere, however, with the sulate removal. Furthermore,
large amounts of wet sulfates are produced and create a
severe disposal problem. There does not appear to be any
provisions made in U. S. Patent No. 3,620,674 for removing
the heat stable salts and/or their sulfur oxyanions from
the absorbent which gradually but inevitably loses
effectiveness because of the accumulation of heat stable
salts therbin.
Anion exchange resins have bean used in the past
~ 7~ 11288
to separate sulfur dioxide from waste gas mixtures. An
example of prior art of this type is United States Patent
No. 3,330,621 which utilizes a mass of solid pyridine group-
containing particles to contact the sulfur dioxide-contain-
ing gas to bind the sulfur dioxide as sulfite groups to the
pyridine groups~ Thereafter, oxygen is added to oxidize
the sulfite groups on the pyridine groups to form sulfate
groups. Then, the sulfate groups on the pyridine groups
are treated with ammonia to form ammonium sulfate which is
then recovered and the pyridine group-containing particles
are recycled for re-contact with the waste gases. This
type of prior art process involves the use of extremely
high quantities of anion exchange resin and excessively
large quantities of ammonia and/or other reagents and
presents a disposal problem for the large quantities o
ammonium sulfate which are produced because the total
quantity of sulfur dioxide in the waste gas is converted
via the pyridine group-containing particles into ammonium
sulfate.
Anion exchange resins have also been used to treat
the total amount of a recycling absorbent, such as sodium
hydroxide or ammonium bisul~ite. In U. S. Patent
3,896,214, the sulfur dioxide and sulfur trioxide in the
waste gases are washed with sodium hydroxide to convert
substantially all the sulfur dioxide and sulfur trioxide
content thereof into sodium bisulfite and/or sodium
sulfite and sodium sulfate which are then contacted with
a hydroxyl-containing weak base or strong base anion
exchange~resin to substitute the hydroxyl groups on the
resin with the bisulfite, sulfite and sulfate anions
~ 112~8
thereby regenerating the sodium hydroxide. The resulting
resin sulfate, sulfite and/or bisulfit:e i9 regenerated by
treatment with aqueous lime hydrate to form calcium sulfate
and calcium sulfite and/or calcium bisulfite and to
substitute hydroxyl anions on the res:in. The calcium salts
are removed as a sludge by dewatering. In U. S. Patent No.
3,833,710, aqueous ammonium sulfite is used as an absorbent
and is converted to aqueous ammonium bisulfite after
picking up the sulfur dioxide in the waste gas. The
aqueous ammonium bisulfite solution is contacted with a
weak base anion exchange resin in the hydroxyl form to
convert the resin to the bisulfite form and regenerate the
ammonium sulfite absorbent solution. Both this and U. S.
Patent No. 3,896,214 are based on the removal from the
waste gases of the total amount of the SO2 con~ent as well
as the SO3 content by utilizing ion exchangeO This requires
the utilization of extremely large amounts of anion
exchLange resins which are expensive and also requires the
use of extremely large amounts of reagents to regenerate
the anion exchange resin which is not only expensive but
presents a considerable waste disposal problem for liquid
wastes that are relatively quite dilute when consideration
is given to the need for ~ashing the resin after each
liquid pass during regeneration.
U. S. Patent No. 2,713,077 discloses the use of
strong base anion exchange resins to remove carbonyl
sulfides from hyrocarbon fluids, such as hydrocarbon gases,
produced by the thermal or catalytic cracking of
petroleum oils or by the reaction of steam with coke or
hydrocarbons. U. S. Patent No. 3,297,401 removes arsenic
11288
and iron contamination from phosphoric acid preparations
wi~h a weak base liquid anion exchange resin. In each
of these patents the spent anion exch.ange resin can be
regenerated with sodium hydroxide. Neither patent relates
to the removal of sulfur dioxide and heat stable salts
from waste gases containing them or their ingredientsO
SUMMARY OF THE INVENTION
The present invention provldes impr~vements
in processes designed ~or the selective remo~al of sulfur
dioxide with respect to carbon dioxide from a gas mi~ture
containing same wherein
a) the gas mixture is contac~ed with an aqueous
absorbent solution containlng as an absorbent
alkanolamine, alkali metal hydroxides,
ammonium hydroxide and/or sulfites thereof
to remove sulfur dioxide from said gas
mixture and form a rich aqueous absorbent
solution enriched in sulfur dioxide removed
from the gas mi~ture,
b) the rich aqueous absorbent solution is moved ?
to a stripping zone where sulfur dioxide is
removed to form a lean aqueous a~sorbent
solution depleted in sulfur dioxide content,
: and
~) the lean aqueous absorbent solution is
recycled and re contacted wit~ said gas
mi~tureO
Processes o~ this type to whîch the improvements
of this invention can be adapted are known in the
art and are illustrated by U. S. Patents Nos,
3,904~735 and 3,620,674 w~ich described processes
j ~,)
.... ~
7~i
11288
utilizing alkanolamines or sulfites thereof as absorbents;
3,790~660 and 3,719,742 which respect:ively use sodium
sulfite and potassium sulfite; 3,833,710 which uses
ammonium sulfite; 3,503,185, 3,561,92!;, 3,839,549 and
3,883,633 which use aqueous ammonia or ammonium hydroxide;
and 3,896,214 which uses sodium hydroxide as absorbent.
It is important to use those operating conditions
that restrain SO2 oxidation, SO~ disproportionation, and
solvent degradation. These include, for example, limiting
the temperature of absorption and stripping to 125QC or
less and maintaining in the solvent at least one mol of water
for every mol of S02 absorbed. It i9 recognized, too,
that even under such constraints there could be some SO~
oxidation to produce the heat stable sulfate ion. In
addition, it is realized that some SO3 exists in the stack
gas as such, and would be simultaneously absorbed along
with the SO2 to give additional sulfate ion. It has been
previously proposed to remove the sulfate ion by potassium
ion precipitation (see U. S. P. 3,904,735). Whereas this
is an excellent step in the situation wherein the major
heat stable salt is the sulfate ~as it would certainly
be if the procass were used to treat sulfuric acid tail gas
containing large quantities o~ SO3), in many other
circumstances this is not the case. The effective choice
of solvent and operating conditions reduces oxidation to
~ery low levels and, if there is little SO3 in the stack
gas, sulfate is no longer the major heat stable anion.
Instead, the products of SO2 disproportionation and other
side reactions make up the bulk of the heat stable anions
and these include, in addition to sulfate, thiosulfate,
--10--
i7~i
11288
dithionate, trithionate and other species. The previously
proposed alkali metal precipitation techniques are not
sufficiently effective against khese other heat stable salts
and a need existed for improvement.
The improvement of this invention relates to the
removal of divalent sulfur oxyanions of heat stable salts
which accumulate in the aqueous absorbent solution, and
comprises the step of contacting a small portion of the
lean aqueous absorbent solution as a slip or purge stream
with an anion exchange resin having hydroxyl anions
displaceable by the divalent sulfur oxyanions of the heat
stable salts to remove same from the lean aqueous absorbent
solution whereby the hydroxyl anions of the resin are
replaced by the divalent sulfur oxyanions o the heat stable
sal~s. The hea~ stable salts (collectively designa~ed HSS)
exist in the absorbent primarily in the dinegative anionic
=
forms SO4 , S2O3 , S2O6 , etc., and are collectively
designated as HSS=. The present invention is based upon the
ability of the hydroxyl anions of an anion exchange resin
containing them to be displaced by HSS in the lean aqueous
absorbent solution which thus become bonded to the resin.
The present invention provides the further advantage that
the anion exchange resin is used to treat the lean aqueous
absorbent to avoid as much as possible the useless
consumption of ion exchange capacity. This permits the more
- selective removal of heat stable salt anions with as little
waste as possible of ion exchange capacity and regenerant
costs on needless aqueous sulfur dioxide or bisulfite ion
removal~
7~
11288
These discoveries are used by the present
invention to treat a slip stream, or purge stream, of the
aqueous absorbent, such as aqueous triethanolamine sulfite
or bisulfite, contaminated with heat stable salts, with an
anion exchange resin in the hydroxyl form after the aqueous
absorbent has been used to treat gases containing sulfur
dioxide which is absorbed thereby and after stripp:ing the
absorbed SO2. The absorbent normally is regenerated by
heating it in an evaporator or still or other suitable
stripping apparatus and is recycled for further contact
with the yases containing sulfur dioxide. In this way,
the gases are depleted in sulur dioxide content and sulfur
dioxide is concentrated for use or disposal. However,
heat stable salts and/or the divalent sulfur oxyanions
thereof form in the system, or enter it with the waste
gases, and, unless removed, they can build up to the point
where the absorbent no longer functions efficiently and
ultimately becomes incapable of absorbing sulfux diox:ide.
Ion exchange has been defined as the reversible
exchange of ions between a solid and a liquid in which there
is no substantial change in the structure of the solid.
~nion exchange resins are high molecular weight polybases
containing large numkers of polar exchange groups that are
chemically bound to a three-dimensional hydrocarbon net-
work. Anion exchangers usually contain primary, secondary
and/or tertiary amine groups, and/or quaternary ammonium
groups. However, sulfonium anion exchangers have been made
and phosphonium anion exchangers are feasible. The mode of
operation of anion exchange resins is the same as a
solution phase reaction, with the exception that one of the
-12-
V~ 7~
l1288
ionic species is permanently fi~ed to the resin be.ad it-
self, and the reaction is shifted back and fourth by strong
concentration differences of the various solutions used.
The three dimensional hydrocarbon network to which
the polar exchange groups are bound usually is of a
polymeric nature. Polystyrene crossli.nked with divinyl-
benzene to provide the needed or desired dimensional
stability is most frequently used although other polymeric
forms t such as the acrylics and other vinyl polymers, are
used or are available, for example, phenol-formaldehyde
resins have been widely used. The polar groups are bonded
to the resin or polymer, usually, by a procedure involving
chloromethylation and amination. The physical form of the
anion exchangers is preferably a bead form as is obtained
by suspension or dispersion polymerization of styrene and
divinylbenzene, although other particle forms are used,
for example, granulated phenol-formaldehyde resins.
Commercial anion exchange resins are available in
two major types, macroreticular and "gel" type. The macro-
reticular resins are small, tough, rigid plastic beadshaving large discrete pores in the range of 1300 A. Because
o~ their large uniform pore structure, these resins are
useful for the absorption and elution of high molecular
weight ions and have good hydraulic characteristics.
Furthermore, because of their tough structure ~they have a
higher degree of crosslinking), they are less susceptible
to physical atbition and have a long operating life. They
do, however, have somewhat lower capacity than the "gel"
resins. The latter do not contain any true porosity (and
usually have a much lower degree of crosslinking3 and the
11288-C
ions to be exchanged must "diffuse" throug~ the gel
structure, thus limitîng the size oE the ions that can
be handled. The advantage to this type o-f resin is i~s
higher exchange capa~ity and somewhat lower capltal cost.
Both of these two major types are broken down
into two sub-groups: The strongly basîc anion exchangers
with quaternary ammonium functionalit:y and the weakly basic
anion eschangers with polyamine functionality. The strong
base resins ofer much be~ter removal o anions in the
e~haustion step w~ereas the weak base resins ofer greater
ease o~ regeneration.
Suitable anion exchange resins for u~e in this
i~vention include the strong base and weak base anion
exchange resins capable of containing bisulflte anlons as
the anion component. Preferred anion exchange resins are
the styrene-divinyl~enze~e copolymers, usually in bead form.
The strong base exchangers have quaternary ammonium func-
tionality and are preferred. Macroreticular strong base
anion exchange resins, which are especially preerred, are
commercially available, for example from Rohm and Haas as
Amberlite IRA-900, IRA~9OOCg IRA-904, IRA-910, IRA-911, and
IR~-938. Cel-type s~rong base ion e~changers that can be
used include Rohm and Haas' Amberlites IRA-400, ~RA-400C,
IRA-401S~ IRA-402, IRA-410, IRA-425 and IRA-458 and
Stratabed 402. The names Amberlite and Stratabed are
trademarks of Rohm and Haas.
; The weak base anion exchangers, i.e., those having
polyamine functionality can also be used in this i~vention
and ~hese in¢lude Rohm and Haas' macroretlcular Amberlite
IRA-93 and S~ratabed 93 and rom Dow Chemical Company as
Dowex 3 and ~owex W~R. Dowe~ is a trademark of Dow
Chemical Company. Gel type weak base anion e~change
7~i
112~8
resins, which can be used herein, are also commercially
available from Rohm and Haas as Amberlites IR-45, IR-47
and IR-68 and are also available from Dow. Any of the
weak base or strong base anion exchange resins described
in "Ion Exchange Technology", F. C. Nachod and J. Schubert,
Editors, Academic Press, New York, 19';6, and "Ion Exchange
Resins", Robert Kunin, Robert E. Krieger Publishing Company,
Huntington, N. Y., 1972, can be used.
In summary, the operation of an ion exchange
resin bed typically consists of the following steps:
(1) Backwashing, ~2) Regeneration, and (3) Exhaustion,
which are preferably used in this invention.
1. Backwashing - This is accomplished by revers~
ing the flow of liquid (usually water) through the column.
This step rinses out any sediment trapped among the resin
beads and it allows thè beads to re-settle, thus allevia-
ting any pressure built up in the column due to expansion
and contraction of the resin.
2. Regeneration - The regeneration is simply a
reverse shift in the reaction equilibrium due to the
passage of a suitable concentrated ionic specie through
the resin. A5 an example, assume the resin is in the
heat stable salt form, represented by (R+)2~SS , and it is
being regenerated to the hydroxyl form, R+OH-, with a
concentrated caustic solution. This can be represented
; by the following equilibriu~ reaction:
(R~2~SS + 2NaOH r~ 2R+OH + Na2HSS.
3. Exhaustion - This is the part of the cycle
~here the unwanted anions, such as HSS , are`removed from
the solution and replaced with acceptable anions from the
~ 5~ 11288
resin. In the case of triethanolamine (TEA) solvent
system, this can be represented as follows:
2R OH ~ (TEAH )~SS ~ ~ (R+)2HSS ~ 2TEA t 2H2O.
A rinse step is used to wash the excess
regenerant (e.g., NaOH or other inorganic hydroxide~ from
the resin after the regeneration step. This usually
requires large quantities o~ water which are disposed of
by appropriate means as described herein. A rinse step is
also used after the exhaustion step to minimize loss of
absorbent which is recovered in the rinse waters.
A single bed or column of anion exchanger or a
plurality of beds or columns thereof can be used. Each
bed, cf course, must be taken out of service periodically
for regeneration. If one bed is used, the purge stream
is continuously removed and stored while reclaimed
absorbent (reclaimed by previous anion exchange resin
treatment3 is fed back into the system as make-up.
Pbriodically, for example once a day, the stored purge
stream is reclaimed and the reclaimed absorbent is stored
and gradually fed back into the system to continuously
make up for the purge stream continuously removed. After
the stored purge stream has been reclaimed by anion exchan~e
treatment there remains ample time for the regeneration of
the bed before the next reclamation operation.
For continuous anion exchange operation, it is
necessary to have at least two anion exchange resins beds,
one or more in anion exchange operation reclaiming
absorbent while one other is being regenerated. The use of
multiple i.on exchange resin columns or beds joined in
series is preferred and it is most preferred to have four
series-connected beds in anion exchange operation while
-16-
11288
one or more others are being regenerated. By series
connection, is meant that the output of the first bed is
fed as input to the second bed, the output of the second
bed is fed as inpu-t to the third bed and the output of the
third bed is fed as input to the fourth bed. The input
to the first bed, of course, is the aqueous absorb~nt purge
stream and the output of the fourth bed is returned to-
service in the absorption of SO~ from the waste gasO
The ion exchan-~e resin beds or columns are
operated under the best conditions for ion exchange.
Ambient temperatures and pressures are usually acceptable,
for example, room temperature and existing pressure of the
atmosphere can be used. The fluids coming into con-tact with
the beds or columns can be preheated or precooled as desired
or necessary to provide the desired contact temperatures.
I~ is important that the absorbent purge stream
and wash, rinse and regeneration fluids passed through the
anion exchange resin bed or column be of sufficiently low
density and viscosity so as not to float off the resin or
otherwise disrupt the bed. The densities and viscosities
of fluids coming into contact with the bed can be lowered
by dilution with water. For example, the incoming
absorbent purge stream if oftoo high a density or viscosity
can be diluted with water, for example, the wash water
obtained by washing the bed or column with water after
contact with the absorbent purge stream. Usually the density
of the purge stream, regenerant and wash waters should be
kept below 1.~5 g/cc which is below the normal density of
the anion exchange resins.
Since each stream -that is passed through the ion
'
-17-
6;5~
11288
exchange bed is preferably followed by a water wash step
to prevent cross contamination of process streams, there
is a substantial quantity of more or less contaminated
waste streams. However, this difficulty is readily
overcome. For instance, in regenerati,ng the exhausted
r~sin bed with sodium hydroxide there is first displaced
a free column volume of virtually pure water which can
be returned to the water reservoir or used elsewhere in
the process. The tail cut from the caustic stream contains
almost pure aqueous sodium hydroxide (with traces of
bisulfite and heat stable salts) and can be returned to
the sodium hydroxide reservoir in which extraneous salts
reach a small and wholly innocuous steady state concentra-
tion. The absorbent stream containing heat stable salts
again displaces a column volume of virtually pure water
which again is returned to the water reservoir for
subsequent resin washing or otherwise employed in the
process. The remainder of the absorbent stream, now
depleted in heat stable salts, is returned to the ab-
sorption-stripping cycle of the process.
The present invention is advantageously applied
to a continuous process for the selective removal of sulfur
dioxide with respect to carbon dioxide from a gas mixture
such as a waste combustion gas containing same wherein the
gas mixture is contacted with a mainstream of an aqueous
absorbent solution to remove sulfur dioxide from the gas
mixture and form a mainstream of rich aqueous absorbent
solution enriched in sulfur dioxide removed from the gas
mixture; the mainstream of the rich aqueous absorbent
-18-
~ 6~ 288
solution is moved to a stripping ~one where it is
stripped of sulfur dioxide to form a Inainstream of lean
aqueous absorbent solution depleted in sulfur dioxide
content; and the mainstream of the lean aqueous absorbent ~;
solution is recycled and re-contacted with the gas mixture
in the absorbing zone. The improvement of this invention
i5 employed in this context for removing divalent sulfur
oxyanions of heat stable salts which accumulate in the ~:
absorbent solution by separating a portion of the aqueous
absorbent solution containing the divalent sulfur oxyanions
of heat stable salts in a slip stream or purge stream from
the lean mainstreams, contacting the slip stream with an
anion exchange resin having hydroxyl anions displaceable
by the divalent sulfur oxyanions of the heat stable salts
to remove same from the separated slip stream whereby the
hydroxyl anions of the resin are replaced by the divalent
sulfur oxyanions of the heat stahle salts; returning the ;
separated slip stream after contact with the anion exchange
resin to one of the mainstreams, and regenerating the anion
exchange resin by contacting it with aqueous inorganic
hydroxide t such as the alkali metal hydroxides including
sodium hydroxide and potassium hydroxide to replace the
divalent sulfur oxyanions of the heat stable salts on said
resin with hydroxyl anions therQby forming a waste stream
containing the divalent sulfur oxyanions of the heat
stable salts~
: It has been discoYered that whereas most organic
bases are not good absorbents for S02 absorption, aqueous
solutions of many of their sulfite salts are. In particular,
the sulfite salts of alkanolamines in general, and
19
~ 57~ 11288
especially triethanolamine sulfite, are good absorbents.
As shown by U. S. Patent 3,904,735, if excess stripping
of the absorbent is avoided, an aqueous solution of
triethanolamine sulfite is indeed being used as the
absorbent. It has now been discovered, that, although
most organic bases are too strong to selectively absorb
S2 in preference to CO2, their soluble sulfites and,
in particular, all of the alkanolamine sulfites are suitable
absorbents which are selective to S02 absorption as opposed
to CO2 absorption. Kinetic factors make the absorbent
appear difficult to strip and increase its "apparent"
basicity. In truth, the basicities of these organic solvents
in the sulfite form are about the same, exhibiting a pH
range of 5-7.5 in the useful loading range of 0.55-0.95
mols SO2 per mol base. The pH of sodium sulfite is slightly
lower, a reflection of the limited aqueous solubility of
Na2SO3, and an indication that a lower absorption
efficiency could be expected. This is significant in that
processes have been suggested for the use of Na2SO3 for
stack gas scrubbing as discussed hereinabove.
The tri- and tetra-alkanolamines, such as those
disclosed by U. S. Patent 3,904,735 are further preferred
because they are higher boiling, e.g., less volatile
than other alkanolamines and less likely to be lost by
evaporation into the stack gas. In addition, the higher
hydroxyl functionality acts to inhibit the oxidation o
S2 to SO3 in the absorber. It has been found that the
oxidation rate is lower for the higher hydroxyl furlctionali-
ty, and is highest for the inorganic sodium sulfite.
Accordingly, whereas all alkanolamine sulfites can be
-20-
1128~
used, di-, tri- or tetrahydroxyl functionality is preferred,
and triethanolamine sulfite is an especially preferred
absorbent.
The aqueous absorbent solution can contain about 5
to about 50 ~t.% water, preferahly about 10 to about 40 wt.%
water. ~igher and lower amounts can be used, if desired.
The volume fraction of the circulating absorbent
(in the absorption-stripping cycle) wh:ich is diverted to the
absorbent reclamation step is directly proportional to the
rate of HSS generation and absorption in the circulating
absorbent and inversely proportional to a) the desired ratio
of HSS to active absorbent tolerable or acceptable in the
circulating absorbent and b) to the Eraction of the HSS
present in the 51ip stream which is removed for reclamation.
The volume fraction o the circulating absorbent diverted as
the slip or purge stream to the absorbent reclamation step
(i.e., contact with the anion exchange resin) i5 about 2 to
about 20 volume percent, preferably about 5 to about 8 volume
percent, and can be more or less depending upon particular
conditionsO It is preferred to maintain the volume fraction
of circulating absorbent diverted as the slip or purge stream
as low as possible to minimize absorbent losses in the wash
water of the resin regeneration step not recirculated to the
; purge stream for density adjustment. It is desirable, of
course, to conduct the overall process in such a manner that
HSS formation and accumulation is minimized.
It is preferred to remove for reclamation a purge
stream from the lean absorbent stream on its way from the
stripper to the absorber because it is desirable to use
the capacity of the ion exchange resin for picking up the
heat stable salt content of the purge stream and not the
-21-
fi~j5~ 288
sulfur dioxide content of the purge stream. A particular-
ly advantageous procedure in some instances is to take a
purge stream of the rich absorbent com:ing from the absorber
and to contact it with incoming waste qas before sending
it to the ion exchange resin. In this way the incoming
waste gas is quenched to a more desirable absorption
temperature and humidified or dehumidiEied to the desired
moisture content while the absorbent purge stream picks up
sulfur trioxide and acid mist from the waste gas to reduce
the formation and accumulation of heat stable sal~s in the
absorption~stripping cycle and, at the same time, gives up
some of its sulfur dioxide content to the waste yas for
removal and concentration in the absorption-stripping
cycle.
Because ion exchange resins require periodic
regeneration, there results normally an effluent stream
of dilute salts which itself could be a pollutant if
disposed of in an arbitrary manner. A novel way has been
found to avoid this by utilizing the effluent to provide
steam for use in the solvent regeneration step of this
process. Accordingly, this invention has interlocked the
solvent reclaiming and re~n regeneration steps in a unique
manner.
Specifically, the waste stream produced by
treating the exhausted anion exchange resin (loaded with
heat stable alions) with inorganic hydroxide contains the
heat stable salt anions, water, some inorganic cations
(e.g. Na ) and some SO2. It is preferred to recover the
water and, at the same time, to further concentrate the
waste stream containing the HSS- ~o simplify disposal there-
--22--
11288
ii7~i
of. This can be accomplished by contacting the incoming
waste gas with the waste stream whereupon sulfur trioxide
and acid mist are absorbed in the waste stream and SO2 is
displaced to the waste gas. In addition, or alternatively,
the waste stream can be heated in an evaporator or a
stripper to drive off water as steam which can be used in
the absorbent stripper as a source of heat and the
condensate of which can be re-used in the resin regeneration
procedure.
The prior art appears to teach that a stripping -~
tower having multiple stages should be used to regenerate
the absorbent by driving off the SO2 to provide ample
contact staging. It has been found, however, that high
liquid residence time of the absorbent being regenerated
is more important than the contact staging and that the
regeneration can be carried out in a simple still or
evaporator, provided it is designed for an ample amount of
liquid residence time. It appears as though the desorption
f S2 in a sulfite solvent is a kinetically hindered
process. Accordingly, attempts to strip rapidly~ even at
excessive temperatures, and with gross excesses of stripping
steam, will fail~ It normally would be concluded that more
stages were needed. It has been now discovered that
increasing the siæe of a single stage boiler or evaporator
which increasesthe hold-up time, i.e., residence time to
about 7 to about 20 minutes, preerably 10 to 15 minutes,
made regeneration possible with minimum quantities of
stripping steam (e.g., 7 lbs. steam/lb. SO2) and at
reasonable temperatures te.g., as low as 80C). To
illustrate this important discovery reference is made to
~ 5~6 11288
Fig. 1 which is a plot of stripping steam requirements vsO
residence time fGr a simple pot type boiler at 90C and
100 mm Hg wherein the average loading of rich aqueous TEA
absorbent was 0.87 mol So2 per mol TEA sulfite, that of
the lean aqueous TEA absorbent was 0.34 mol SO2 per mol
TEA sulfite and the average water in the lean absorbent
was about 8.7 wt.%. The overall relationship is quite
clear that by increasing the residence time the amount of
required stripping steam is indeed reduced.
It is, therefore, preferred to increase the re-
sidence time of the rich absorbent in the absorbent stripper
as much as practical. However, inasmuch as an increasa in
residence time also increases heat stable salts formation,
the residence time should not be increased to such an extent
that excessive amounts of HSS are formed. This can be done
in a number of ways one of which, as mentioned above, is to
enlarge the stripper boiler or evaporator to provide a high
holdup to throughput ratio. Another way o increasing re-
sidence time is to use conventionally sized equipment (for
example, a alling film evaporator which itself has very low
holdup and, hence, provides little residence time) and to
recirculate the absorbent being stripped through an external
circulation loop containing an amply sized reservoir so that
the total holdup time in the stripper and its circulation
loop is adequate to provide the desired high residenca time,
for example of 10 to 15 minutes.
It also has been found that with an 80-90C strip-
ping ~emperature, the quan tity o~ S02 disproportionation
products produced is reduced to only a fraction of that
produced at 100-125C stripping and, accordingly, the lower
stripping temperature puts less load on the ion exchange
resin. For this reason, it is preferred that the stripping
temperature should not exceed 100C.
-2~-
11288-C
There are many variations of the steps o~ ~he
process described above whi~h do not change the essential
nature of the in~ention described and claimed herein. One
such variation would be to introduce ~he bulk o~ the
stripped absorbent from the stripper înto the absorber at
one stage below the top and to introduce the reclaimed ;~
portion of absorbent coming from the ion exchange resin step
to the top stage of the absorber. In t~is way9 the reclaimed
absorbent, which is much leaner ~han ~he stripped absorbent,
would be able to more effecti~ely remo~e SO2 at t~le small
concentrations existing in the waste gas at the top of the
absorber, permitting a more ef~icient clean-up of the gas.
BRIEF DESCRIPTION OF THE D~AWING
Fig. 1 is a plot of stripping steam requirements
versus residence time in the stripper for a rich so2-~aden
aqueous absorbent ha~ing an ave~age loading of 0.87 mol
S2 per mol of triethanolamine sulfite, and a lean aqueous
absorbent having an average loading of 0.34 mol SO2 per mol
o~ triethanolamine sulfite and ha~ing an average water
content of about 8.7 wt.% (in the lean absorbent). Reerence
to Fig. 1 is fo~md on page 24;
Fig. 2 is a 1Ow diagram illustrating one
embodiment of this -invention wherein the tail gas from the
incinerator of a claus plant is trea~ed ~o remove and
concentrate SO2 contained thereby;
Fig. 3 is a 1Ow diagram ilLustrating another
embodiment of the invention wherein waste flue gases from
an 800 megawatt power plant burning coal contai~ing 3 wto%
~; sul~ur are treated according ~o this invention; and
Fig. 4 is a flow diagram illustra~ing a
~ 25-
-
~ 288
laboratory simulation of the process of this invention
using air and SO2 as the waste gas being treated~
DESCRIPTION OF SPECIFIC EMBODIMENTS
In one embodiment the invention comprises an
improved process for the removal of sulfur dioxide from an
effluent gas stream (power plant stack gas or the like)
by means of selective absorption in an aqueous solution
of an organic base sulfite, preferably an alkanolamine
sulite and especially preferably triethanolamine sulfite,
comprising about 50-90 wt. percent of the organic sulfite,
about 10-40 wt. percent water, and some quantities ~about
5-30 wt. percent) o recirculated heat stable SO2 side
reaction products, i.e., heat stable salts. The improved
process in one full embodiment comprises the steps of
a) absorption by countercurrent contact in an
absorption tower, for example, containing
the equivalent of three (3) or more stages,
such absorption being carried out near
atmospheric pressure (or above if the feed gas
is so available), and at about 20 to about
70C;
b) desorption of the SO2 from the absorbent by
steam stripping in a single stage still or
evaporator providing a liquid residence time
of about 3 to about 30 minutes, preferably
about 10 to about 15 minutes and operated at
a reduced pressure of about 50 to about 350
mm Hg, preferably about 100 to about 150 mm
Hg and a temperature of about 80 to about
-~6-
~ 112~8
100C;
c) removing the So2 overhead from step b),
condensing the water therefrom, and compressing
the SO2 to atmospheric pressure or above for
further use or storage;
d) recirculating the bulk of the stripped
absorbent back to the absorber;
e) removing from~the stripped absorbent a purge
- stream of about 2 to about 20 wt. percent,
preferably about 5 to about 8 wt. percent, for
reclaiming with respect to heat stable salts,
reclaiming this purge stream via passage through
an anion exchange resin and returning the
reclaimed absorbent, including resin bed
washings, to the top o the absorber;
f) regenerating the anion exchange resln with,
for example, an aqueous solution of sodium
hydroxide, taking the effluent, including resin
bed washings, to a boiler and boiling to
concentrate the salts to a slurry of about
30 to about 7Q wt. percent solids, and to
provide about 15 to about 30 lbs. steam;
g) using the steam generated in step f) for at
least part of that required in the stripping
step b), either by direct introduction into
- the still or evaporator or indirectly via heat
exchange~ and using part or all of the conden-
sate thererom to provide water for further
use in the ion exchange step, both for the
dissolution of the sodium hydroxide (caustic~
~-27-
~ i 11288
and for resin bed washing.
Temperatures and pressures above and below those
specified in the above steps can be ernployed if desirable
or advantageous.
-28-
~ 6 11288
The following examples are presented. In the
examples, the following designations have the meanings
given below:
cc cubic centimeter
g grams
wt. ~ weight percent
ppm parts per million based on weight
ppmv paxts per million based on volume
M lb mols thousand pound mols
LTPD long tons per day
; M. Wt. molecular weight
M Molar
meq milliequivalent
ml milliliter
MSCFM thousand standard cubic feet per minute
MMSCFD million standard cubic feet per day
psia pounds per square inch absolu~e
psig pounds per square inch gauge
TEA triethanolamine
HSS heat stable salts having divalent
sulfur oxyanions, e.g., SO4 , S2O3 ,
SxO6~ and the like, as described
hereinabove.
Also r unless otherwise specified all parts and percentages
are on a weight basis, and all temperatures are on the
Fahrenheit scale.
EXAMPLE 1
Fig. 2 illustrates a unit and process for treating
the tail gas effluent from a typical three-stage Claus
-29
.
11288
plant producing 280 LTPD sulfur at 96% efficiency. The
composition and conditions of the feed gas before and after
incineration are given in Table I. The treated gas vented
by stack has So2 and SO3 concentrations less than 500 and
15 ppmv, respectively. The moisture level is about
100F dew point.
The purpose of the gas pretreatment section is to
provide an absorber feed gas at 100F containing minimal
SO3 and H2SO4 mist. A stream 1 of tail gas from the Claus
plant incinerator enters a waste heat boiler 2 at 1000F
and exits through stream 3 at 600F producing 250 psi steam
; in line 4 from water entering the boiler 2 through line 5.
The exlt temperature of stream 3 is well above the 500F
dew point of H2SO4 thus avoiding acid condensation which
would create a difficult corrosion problem. The gas is
then quenched in quench tower 6 with an alkaline waste-
water stream 7 (produced as described hereinafter) to
remove SO3. The liquid and gas are separated in an SO3
knock-out drum 8 at the top of the quench tower 6.
The quenched gas stream 9 is cooled to 100F
(20F below the absorber temperature) in the feed cooler
10 using cooling tower water. Condensation is removed
in a high-efficiency demister 11 and is disposed of
; through waste stream 12. This step is necessary to
minimize the carryover of sulfuric acid mist to the
absorber.
The cool gas stream 13 enters the base of a
six-tray absorber 14 where it countercurrently contacts
the absorbent, TEA. The So2 content of the gas is reduced
from about 10,500 to ~ 500 ppmv (more or less depending on
-30-
~ 57~ 11288
pollution regulations). The treated gas stream 15
leaviny the top of the absorber has a 100F dew point.
It is sent to an induced draft blower 16 and then to the
stack. The treated gas can be rehéated with the
incinerator tail gas stream 1 .in a regenerative heat ex-
changer if required to control pluming and ensure
buoyancy.
T~BLE I
After Incineration
Feed Specifications Cl-aus Tai-l Effluent And Cooling To 100F
Feedrate, MSCFM 18.338 16.2 (at 100F)
Feedrate, MMSCFD 26.407 23.3 (at 100F)
Pressure, psia 17.7
Temperature, F 280
Density, lbs/cu.ft. 0.0558
Molecular Weight 25.05
LTPD of sulfur 10
-31-
;it7~i 1128~3
olP tl') C~ r
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-32- .
576
11288
Lean absorbent stream 17, which has been cooled
to 120F by cooler 18 and to which regenerated absorbent
stream 19 and make-up absorbent stream 20 from make-up
reservoir 21 and water stream 22 have been added, ls fed
into the top of the absorber 14. It passes downward,
countercurrent to the gas stream, absorbing SO2. The rich
absorbent stream 23 exits the bottom of the absorber 14 and
is heated to about 160F in.heat exchanger 24 by the lean
absorbent stream 17 which is cooled to about 150F. The ::
rich absorbent stream 23 then enters the absorbent stripper
25 where water and SO2 are evaporated at low pressure and -::
moderate temperature. Stripper 25 is a falling-ilm type
evaporator with indirect steam heating. The lean absorbent
stream 26 is sent back to the absorber 14 via exchanger 24
and cooler 18, while a purge stream 27 is sent to regene-
ration for HSS removal (to be described hereinafter~. The
stripper overhead stream 28 which is laden with water, is
condensed in condenser 29 and separated in separator 30.
The wet SO2 vapor stream 31 from the separator 30 is pumped
by vacuum compressor 32 to the Claus plant. The condensed
water stream 33 is sent to the ion exchange regeneration
system to be hereinafter described and/or to the purge
stream 27.
The purge stream 27 of the circulating absorbent
is treated by Amberlite IRA-910, a strong base anion
exchange resin, to remove sulfate and any HSS buildup in
the system. The resin unit comprises four beds 34 operated
on a four-hour total cycle of four equal time steps each
comprising: 1. *bsorbent exchange, 2. Absorbent displace-
ment by water, 3. Regeneration by caustic, and 4. Caustic
-33-
~ 76 11~88
displacement by water.
The purge stream 27 is diluted with tripper
overhead condensate 33 to reduce its density and viscosity
for easier flow through the ion exchange beds 34 to avoid
bed disruption. The diluted purge stream 27 is then
passed through an ion exchange bed 34 to remove 80~ of the
HSS, 65% of the SO2 present and thus regenerate 80~ of the
TEA. The regenerated absorbent stream 19 is returned to
the absorber 14. Because the output from resin bed 34 is `
necessarily intermittent and because the water content will
vary, a surge vessel 36 and a circulating pump 37 are
provided to induce mixing via recycle loop 38.
While one ion exchange bed 34 is used for absorbent
reclaiming, the others are regenerated by a three-step
process which consists of: ;
1) Water washing with condensate stream 33 to
displace the absorbent from the bed 34~ This
water effluent is combined with the reclaimed
absorbent in surge vessel 36 and mixed via the
recycle loop 38.
2~ Regenerating the bed 34 with a 10 wt. % sodium
hydroxide solution from sodium hydroxide
reservoir 39 and
3) Water washing the bed 34 with process water
stream 40 to displace and wash out excess
; caustic.
The alkaline effluent streams 41 and 42 from steps (2) and
t3), respectively, are used in the quench tower 6. An
alkaline effluent surge tank 43 is provided to accommodate
the intermittent nature of effluent streams 41 and 42.
-34-
~ 7~ 11288
Make-up sodium hydroxide i5 supplied through line 44.
Circulating pumps 45 are provided at appropriate locations
to provide adequate circulation.
EXAMæLE 2
Fig. 3 represents the process of this invention
applied -to 1ue gas of an 800 megawatt power plant burning
coal containing 3 wt.~ sulfur. The flue gas feed, which
has been passed through an economizer (not shown), and
from which 99~ of the fly ash has been mechanically
removed, is represented by stream 51t which typically
consis-ts of about 243 M lb mols~hr of gas at about 150C
having the composition yiven in Table II which also gives
the approximate compositions of the process streams herein-
after described.
:
-35-
7~ 11288
a~ r~ I~ I~ r~ o o o o
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o ~ ~C
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O rl O~
æ ~
C~ O O ~ ~, ~,
I I I ~ ~ ~ O O O
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I~ r~ o In O '
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r~ o ~I
o o~ ~ o ~ ~C~ o o O
o o o O~1 0 0 0 0
U~
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o~ ra 00 ~0 00
U~
U~C ~1 ~ Ul ~ `D ~ -1 0
Cl~ O o I ~C~l ~ O O O
~ JJ , , I
o ~ ~" o o I o O o O O O
C!. aJ
H ~U~
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O I : ~ C`J ~ ~
~1 O ~~1 ~1 0 0 0 0 0
OO O O I O O O O
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¢ a O o O I o O O O
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ta O
^ e.~ 0~ ~0 ~ I I I I t
C~ ~ ~ C~l I I I I I
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J- aJ ~ ~ 0 5~ ~
C~ ~ O aJD ~ D h D ~ ~'J h H ~:1 '~ 00
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c~ ~JJ ¢ 1~t/~ ~ trJ P. U~ h 1- JJ
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--36--
;CJ~ 11288
~ C~ o o oU~
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C~ I ~ ~ C~
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a) o ,i co ~`I
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-37-
~ 57 ~ 11288
Stream 51 can be cooled and dried via quench
tower 5~, in which a portion of the water condensate is
recirculated through cooler 53 and reintroduced to the top
of the tower as stream 54 as the quench media. Excess
water stream 55, containing most of the remaining fly ash,
is used in the hereinafter-described absorbent reclaiming
step or sent to waste, e.g., an ash pond or pile.
The resulting cool gas stream 56, at 37C, enters
the base of absorber 57 wherein it flows countercurrent to
the downcoming absorbent TEA, and its SO2 content is
reduced from about 2000 ppm to about 100 ppm (more or less
as is required to meet pollution restrictions). It exits
as gas stream 58 from the top of the absorber 57 and is
sent to the stack (not shown) with or without an inter-
mediate reheating step as may be desired.
Recycle lean absorbent stream 59 from the
absorbent stripper 67 is cooled to 37C by cooler 60 and has
added to it a reclaimed absorbent stream 61 and any necessa-
ry makeup absorbent through stream 62 from makeup absorbent
reservoir 63. Lean absorbent stream 59 is fed into the top
o~ the absorber 57 and passes downward, countercurrent to
the gas stream, during which time it picks up SO2 from the
gas stream, increasing its loading from 0.2 to 0.8 mol
S2 per mol of triethanolamine sulfite. It exits the
bottom of the absorber 57 as rich absorbent stream 64
into a reservoir 64A from which it is passed through the
heat exchanger 65 where it is heated by the stripped
lean absorbent stream 59 to about 75C, the lean stream 59
being simultaneously cooled from about 90C to about 55C.
The heated rich absorbent stream 64 ls joined by
~38-
~ 7~ 11288
stripper recycle stream 66 and enters the absorbent thin
film evaporator, where water and SO2 are evaporated at
100 mm Hg and 90C in a ratio of 7 lbs. H20 per lb. S02.
The lean absorbent effluent stream 68 exits the reservoir
of the stripper 67 and 90% by volume is recycled via
stream 69 back to the stripper 67. Most of the remaining
exit stream 70 from stripper 67 is sent via stream 59 back
to the absorber 57 via the heat exchanger 65 and cooler 60,
while a 7% (by volume of the remaining exit stream 70)
purge stream 71 is removed for reclaiming with respect to
heat stable salts.
The H2O/SO2 overhead stream 72 from the stripper
67 passes through condenser 73 and thence via stream 74
into separator 75 from which the SO2 can be pumped via
line 76 to storage or any desirea chemical treatment step.
The water condensate stream 77 from the separator is used
in the hereinafter-described absorbent reclaiming step.
The purge stream 71 is diluted with stripper
condensate via lines 77 and 78 and quench tower condensate
from stream 55, and filtered in filter 79 to remove any
fly ash which may have bypassed initial mechanical removal
and been entrapped either in the absorbent or the quench
tower condensate. The filtrate stream 80 is passed through
one of two (or more) anion exchange beds 81 containiny
Amberlite IRA-910 a strong base anion exchange resin to
remo~e 90% of the heat stable anions, 70~ of the SO2
present as sulfite-bisulfite, and thus regenerate 80% of the
TEA as a free base. The ion exchange resin bed effluent 82
containing reclaimed absorbent is passed into a reservoir 83
fxom which it can be returned via lines 84 and 61 to the
-39
~6~ 11288
absorber 57. A portion 85 of this stream may be sent
to the stripper 67, if desired, to provide addltional
stripping water. Because the output from anion exchange
resin bed 81 is intermittent~ and because the water
content will vary, it is beneficial to induce mixing in
absorbent reservoir 83 via a recycle loop 86.
While one anion exchange bed 81 is used for
absorbent reclaiming, another ion exchange resin bed is
being regenerated by a three-step process which comprises
1) water washing with water condensate stream 77, the
effluent of which is combined with the reclaimed absorbent
in reservoir 83 and mixed via the recycle loop 86, 2)
regeneration with 10 wt.% aqueous sodium hydroxide from
holding tank 88 via line 89, and 3) water washing with
steam condensate via line 90 from steam condensate reservoir
91 to remove excess sodium hydroxide from the bed. The
effluents from steps 23 and 3) are sent via line 92 to
water stripper 93 where the water is boilded off to leave
a waste stream 94 of concentrated sodium salts o the heat
stable anions. The steam rom water stripper 93 can be
effectively utilized via line 95 to provide the heat in the
absorbent stripper 67, and the condensate therefrom can be
cooled via cooler 96 and sent via line 97 to reservoir 91,
from which it can be recycled as wash water for the resin bed
87 through stream 90 and/or can be sent via line 98 to the
sodium hydroxide reservoir 88 to adjust the concentration of
aqueous sodium hydroxide therein. Additional sodium hy-
droxide (caustic) is added to the reservoir 88 through line
99.
-40-
~ S7 ~ 112~8
Water condensate stream 77 separated ~rom the
S2 gas emanatin~ from the stripper 67 can be diverted to
the sodium hydroxide reservoir 88, as needed, via line 100,
or it can be used to wash the ion exchange resin bed 81
via line 101, or it can be diverted, as needed, via line
102 to the recycle loop 66 of the absorbent stripper 67,
or it can be used via line 103 to dilute the purge stream
71 before it enters the ion exchange resin bed 81. Pumps
104 are appropriately placed to provide adequa e
circulation.
Steam is provided to the water stripper 93 ~ia
line 105 and condensate removed via line 106.
EXAMPLES 3 AND 4
_.
Referring to Fiy. 4/ the feed gas employed was
made up of an airstrea~ 107 to which was added sulur
dioxide initially from a suitable source via stream 108.
During operation recycle sulur dioxide was used via
stream 109 with enough make up sulfur dioxide if needed
from stream 108 to provide a sulfur dioxide level of 2000
to 3000 ppm. The resulting feed gas was passed via stream
110 into the bottom of a 4-stage bubble cap absorber 111,
and exited the top via stream 112. Gas samples were taken
at streams 110 and 112 for analyses by both the Dr~ger tube
technique and gas chromatography.
The lean absorbent stream 113 was passed into the
top of the absorber 111, and the rich stream 114 exited
the bottom. The rich ~bsorbent stream 114 was passed through
a let-down valve into the vacuum portion of the stripper
system and pumped by pump 115 into the stripper 116. The
.
11288
6~;i7~
stripper 116 comprised a stirred steam heated pot from
which both vapor and lean absorbent exited via stream 117
into the condenser 118. The water vapor was condensed in
the condenser 118 and rejoined the lean absorbent stream
113 which was pumped by pump 119 out of the vacuum portion
of the system and recirculated to the absorber 111. A
portion of this stream 120 was withdrawn for reclaiming
with respect to heat stable salts, and the absorbent
volume was made up by the addition of reclaimed
absorbent via line 121.
The SO2 stream 122 emanating from the top of
condenser 118 normally would be treated further to produce
sulfuric acid or su].fur. However, it was, in this case,
recirculated to the incoming air stream 107 through the
vacuum pump 123.
The purge stream 120 was passed through a strong
base anion exchange resin, identified as Amberlite IRA-910
which treatment removed 80% of the heat stable salts and
65% of the SO2 present as sulfite or bisulfite. The
2n resin was subsequently washed and regenerated with 10 wt.%
aqueous sodium hydroxide yielding an aqueous effluent of
the sodium salts of the heat stable anions. This effluent
was not distilled to produce steam as it would be in an
industrial application to concentrate it prior to disposal.
Opearting parameters for Examples 3 and 4 are
given in Table III.
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' '. ' ~'
11288
TABLE III
OPERATING PARAMETERS FOR E~MPLES 3 AND 4
Example No.
Stream No. ~
107 Air Flow, liter/min 100 50
110 SO2 to Absorber, ppm 2230 2590
113 Lean Absorbent Flow, g/min 7.6 8.0
Lean Absorbent Composition, wt.%
Triethanolamine sulfite 58.9 56.8
Heat Stable TEA Salts 11.7 20O1
Water 26.0 20.0
S2 3.37 3.06
Lean Absorbent Loading, mol
SO2/mol (TEAH)2SO3 0.34 0.32
Absorption Temperature, C 37 37
112 SO2 output, ppm 460 200
S2 removed, ~ 79 92
~verage tray efficiency, % 33 47
114 Rich Absorbent loading,
mol SO2/mol (TEAH)2 SO3 0.86 0.76
Loading change 0.52 0.44
Stripping pressure, mm Hg 100 100
Stripping temperature, C 90 90
Residence time in stripper,
minutes g 9
Steam consumption, gms ~20/gm
2 8.4 1005
Rate of heat stable salt
formation 0.44 0.20
120 Purge stream out, g/min 0.7 0.2
Percent purge 9 2.5
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