United States
Environmental Protection
Agency
Industrial Environmental Research
Laboratory
Research Triangle Park NC 27711
Research and Development
EPA-600/S7-81-140 Sept. 1981
Project Summary
Demonstration of Wellman-
Lord/Allied Chemical FGD
Technology: Final Report and
Demonstration Test Second
Year Results
R. C. Adams, S. W. Mulligan, and R. R. Swanson
Performance of a full-scale flue gas
desulfurization unit to demonstrate
the Wellman-Lord/Allied Chemical
process was evaluated for a period in
excess of 2 years. The Wellman-Lord/
Allied Chemical process is a regener-
able process employing sodium sulfite
wet scrubbing, thermal regeneration
of the spent scrubber solution, and
reduction to elemental sulfur of the
recovered SO2.
Test program results indicate that
89 to 90% of the SO2 can be readily
removed from the flue gas in a long-
term dependable manner. Reliability
of the coupled absorber and regenera-
tion system for the second year was
61%. For the last 7 months, it was
74%. The major operating limitations
were reduction unit problems, but
unscheduled outages of the evaporator
and the booster blower and start-ups
and shutdowns also contributed to
down time.
As expected, the energy require-
ments of the process, primarily for
thermal regeneration of the scrubber
solution and subsequent recovery of
SOa, were quite large, amounting to
12% of the boiler heat input derived
from fuel. Actualannualized operating
cost was 14.9 mills/kWh, using 1978
prices for raw materials and utilities.
Credits for the sale of byproduct sulfur
amounted to only 0.2 mills/kWh.
The reported operation and per-
formance occurred after some modifi-
cation to the boi|er to increase inlet
flue gas temperature and after needed
improvements to the FGD plant
identified during initial operation were
implemented. Design limitations af-
fecting overall performance were lack
of redundancy, regeneration area
capacity of only 80% of full load,
underdesign of the purge solids dryer,
and limited turndown capability.
This Project Summary was devel-
oped by EPA's Industrial Environmen-
tal Research Laboratory, Research
Triangle Park, NC, to announce key
findings of the research project that is
fully documented in a separate report
of the same title (see Project Report
ordering information at back).
Introduction
In 1972, the EPA entered into a cost-
shared contract with Northern Indiana
Public Service Company (NIPSCO) to
design, construct, and operate a flue gas
desulfurization (FGD) plant that uses
the Wellman-Lord/Allied Chemical
(WL/A) FGD process. NIPSCO entered
into contracts with Davy Powergas (now
Davy McKee) to design and construct
the unit and with Allied Chemical (now
Allied) to operate the plant. The FGD
unit was retrofitted to NIPSCO's Mitchell
No. 11 boiler in Gary, IN.
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The WL/A process developed by the
two design organizations is a regener-
able process, based on the recovery of
concentrated SOz and its subsequent
reduction to elemental sulfur. The
product is sold to partially offset the
process costs. This was the first coal-
fired WL application, as well as the first
joint WL/A installation.
To ensure that potential users are
fully aware of the commercial practical-
ity of the WL/A process, the EPA
conducted a test program that included:
• A baseline characterization of the
host boiler.
• A performance test required by
contract to demonstrate compliance
with the performance guarantees.
• Testing for a 2-year period to
demonstrate long-term perform-
ance and dependability.
This summary highlights results and
conclusions of the test program.
Principal objectives of the test pro-
gram were:
1. Verification of the reduction in pol-
lutants achieved by the WL/A
process FGD unit.
2. Validation of the estimated technical
and economic performance of the
demonstration unit.
3. Assessment of the applicability of
the WL/A process to the general
population of utility boilers.
The test program was designed to attain
these objectives to the maximum extent
possible. Emphasis on the various
objectives was sometimes redirected
due to test program findings and
operating difficulties but, in general, the
program goals remained unchanged.
The test design featured continuous
monitoring of SO2 removal performance.
Evaluation of the data was in response
to the test objectives and focused on the
dependability and economics of SOz
removal capability while operating the
boiler to provide the expected range and
variability of the flue gas properties with
a single coal type.
Overall Process Design
Mitchell No. 11 is a 115-MW pul-
verized-coal-fired, balanced-draft boiler
with cold-end electrostatic precipitator
(ESP) particle control. The boiler was
designed to use a coal with a nominal
sulfur content slightly above 3% by
weight. The FGD unit was designed to
accept flue gas at SOz concentrations
equivalent to this sulfur level in the coal.
Flue gas was fed to the FGD plant by the
boiler's two induced draft (ID) fans.
Before retrofit of the FGD plant, the flue
gas went to a stack shared with another
boiler. A quick opening'damper was
installed in the duct to that stack to
bypass the FG D pla nt when not operati ng
and to protect the boiler from damage
during upsets. Normally, the FGD plant
operated with the bypass damper
closed.
The WL/A FGD process removes S02
from the flue gas stream by scrubbing
with an aqueous sodium sulfite solution
and subsequent thermal regeneration
to recover the SO2. The solution is free
of solid material. The liberated SO2 is
then reduced to elemental sulfur which
is sold. The FGD unit was designed to
remove 90% of the SOz delivered with
the flue gas at flue gas rates equivalent
to a boiler load of 92 MW (80% of full
boiler load). The processes are propri-
etary designs of Davy McKee and Allied.
Logical separation of the various
process steps are:
• SOz absorption - Davy.
• SOz recovery and scrubber solu-
tion regeneration - Davy.
• Purge treatment - Davy.
• SO2 reduction - Allied.
The following description is based on
Davy and Allied non-proprietary design
data.
Figure 1 shows the process steps. The
FGD plant accepts the total flue gas
stream from the discharge of the
boiler's ID fans using a booster fan to
overcome the flow resistance through a
prescrubber and an absorber. The pre-
scrubber is a single-stage orifice
contactor designed to remove paniculate
matter and cool the flue gas before the
S02 absorption step. A pump recircu-
lates the scrubber water from a sump
back to the contactor. In order to control
solids buildup in the liquid stream, a
purge stream is withdrawn and makeup
water is added to the prescrubber to
compensate for this loss as well as to
humidify the flue gas. The purge stream
is sent to the power station's fly ash
settling ponds. Particles not removed in
the prescrubber are removed with a
filter in the spent absorbing solution line
leaving the absorber. The wash water
from periodic washing of this filter is
also discharged to the power station's
fly ash settling ponds. These are the
only waste streams expected to be
discharged from the FGD plant.
The cooled, humidified flue gas leaves
the prescrubber and enters the bottom
of a three-stage absorber where the gas
is contacted with the sulfite solution
flowing countercurrent to the gas
stream. The solution absorbs the SO2
and the treated flue gas, saturated at
about 54.4°C (130°F), is then discharged
to the atmosphere through an integrally
mounted stack. Direct-fired reheat of
the flue gas with natural gas as fuel was
provided but never used because of
limited gas supplies.
The absorber is a three-tray column.
An absorber demister pad above the top
contact stage prevents entrained ab-
sorber solution from being exhausted
with the treated gas. Each contact stage
of the absorber hasa separate recircula-
tion system to promote good gas/liquid
contact. The absorbing process is con-
ducted at about 54.4°C (130°F), and
spent absorbing solution is withdrawn
from the bottom contact stage. Re-
generated absorbing solution is added
to the top contact stage. The process
chemistry for absorption is:
SOz + Na2S03 + H20 - 2NaHSO3
The bisulfite-rich solution is then
passed to the evaporator-crystallizer
unit. The evaporator-crystallizer is a
single-effect forced circulation unit. The
heat exchanger employs steam and the
clean condensate is discharged for
reuse by the host boiler. The heat ,
supplied to the liquor decomposes the |
sodium bisulfite solution to sodium
sulfite (which crystallizes out), SO2, and
water:
heat
2NaHSO3
Na2S03 + S02 + H20
The SO2 and water vapor are discharged
overhead from the evaporator. The wet
gas stream is cooled to condense and
separate the water vapor from the gas
stream, providing a S02 feed stream of
low humidity for the SO2 reduction unit.
The condensate that was removed is
stripped of dissolved S02 a nd is added to
the sodium sulfite slurry discharged
from the evaporator salt leg. The sodium
sulfite and condensate, along with
makeup sodium carbonate, are mixed to
provide a solids-free sodium sulfite
solution suitable for reuse in the
absorber.
Oxidation of some of the sodium
sulfite during the absorption step to
form sodium sulfate is unavoidable:
Na2SO3 + O2
2Na2SO4
A portion of the spent absorber solution
leaving the absorber is passed through a
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Treated Flue Gas
Figure 1.
\FGD ProceddBoundary
Block flow diagram of major process steps.
J
purge treatment system to separate and
dry the NasSCh for eventual sale. The
purge treatment unit was designed to
minimize the amount of purge and to
produce a salable byproduct of sodium
sulfate. Purge requirements are reduced
by subjecting part of the absorber
product stream to a fractional crystalli-
zation process. Sodium sulfate crystals
are produced by chilling the solution
under very specific conditions. This
treatment allows most of the active
sodium compounds (sulfite and bisulfite)
to remain in the mother liquor which is
separated from the crystals by a
centrifuge. This mother liquor is then
returned to the main process and the
crystals are dehydrated in a drying
system. The final product is a dry
sodium sulfate suitable for marketing.
Purge treatment requirements depend
on the amount of sodium sulfate and,
possibly, sodium thiosulfate formed in
the process and on the allowable con-
centration in the absorbing solution of
these inactive components. The amount
of sodium sulfate formed may be a
function of the amount of excess air in
the flue gas.
The S02 discharged overhead from
the evaporator-crystallizer is reduced to
elemental sulfur during the S02 reduc-
tion step, a two-stage process employ-
ing a primary catalytic reaction of SOa
with natural gas to produce sulfur and
some hydrogen sulfide(H2S) followed by
a secondary Claus conversion system in
which the H2S is reacted with residual
S02 to produce additional sulfur. The
primary reaction system consists of two
packed-bed regenerative heat exchangers
and a catalyst-packed reduction reactor.
These vessels and their connecting
flues are refractory-lined for protection
against high temperatures and corrosive
gases.
The regenerative heat exchangers
remove the heat from the gases leaving
the reactor, and utilize this heat to raise
the temperature of the gases entering
the reactor. At appropriate intervals, the
duties of the two regenerative heat
exchangers are alternated; i.e., the
packed bed heat exchanger that was
heating the entering gases becomes a
cooler for the gases leaving the reactor.
Essentially half of the S02 in the feed
is reduced to elemental sulfur by direct
reaction with the reducing gas:
2S02 + CH4 —»• C02 + 2H20 + 2S
Simultaneously, most of the remaining
S02 is converted into H2S and additional
sulfur by a similar reaction:
SSOs + 2CH* —*• 2C02 + 2H20 + 2H;
The hot gases from the primary reaction
system pass through condensers where
the sulfur is removed and sent to
storage. The gas flows and operating
conditions in the primary reaction
system are carefully controlled so that
the mixture of H2S and unconverted SOa
in the product gases from the primary
reaction system closely approximates
the ideal volumetric ratio (two parts H2S
to one part SO2) required for the sub-
sequent Claus reaction.
2H2S + S02
At the same time, maximum utilization
of the reducing agent is achieved and
the formation of undesirable side-reac-
tion products is minimized.
The Claus con version system is a two-
stage unit. Associated with the Claus
unit is an interstage condenser from
which elemental sulfur also passes to
storage. A final condenser follows the
second stage converter to recover the
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last portion of the elemental sulfur
product.
The tail gases from the SOz reduction
unit pass through an incinerator where
natural gas is burned in air to oxidize the
remaining H2S to SOz. The hot gases
from the incinerator are admixed with
the untreated flue gases at the booster
blower inlet, thus avoiding having to
discharge a stream of small volume but
containing a relatively high concentra-
tion of S02.
Design Limitations
The FGD demonstration would have
benefited from a more conservative
design. Deficient performance was due
partially to design related causes. The
most significant design decisions
affecting performance were:
• The booster blower and the ab-
sorber were designed for flue gas
flows in excess of boiler full-load
flow. The absorber was designed to
remove the expected amount of
SO2 at boiler full load. The effect on
performance was positive.
• Capacity for recovery of the SOa
(evaporator, purge treatment, re-
duction) was based on an expected
load factor of 80% of full load.
Surge capacity provided by storage
tanks for absorber feed solution
and absorber spent solution was
limited to about 4 hours. The effect
was to limit either the amount of
S02 removed or maximum utiliza-
tion of the boiler. The latter pre-
vailed because of a policy to
operate with a closed bypass and
the boiler limited to 80% load factor
(92 MW).
• Baseline testing revealed that flue
gas flow at a given load exceeded
the design flue gas flow signifi-
cantly. High excess air levels in the
flue gas and additional flue gas
due to higher-than-expected boiler
heat rates were identifiable causes.
• The FGD plant was designed with
virtually no redundancy as installed
spares.
• Initially, the evaporator circulating
pump had a steam-turbine drive.
Loss of high pressure steam during
boiler shutdowns delayed start-
ups because slurry in the evapo-
rator had to be removed and
diluted. The deficiency was cor-
rected after 1 year of demonstra-
tion with the installation of an
electric-motor drive.
• The purge solids dryer was under-
designed or a misapplication and
this prevented full recovery of
sodium sulfate from the purge
stream.
• The reduction unit had less turn-
down capability than the absorber/
evaporator.
• The FGD plant was designed to
take flue gas at 149°C (300°F).
Initially, actual temperatures were
substantially lower due in part to
cooling of the flue gas by in leakage
air. Some of the baskets were
removed from the air heaters to
provide flue gas at temperatures of
149°C(300°F) and above.
Test results that were design-related, as
opposed to process related causes, have
been identified as such in the discussion
of test results that follows.
Conclusions
The WL/A demonstration test pro-
gram consisted of three major test
phases:
• Baseline testing.
• Acceptance tests to verify per-
formance guarantees.
• Two-year demonstration test pro-
gram.
Baseline Testing
Major baseline testing occurred in
two time-separated stages. The baseline
test was conducted prior to completion
of construction of the FGD plant. Flue
gas characterizations correlated with
boiler operating settings showed that
some of the flue gas properties, par-
ticularly flow and temperature, differed
from those used to design the FGD
plant. The differences had a profound
effect on the criteria for acceptance
testing, and the baseline data were very
useful for defining those criteria. The
data also were useful for defining the
range of testing.
Following completion of the first year
of the FGD process demonstration,
baseline tests were repeated with
several objectives:
• Establish an updated boiler per-
formance baseline for comparison
with boiler performance when the
FGD plant is operating.
• Obtain an updated characterization
of the flue gas leaving the boiler.
The FGD plant was down and com-
pletely isolated from the boiler. The
collection of additional baseline data
was necessary to establish to what
extent a substantial rebuild of the boiler
since the first baseline test and a
change in the type of coal being burned
had affected the performance of the
boiler. Flue gas volumes, flue gas
temperatures, and coal quality were of
particular interest because of their
impact on FGD operation and per-
formance.
Flue Gas Volume
Flue gas volumes were 27% higher
than that used for FGD plant design,
151 mVs at 148.9°C (320,000 cfm at
300°F). An excess of inleakage air
appears to be the major contributor to
the high flue gas volumes, but heat
rates that were higher than the new
boiler design heat rates probably had
some effect. Factors contributing to the
higher heat rates were high turbine
steam rates and, at low load factors,
combustion air in excess of the operating
set point. Both the high excess air and
the high steam rates add to the volume
of flue gas that the FGD plant must treat.
The high steam rates would be expected
to add to the amount of SOa to be
removed.
Flue Gas Temperature
Flue gas temperature averaged
148.9°C (300°F) at all load levelsduring
the second baseline test. This was
substantially higher than the tempera-
tures measured during the first baseline
test. The higher temperatures were due
to the removal of part of the heat
transfer surfaces of the air preheaters in
a deliberate attempt to raise the
temperatures above the dew point and
thus prevent the scaling and corrosion
that was occurring at the FGD booster
fan. The temperatures were well above
the sulfuric acid dew point.
Coal Quality
The high sulfur (—3%) coal used
during FGD operation was burned
during the second baseline test and
during subsequent demonstration test-
ing. The coal sulfur content was 0.2 to
0.3 wt. % lower than that of the coal
burned during the first baseline test.
Except for expected differences in
variability of the ash and moisture
contents, the quality of the two coals
was about the same.
Acceptance Test
Process performance guarantees
were met or exceeded as confirmed by
acceptance testing. The boiler was
operated to provide the flue gas flows
used to design the FGD plant. Despite
flue gas dilution from high excess air
levels that was found during baseline
testing, the SOa feed rates equalled or
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exceeded the design expectation of
0.6101 kg/s(4842 Ib/h).
Specific performance criteria were
met or exceeded:
1. S02 removal of 90%, 2-hour
average, or better was achieved
for 261 hours at flue gas flows of
151 mVs (320,000 cfm) or higher
and was achieved for 84 hours at
about 183 mVs (388,000 cfm) or
higher.
2. Paniculate emissions did not
exceed 40 ng/J (0.1 lb/106Btu)of
boiler heat input at either the
lower (design load) or higher load
conditions.
3. The consumption of steam, natural
gas, and electrical power averaged
76% of the performance guarantee
requirements at design load con-
ditions.
4. Soda ash consumption averaged
less than 0.069 kg/s (6.6 tons of
sodium carbonate per day) which
was the limit set in the perform-
ance guarantees that was not to
be exceeded during the design
load period. Sodium carbonate
consumption of 0.069 kg/s (6.6
tons/day) is equivalent to 0.152
mol Na/mol S removed from the
flue gas at a SO2 feed rate of
0.6101 kg/s (4842 Ib/h) and 90%
removal.
5. Sulfur product purity was greater
than 99.5% at both design- and
high-load conditions.
Two-Year Demonstration
Program
The test program, as originally
planned, was to monitor performance of
the FGD plant for 1 year. The FGD plant
achieved only 90 days of accumulated
operation during the first year due in
part to boiler operating problems. The
principal boiler problems that prevented
FGD operation were unstable flue gas
flows and steam pressures. They were
the result of poor coal quality, that
exacerbated some problems with the
coal feeding equipment, and of boiler
feedwater quality problems. Problems
were also encountered at the boiler/FGD
interface, in particular, booster blower
and damper problems. A midyear
review resulted in the initiation of a
plant improvement program with the
goal of correcting the major problems.
The program was targeted for comple-
tion during a scheduled boiler shutdown
which coincided with the end of the 1 -
year demonstration. The test program
was continued for another full year to
evaluate the effects of the boiler and
FGD plant modifications.
Table 1 lists the major improvement
projects. The need for these improve-
ments are identified with the following
FGD plant process and design limita-
tions:
• Steam and flue gas fluctuations
limited or prevented reliable
operation.
• Flue gas temperatures below the
acid dew point caused unbalancing
of the booster blower and (ulti-
mately) severe corrosion/erosion
damage to the fan blades.
• The guillotine-type damper installed
to isolate the FGD plant from the
boiler during shutdowns became
inoperable, after binding from the
accumulation of aggregated fly ash
along the tracks.
• Energy supply to drive the evapora-
tor Circulating pump must not be
interrupted. Originally, this drive
was a steam turbine operating on
high pressure steam supplied by
the boiler. During a boiler shut-
down and with no evaporator
circulation, slurry in the evaporator
Table 1. Plant Improvement Projects
Item Date completed
had to be drained and diluted at
considerable penalty in additional
startup time required.
• Sulfur condenser leaks were a
recurring problem throughout the
2-year demonstration.
Test program results and operating
experience during the second year
showed a substantial improvement in
FGD performance; since boiler utilization
was high, the results more nearly
duplicated those expected during com-
mercial application. Flue gas character-
istics were essentially unchanged from
those of the second baseline test. The
conclusions that follow are based on
second-year results.
Dependability
Reliability of the FGD unit (hours
operated/hours called upon to operate)
was 61%. However, FGD plant utiliza-
tion was consistent for the last 7
months: reliability averaged 74% for
that period. Only full operation was
counted as hours operated, full opera-
tion being integrated operation of the
absorber/evaporator loop with the
reduction unit when the flue gas bypass
was closed. The purge treatment unit
Action
Coal supply
Air heater
Duct insulation
Blanks
Completed June 78
During September 78
shutdown
After September 78
shutdown
During September 78
shutdown
Booster blower Not completed
Evaporator pump
Absorber
Booster blower
turbine
Sulfur condenser
During September 78
shutdown
During September 78
shutdown
After September 78
shutdown
During September 78
shutdown
Provide an uninterrupted supply of
Captain coal for Mitchell No. 11 use.
Remove part of baskets which provide
heat storage, to raise inlet duct
temperature.
Insulate duct before and after booster
blower.
Provide way to install blanks rapidly
at inlet of booster fan as an alternative
to the isolation damper.
Install a sparger pipe in the booster
blower to periodically steam clean
blades while running.
Install an electric motor as an
alternative to steam turbine drive.
Recoat and repair leaks.
Provide enclosure to protect against
SOz and weak acid attack.
Plug leaking tubes.
Note: Dates indicated were during first year of demonstration test program,
September 1977-September 1978.
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may or may not have been operating.
The reliability record was established
with virtually no redundancy built into
the FGD unit. Also, the evaporator was
designed for the equivalent of only 80%
of full boiler load. With limited surge
capacity in the regeneration loop, the
FGD plant was unable to effect complete
SOz recovery during evaporator or
reduction unit shutdowns. The 'FGD
plant was utilized (hours operated/hours
in period) 56% of the time. The downtime
was due to:
• FGD repair - 33%.
• Start-up and shutdown - 5%.
• Boiler down - 6%.
The S02 absorber was essentially
trouble free. The reduction unit required
the most downtime for repairs, followed
by the evaporator circulating pump and
the booster blower (Table 2). Called
upon time is defined as the time the
boiler operated to provide flue gas and
utilities and the time other feed streams
were available within the specific
design criteria of the FGD plant. These
include:
• Flue gas at rates not less than 46
MW equivalent.
• Stable steam pressures within an
operable range around a design
pressure of 3,790 kPa (550 psig).
• Electricity.
• Natural gas.
• Soda ash.
• Boiler stable within limits of
greater than 46 gross MWe, and
coal sulfur content greater than 2.8
and less than 3.5 wt. %.
SO2 Removal
The process was controlled to remove
89% of the inlet SO2. Thirty-day average
removal efficiencies varied from 88% to
93% (Figure 2). The 24-hour and 1 -hour
average data for removal efficiency
were:
Table 2. Reasons for Interruption of Operations
SOz
removal
Percent of
time operated
24-h 1-h
average average
90% and greater 60 52
89% and greater 84 78
85% and greater 97 97
On a boiler heat input basis, SO2
emissions were controlled in the range
of 110 to 400 ng/J (0.25 to 0.94 lb/106
Btu).
S02 removal was attained at electrical
generating outputs in the range of 53 to
85 MW of the 115 MW boiler. The lower
Equipment or
reasons
Reduction unit
Evaporator circulating pump
Booster blower
Start-up and shutdown
Other equipment, including absorber
Evaporator
Days of
interruption
59
28
18
17
10
8
% of called
upon time
17
8
5
5
3
2
limit was set by the limiting turndown
capability of the reduction unit. The
upper limit was set by the 80% capacity
limitation of the evaporator, as designed.
Because a substantial amount of energy
(primarily as boiler main steam) was
consumed by the FGD plant, the
generating potential of the boiler was
actually about 95 MW at the FGD
maximum capacity limit of 85 MW.
"Generating potential" refers to the
gross megawatts that the boiler is
capable of generating during FGD
operation but cannot attain due to the
boiler main steam consumed by the
FGD plant.
Confirmation of Design
Limitations
Design limitations were identified in
the Introduction. The effect on FGD
plant performance was:
• Performance at maximum design
flue gas flow rates and SO2 feed
WO
90
o
I
§«>
70
05
3
c
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.G
1
rates could not be adequately
tested because of the design
limitations of the regeneration
system. Flue gas volumes up to
95% of full boiler load were treated
successfully during the acceptance
test. During the demonstration test
program, the maximum SOz feed
rate that was sustainable for 24
hours was 0.750 kg/s (5950 Ib/h),
exceeding design expectations.
Capacity of the recovery unit met
design expectations. The recovery
unit successfully processed 90% of
an SOz feed rate of 0.588 kg/s
(4670 Ib/h) at a flue gas flow rate
of 152 mVs (323,000 acfm). The
boiler generating potential was 89
MW and actual gross electrical
generating output was 80 MW.
This output was sustained for 65
hours during maximum load tests|
Baseline flue gas flows were
confirmed. At a generating poten-
£_ Desigt
Oct Nov Dec Jan Feb Mar Apr May Jun Jul Aug Sep Oct
1978 1979
Period
Figure 2. Thirty-day average SOz removal.
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tial of 92 MW, the design point, the
flue gas flow rate was 170 mVs
(360,000 acfm). This was 13%
higher than design expectations.
Oxygen in the flue gas averaged
8.0% by volume, compared to an
expected oxygen level of 5.6%.
• Installation of an electric drive for
the evaporator circulating pump
reduced start-up time significantly.
• Tests of absorber turndown, with
the reduction unit not operating,
established a minimum throughput
equivalent to a generating poten-
tial of 50 MW. The limiting factor
was a flue gas flow rate of 100
mVs (220,000 acfm), limited by
the minimum governor setting of
the steam turbine drive of the
booster blower. The minimum
sustainable load at full operation of
the FGD plant, with the reduction
unit operating, was a 62 MW
generating potential. Ninety per-
cent S02 removal was achieved for
4 days at an average S02 feed rate
of 0.426 kg/s (3380 Ib/h).
Process Economics
The projected annual operating cost
in 1978 dollars was 13.2 mills/kWh.
Actual annualized cost for the second
year of the demonstration, using 1978
prices for- raw materials and utilities,
was 14.9 mills/kWh. Byproduct sulfur
production averaged 1.81 kg/s (17.3
ton/day). Credits for the sale of the
sulfur amounted to only 0.2 mills/kWh.
Actual annualized cost was based on
the 82.2 MW of generator output that
was possible at the FGD design capacity
of 92 MW, assuming 351 days of boiler
operation per year. On this basis, the
actual boiler capacity in total kilowatt
hours was 92% of the projected capacity.
Annualized costs are quite sensitive to
lower-than-projected capacities because
fixed costs, about 50% of the annual
costs, and labor costs continue to
accrue whether or not the boiler or the
FGD plant is operating at full capacity.
Energy and Raw Material
Consumption
A significant amount of steam pro-
duced by the boiler was consumed by
the FGD plant, used primarily by the
evaporator to recover S02 and regen-
erate the scrubber solution (Figure 3).
Electrical power consumption amounted
to about 1 MW after the evaporator
circulating pump had been converted
from steam-turbine to electrical drive
for improved operability. Actual average
steam consumption was 105% of
design expectations. The energy equiv-
alent of this steam was 11% of the boiler
input energy derived from fuel. Since
the average generating output of the
boiler was 77 MW, the equivalent loss
in electrical generating capacity
amounted to an 8% derating of the
boiler from a nameplate capacity of 115
MW. Including 1 MW of electricity
consumed, the total energy requirement
was 1*2% of the boiler heat input derived
from fuel at an average boiler load of 77
MW.
Soda ash was used as makeup
sodium carbonate for the scrubbing
process. Makeup is made necessary by
the buildup of inactive constituents in
the absorber/evaporator loop, such as
sulfate and thiosulfate, that must be
purged. Any loss from the system due to
leaks also would require soda ash
makeup. High soda ash consumption
during the first demonstration year was
due to leaks at the bottom collector tray
of the absorber that were repaired
before commencing the second demon-
stration year. Average daily consump-
tion of soda ash for the last 7 months of
operation was 0.091 kg/s (8.7 tons/day),
using the total operating days of the
absorber/evaporator as the time base.
Soda ash consumption as a function of
S02 removed was 0.217 mol Na/mol S
removed. The performance guarantee
for acceptance was 0.069 kg/s (6.6
tons/day) at the design levels for flue
gas flow and inlet SOj. Soda ash
consumption of 0.069 kg/s (6.6 tons/
day) is equivalent to 0.152 mol Na/mol
S removed at a design feed rate of S02 of
0.6101 kg/s (4842 Ib/h) and 90%
removal.
Natural gas was used as the reductant
for converting the SO2 to elemental
sulfur. It also was the fuel used to
incinerate the tail gas emitted from the
reduction process. The tail gas was
returned to the inlet of the absorber
after incineration. It was necessary to
continue incinerator operation during
shutdowns to destroy the reduced
sulfur forms that desorb from the
reduction unit refractory materials.
Thus, there was a corresponding
improvement in unit consumption of
natural gas with improvement in reli-
ability. About 0.2 m3 (7 ft3) of natural gas
was consumed per pound of sulfur
produced which was in accordance with
the design expectations. During opera-
tion, the incinerator consumed 7.5% of
the natural gas. In contrast, the inciner-
ator consumed over 12% of the gas
overall because it continued to operate
during shutdowns, demonstrating the
Design
11.5 Ib/lb
Oct Nov Dec Jan Feb Mar Apr May Jun Jul Aug Sep Oct
1978 1979
Period
Figure 3. Comparison of actual steam consumed as a function of SOz feed rate
with design value.
U. S. GOVERNMENT PRINTING OFFICE: 198I/559-092/33!0
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result of a 61% FGD plant reliability
factor
Purge Treatment
Considerations
The purge unit, as initially designed,
was to have treated a small purge
stream removed from the regeneration
loop, to separate sodium sulfate from
most of the sulfite/bisulfite components,
and to dry the sodium sulfate to produce
a marketable product. The "wet" end of
this purge treatment system performed
satisfactorily. The limitation of the
purge dryer already has been men-
tioned.
The amount of purge to be treated is a
function of the formation of sulfate and
possibly thiosulfate. Attempts to deter-
mine the amount of sulfate formation
during absorption were frustrated by an
inability to obtain correct flow measure-
ments and uncertainties about the
specific water balance across the
absorber. However, the data seem to
indicate that sulfate formation is a
function of the oxygen concentration in
the flue gas. Thus, higher than design
purge rates might have been due to the
high excess air levels in the flue gas.
The average purge rate* for the last 7
months of operation has been estimated
to be 18.2% to 24.8%, substantially
higher than expected. The estimate was
determined from soda ash consumption
and the calculated amount of S02
removed. A purge rate of about 10% was
the value indicated during the design
phase of the project. In summary, the
information seems to indicate actual
purge rates much higher than design.
Purge rates of this magnitude put a
further load on the purge solids dryer.
Dryer tests performed by Davy McKee
determined that the dryer did not have
the needed capacity, even at design
rates. The maximum dryer capacity
achieved during the test, approximately
66% of the design heat duty, could not
be sustained because of a buildup of
solids at the discharge end of the dryer.
Maximum capacity that could be sus-
tained without this buildup was only 45-
50% of design.
Davy McKee is investigating the use
of an antioxidant that shows promise for
reducing the oxidation rate and, thus,
the amount of sodium values that must
be purged as sodium sulfate.
Recommendations
Overall performance of the FGD
demonstration unit was affected signifi-
cantly by design limitations that were
known but not eliminated for various
reasons. Typically, capital cost saving
funding limitations, and need for furthi
development are the incentives for le;
conservative design. For installatior
demonstrating new FGD technolog
design criteria should be established .
the start of the program that focus c
the advantages and limitations of th
process rather than having to repo
poor performance solely because <
design limitations. The WL/A demor
stration plant was design-limited by lac
of regeneration capacity and by almos
complete lack of installed spare;
Improved performance would be e>
pected with a more conservative desigr
It is recommended that full or exces
capacity and redundant equipment b
designed into future demonstrations t
the maximum extent possible. Th
demonstration test and evaluatio
would have to indicate the degree c
overdesign and associated costs, if an\
so that installed costs relative t
performance are demonstrated.
•Purge rate was determined as the ratio of moles
sodium consumed to moles SO? removed from flue
gas, expressed as a percentage
R. C. Adams andS. W. Mulligan are with TRW, Inc., Research Triangle Park, NC
27709; R. R. Swanson was with Aide, Richmond, VA 23225.
Norman Kaplan is the EPA Project Officer (see below).
The complete report, entitled "Demonstration of Wellman-Lord/Allied Chemical
FGD Technology: Final Report and Demonstration Test Second Year Results," (
(Order No. PB 81-246 316; Cost: $29.00, subject to change) will be available
only from:
National Technical Information Service
5285 Port Royal Road
Springfield, VA 22161
Telephone: 703-487-4650
The EPA Project Officer can be contacted at:
Industrial Environmental Research Laboratory
U.S. Environmental Protection Agency
Research Triangle Park, NC 27711
United States
Environmental Protection
Agency
Center for Environmental Research
Information
Cincinnati OH 45268
Postage and
Fees Paid
Environmental
Protection
Agency
EPA 335
Official Business
Penalty for Private Use $300
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