-------
S02 emission reductions will be obtained by placing a national cap on total utility
emissions and a "bubble" over each utility's emissions. The US strategy allows flexible,
utility-shaped compliance strategies to be realized. The US allowance market, if
operated without significant intervention, has the potential to lower the overall cost of
emissions compliance to the utility industry by allowing each utility to determine its
most cost-effective compliance approach, while nevertheless constraining emissions from
the utility sector as a whole.

Even given the differences in legislations, the selection in technologies is similar, with
principal reliance on a combination of fuel switching and installation of FGD systems.
Implementation and technology selection will remain highly country-specific.

A review of the US and German availability data suggests that FGD systems have, after
a transitional period, had minimal impact on unit availability. For the US, in 70 percent
of the units evaluated, the overall unit EUF, as the result of an FGD system, which
includes at large both planned and unplanned outages, was less than or equal to 1
percent For Germany, less data is available for analysis; however, it appears that,
following a "learning" period, the influence of FGD equipment on unplanned outage
rates was less that 03 percent, as calculated for 1990. Limited data available from Japan
also shows that FGD has a minimal impact on unit availability.

The impact of design measures will have a considerable impact on the unit production
cost, roughly in the range of 8 to 23 percent for FGD with a farther 8 percent impact for
SCR, the accuracy of such estimates being limited by the varying financial, monetary, and
energy conditions present in each country.

As environmental pressures continue to be felt around the world, legislative, political,
economic factors, and technologies selected will undoubtedly change and evolve but,
given current experience to date, such changes are likely to affect costs, but not
reliability, of supply. Nevertheless, it is important for utilities worldwide as they install
FGD systems to employ appropriate data collection systems necessary to confirm the
role of FGD on plant availability.

References

1.	Salvaderi L., C. Reese, H. Schlenker, and S. StaUard, "Thermal Generating Plant
Availability and the Environment," Joint UNIPEDE/World Energy Council
Committee on the Availability of Thermal Generating Plant, September 1992.

2.	Salvaderi L, S. Stallard, and C. Reese, "Emissions Control: Comparison of the US
and Western European Approaches," presented at Power-Gen Europe '93, Paris,
France, May 1993.

3.	"Impact of FGD Systems, Availability Losses Experienced by Flue Gas
Desulfurization Systems," North American Reliability Council, July 1991.

89-24

-------
DEMONSTRATION OF PROMOTED DRY SORBENT INJECTION
IN A 75 MWe TANGENTIALLY-FIRED UTILITY FURNACE

D. Rodriguez
C. Gomez
Intevep SA
Caracas 1070A, Venezuela

R Payne
P.M.Maly

Energy and Environmental Research Corporation
18 Mason
Irvine, California

Abstract

The injection of dry calcium based sorbent materials into the furnace space is an S02
control technique which has been thoroughly studied and well documented in recent
years. Application of the technique has however been limited by the relatively low
utilization of typical sorbent materials, and by an ability to achieve only moderate levels
of SO2 emission control

In the study described here, a promoted Ca(OH)2 sorbent material has been developed,
in which small quantities of chemical additives are introduced into the water of
hydration. In pilot scale experimental testing, the resulting sorbent was found to yield a
reactivity towards SO2 which was 30 - 50% greater than conventional sorbent materials
under typical utility boiler conditions.

Injection of the promoted sorbent (a process called PromiSOx™) has recently been
demonstrated in a 75 MWe utility boiler firing a 3% sulfur bituminous coal. Sorbent
production was achieved in a commercial prototype hydration unit with a capacity of 10
tons/hour. Under optimum injection conditions, the promoted sorbent showed
significant improvement in performance, and recorded greater than 80% SO2 removal
at an injection rate of Ca/S = 2.6. Results from this demonstration program are
presented and discussed.

Introduction

The direct injection of dry, calcium based sorbent materials into the furnace space
(Furnace Sorbent Injection or FSI) is a well known SO2 control technique, which has
been frequently demonstrated in recent years. The sorbents most typically employed
are limestone (CaC03) or hydrated lime (Ca(OH)2), injected at flue gas temperatures
around 2200°F, and have been reported to achieve SO2 removal rates in the range of 50
- 60%. The process has been offered as a low capital cost alternative to scrubbers for

90-1

-------
some units, usually where title capacity factor is low, and only moderate levels of S02
removal are required. Major limitations of the basic FSI process have included: low
utilization of the sorbent material (typically 25% for Ca(OH2)); an ability to achieve only
moderate levels of S02 control; and difficulties associated with the collection of the
spent sorbent material in existing particulate collection devices; all of which have tended
to restrict the number of commercial installations.

Throughout the history of the development and evaluation of sorbent injection as an
S02 control technique, many studies have been conducted into the fundamentals of
those processes which dictate sorbent behavior, and into the development of new
sorbent materials to enhance reactivity. Such studies are weE documented in the
proceedings of recent 'Joint Symposia on Dry SO2 (and Simultaneous S02/N0x)

Control Technologies' (1986 -1991, published by EPRI), and are represented in papers
by authors such as: Slaughter et aL; Muzio et al.; Nolan et al.; and Gullet et aL. ITie
techniques employed have used the direct admixture of various additive materials with
the sorbent, or the use of additives combined with the water of hydration in the
preparation of enhanced Ca(OH)2. Mostly, these different studies met with moderate
success, and resulted in sorbent materials with small or no changes in reactivity towards
^2, particularly under conditions typical of utility boiler systems.

In the present study, a similar approach to sorbent enhancement has been adopted,
where small amounts of chemical promoters have been added to quicklime during the
hydration process, but in such a manner that sorbent reactivity has been significantly
improved. The development of this new approach has taken place over a three year
period. It was initiated with fundamental studies into sorbent sulfation, which lead to
the identification of specific promoter chemicals and processing techniques, followed by
studies in both bench- and pilot-scale facilities. Results from the experimental testing
provided sufficient confidence to proceed with the construction of a commercial
prototype hydrator with a capacity of 10 tons of promoted sorbent per hour. This
prototype unit allowed the program to proceed to a full-scale demonstration in a 75
MWe utility boiler firing a 3% sulfur bituminous coal.

Process Description

The promoted dry sorbent injection process will be marketed under the trade name of
PromiSOx™, and is a relatively simple variation on the well known Furnace Sorbent
Injection (FSI) process, where a dry calcium based sorbent material (limestone or
hydrated lime) is injected into the upper region of a furnace firing medium to high
sulfur fuels. A typical utility bote arrangement of this process is presented in Figure 1.

With PromiSOx™, a promoted sorbent material is used, which is based on a calcific
hydrate that is prepared in a patented process in which chemical additives are
introduced into the quicklime during the hydration process. A specially designed
hydrator has been developed for this purpose. The sorbent may be produced either in a
central hydration plant and shipped in bulk containers to the plant site, or it may be
manufactured on-site from quicklime in a dedicated hydration facility.

As with FSI, the promoted sorbent is injected at a location where flue gas temperatures

90-2

-------
are on the order of 2000°F. Under these conditions, the injected sorbent decomposes
in-situ to form CaO which subsequently reacts with SO2 to form a solid, dry CaS04. The
dry product, which consists of CaS04 and unreacted CaO, mixed with the fuel ash,
passes through the boiler to downstream particulate collection equipment, where it is
removed.

Hie sorbent is typically injected into the boiler through an array of wall mounted
nozzles arranged in the upper furnace region of the boiler, where six to eight
penetrations of no more than 4 inches (100 mm) in diameter are usually sufficient to
ensure uniform distribution of the sorbent throughout the furnace space. Injection of
sorbent from two elevations in the boiler may be necessary to ensure optimum S02
removal across the boiler load range. Sorbent feeding is usually from a small day-silo,
by means of a gravimetric feeding system, with pneumatic transport to the injection
locations. Bulk storage of the sorbent is provided in larger silos with a capacity sufficient
for several days' supply.

The application erf any furnace sorbent injection process can result in increased solids
deposition in the convective passes of the boiler, and in degradation of the performance
of existing electrostatic precipitators (ESPs). Available experimental data, and limited FSI
applications conducted to date, suggest that convective pass fouling should not present
a significant problem, and its occurrence can be managed by modifications to standard
sootblowing techniques. ESP performance may however be affected as a consequence
of the higher solids loading, as well as by the modified electrical properties. Mild
humidification of the flue gases, or conditioning by SC)3 injection, have been
demonstrated to be effective in restoring collection efficiency.

Experimental Studies

The promoted sorbent concept was developed initially as a result of fundamental
studies into the mechanisms of S02 reaction with calcium, oxide. These studies provided
insight into the reaction mechanisms, leading to the possibility of modifying the sorbent
particle morphology for enhanced reactivity, both in the initial 50 msec of reaction time
(the so-called 'prompt capture regime', where utilization occurs very rapidly), and at
longer times where utilization is limited by product layer diffusion. This lead in turn to
the identification of certain chemical additives, and to the development of processing
techniques for the production of highly reactive sorbent materials. Experimental studies
were carried out on: sorbent production; sorbent reactivity characterization; and
process evaluation under representative operating conditions.

Initially, sorbent production was carried out in small laboratory scale facilities where
CaO was hydrated under carefully controlled experimental conditions. Parameters
evaluated included: CaO source; additive combinations and quantities;
CaO/water/additive mixing conditions; and reaction time. The promoted sorbent
materials were subsequently characterized for reactivity in a small laboratory
drop-tube furnace, in order to define the optimum ranges of process conditions. The
manufacturing process was extended to a continuous hydration facility developed
specifically for this purpose, and producing 200-500 lb/hr. The production of larger
quantities of material allowed reactivity testing to be conducted in larger bench-scale

90-3

-------
facilities, and ultimately in a fully integrated pilot plant, more closely simulating the
conditions, geometry and auxiliary equipment of typical utility boiler applications.

Most recently, sorbent production has been successfully scaled to a 10 ton/hour
commercial prototype hydration facility.

Reactivity testing of the promoted sorbent has shown that its most important
characteristic is an ability to generate a CaO with, a high in-situ porosity and surface
area, and to retain these features through the temperature window most favorable to
the sulfation reactions. This is different from sorbent materials which begun with a high
initial surface area, but where this potential reactivity is lost to high temperature
processes, such as sintering, before significant reaction with SO2 can occur. In the bench
scale furnace facilities used in the initial phase erf this project, special sampling and
analysis techniques were developed to provide information on the morphology of the
sorbent particles, in-situ, where the sulfation reactions occur. Typical results are shown
in Figure 2, where surface area and porosity are compared for a promoted sorbent
material and a conventional calcific hydrate. Hie superiority of the promoted sorbent is
clearly evident from these data, as is its ability to maintain enhanced reactivity at
temperatures important to the sorbent sulfation process.

Prior to proceeding to a full-scale demonstration of the promoted sorbent process,
considerable proof-of-concept testing was carried out in a 10 x 106 Btu/hr pilot scale
combustion test facility. The facility was designed to simulate conditions of temperature
and time characteristic of utility boiler operation, and provided representative heat
transfer surfaces and downstream equipment for the evaluation of ash deposition and
particulate collection. Testing has been carried out with medium and high sulfur
bituminous coals, and with Orimulsion®, an emulsified natural bitumen from
Venezuela.

Testing has shown that the optimum location for the injection of the promoted sorbent
is slightly lower than that normally employed in PSI, and occurs where the flue gas
temperature is approximately 2000°F (1100°C). Under optimum conditions, the
promoted sorbent showed a reactivity which was between 30% and 50% greater than
that normally found with commercially available calcific hydrates. Experimental data
for the 10 x 106 Btu/hr pilot scale furnace are presented in Figures 4 and 5, for a 3.6%
sulfur bituminous coal, and for Orimulsion® fuel respectively. Hie significant
improvement in SO2 removal performance can be readily seen from these figures, as
can the tendency for the promoted sorbent to retain its reactivity at relatively high
overall S02 removal rates. SQ2 removal in excess of 80% was achieved with sorbent
addition rates corresponding to Ca/S molar ratios on the order of 2.5.

In addition to parametric testing for S02 removal performance, the promoted sorbent
has been subjected to several long term teste to evaluate potential impacts on furnace
ash deposition and heat transfer performance, and on the collectability of the fly ash in
downstream particulate collection devices. The results of the pilot-scale testing showed
that the promoted sorbent behaves in a manner consistent with that observed in boilers
and other experimental facilities operating with conventional FSL Compared to
operation without sorbent injection, there was an increase in ash deposition on
convective pass surfaces, although the ash deposits were found to remain friable and
were readily amenable to removal by sootblowing. The rates of ash deposition were

90-4

-------
comparable to those for conventional sorbent materials at similar mass addition rates.
Similar conclusions were also drawn with respect to collection of the fly ash particulate
in the facility baghouse and electrostatic precipitator.

Demonstration Testing

The host site for the promoted sorbent demonstration program was Illinois Power's
Hennepin Station Unit 1, which is a 71 MWe (net) tangentially-fired boiler burning a 3%
(as received) sulfur Illinois bituminous coal. In 1990, this unit was retrofitted with
systems for Gas Reburning (for NOx control) and Sorbent Injection (for S02 control), as
part of a demonstration program under Round 1 of the DOE Clean Coal Technology
program. Since that time, the unit has been operated to provide both parametric and
long-term evaluation erf the combined NOx and S02 control technologies, under the
joint sponsorship erf the U S Department of Energy (DOE), the Gas Research Institute
(GRI) and the State of Illinois Department of Energy and Natural Resources (ENR). A
general arrangement of the Hennepin unit is presented in Figure 5, showing the gas
reburn and sorbent injection systems, together with the flue gas humidification system
which provided conditioning in order to maintain performance of the ESP in the
presence of spent sorbent

Testing of the promoted sorbent was carried out over a nominal three-week period in
December of 1992. The promoted sorbent was produced in the 10 ton/hour prototype
hydration facility, trucked to the plant site in bulk carriers, and then off-loaded into the
main sorbent storage silo. The sorbent feeding and injection system developed for the
GR-SI demonstration program was retained un-modified for the promoted sorbent
testing. The feeding system employs a Fuller-Kinyon pump, which provides for dense-
phase transport of the sorbent in compressed air to the injection elevation, at which
point a booster fan provides additional air to support injection. The injection system
itself comprises an array of six three-inch diameter ports, located at the boiler nose-arch
elevation where flue gas temperatures are nominally in the range of 2250 - 2300°F. This
injection elevation was known to be non-optimum for the promoted sorbents, which
are known to prefer somewhat lower temperatures. A significant benefit to this
program however was that a considerable data base on sorbent injection and associated
boiler impacts had been developed during the course of the GR-SI demonstration. Since
the promoted sorbent was expected to behave similarly to the conventional,
commercially available, sorbents, this meant that the program could concentrate on
parametric testing.

During the course of the demonstration program, the effects of a large number of
operating variables were evaluated in order to identify optimum conditions for
promoted sorbent injection. These included parameters such as: burner tilt, excess
oxygen, boiler load, sorbent injection velocity, and whether gas reburning was
implemented or not Results from the testing are presented in Figure 6 for nominally
optimum injection conditions. The figure shows percentage S02 removal as a function
of the injected Ca/S molar ratio, where both parameters are related to the initial actual
S02 level to account for the reduced S02 concentration when gas reburning is
implemented. As can be seen from Figure 6, the promoted sorbent results in a
significant improvement in both sorbent utilization and S02 capture, relative to the
unpromoted material which was a commercial calcific hydrate from Linwood. At a

90-5

-------
sorbent injection rate of Ca/S = 2.6, more than 80% S02 removal was recorded,
resulting in a final SO2 concentration on the order of 500 ppm (at 3% O2).

Parameters which were found to most significantly influence sorbent utilization were
those impacting temperature at the injection elevation. Reduced burner tilt, increased
excess air, and reduced boiler load were all found to improve sorbent utilization to
some extent. The general impact of temperature on utilization is shown on Figure 7,
where temperatures have been estimated from available measurement data and from
boiler heat transfer modelling. Figure 7 suggests that the optimum injection
temperature for the promoted sorbent is on the order of 2050°F, at which point the
measured sorbent utilization ranges from 40% to about 70% higher than the normal,
unpromoted calcitic hydrate.

In addition to measurements of SO2 removal, a number of additional measurements
and observations were made concerning performance parameters such as sorbent
feeding, ash deposition in the boiler convective passes, and ESF performance. Such
observations were generally made relative to the available experience base with
conventional sorbent injection, and these are briefly summarized below.

In terms of sorbent feeding, no significant problems were encountered, and the sorbent
was found to flow freely through all stages of the feeding and injection system. Only
tihe maximum capacity of the system was somewhat reduced due to a slightly lower
bulk density of the promoted sorbent.

Sorbent injection into the furnace can add significantly to the particulate loading in the
boiler convective passes such that tube surface fouling and reduced heat transfer is of
potential concern. In general, although there is an increase in the rate of solids
deposition, experience has shown that the deposits remain friable and easily removed
by sootblowing. During the demonstration testing, heat balances were routinely
performed across individual convective tube banks in order to monitor fouling and to
calculate cleanliness factors. Results from such evaluations are presented in Figure 8 for
the secondary and primary superheaters and for the reheater sections, and for both
promoted and unpromoted sorbents. It can be seen from the rates of reduction in the
cleanliness factors that deposition is comparable for the two different sorbents, and that
application of the sootblowers tends to recover the initial starting values.

As part of the GR-SI demonstration program a humidification system was installed on
the Hennepin unit as a means of conditioning the flue gases in order to compensate for
the impacts of the sorbent modified fly ash cm ESP performance. The humidification
system comprises an array of twin-fluid atomizers which inject water into the flue gas
duct at a location just downstream of the air heater. Approximately 2 seconds of
residence time is available prior to the entrance to the ESP. The normal mode of
operation with sorbent injection is to add sufficient water to maintain the stack opacity
level comfortably below 20%, at which point measurements have shown the total solids
loading to be well below the permitted 0.10 lb /MMBtu. For both the promoted and
unpromoted sorbents the necessary humidification rate was found to be nominally the
same, and equivalent to that necessary to reduce the flue gas temperature to
approximately 250°F.

90-6

-------
Process Economics

Economic evaluations have indicated that promoted sorbent injection can be a cost
competitive S02 control technique for a wide range of applications, including utility and
industrial boilers, and incinerators. The process requires relatively small modifications
to the existing boiler system, and since additional equipment requirements are modest,
tiie initial capital investment necessary is low. The process is therefore ideally suited to
retrofit applications, and can be applied with most medium to high sulfur fuels (coal, oil,
Orimulsion®, etc) As with FSI, it may be most readily applied in coal fired units, or
plants designed for coal, where there is an existing particulate control device.

Economic analyses relative to wet-FGD and FSI have been carried out using the EPA
Integrated Air Pollution Control System (IAPCS) cost estimating model, and using data
obtained in the demonstration program. The results of selected economic comparisons
are presented in Figure 8 for a typical 100 MWe boiler application firing a 2.6% sulfur
coal, and where the retrofit difficulty for a wet-FGD system would be considered
moderate. The figure compares initial capital and operating costs (on a relative basis)
for wet-FGD, ESI, and PromiSOx™. The significant potential advantages of the
promoted sorbent process, in terms of both capital and operating costs, can be seen for
this particular application. Particularly attractive applications of the process are likely to
be found in smaller boilers (<300 MWe), where moderate levels of S02 control are
required, and in situations where the high capital costs of flue gas scrubbing cannot be
justified.

Acknowledgements

PromiSOx™ is a trade mark of Intevep SA (an affiliate of Petr61eos de Venezuela). The
authors wish to thank Intevep SA for their authorization to publish this paper.

The authors also wish to thank the staff at Hennepin Station for their cooperation and
assistance, and to acknowledge the US Department of Energy (DOE), the Gas Research
Institute (GRI), and the State of Illinois Department of Energy and Natural Resources
(ENR) for making the site and equipment available under the Clean Coal Technology
Program.

90-7

-------
Figure 1. Conceptual sorbent injection system.

1200	1400	1600

Injection Temperature (K)

0.4-

0.35-

«

o
»-
o

a

0.3-

0.25-

0J2-

Promoted
Sorbent

%

Unpromoted V
Sorbent	&

1200	1400	1600

Injection Temperature (K)

figure 2. Comparison of in-situ surface area and porosity data.

• 90-8

-------
Figure 3. SO2 removal performance - Coal. Figure 4. SO2 removal performance - Orimulsion®.

Sorbent Injection System

Transport Air Blower

Figure 5. Overview of the GR-SI and Humidification systems at Hennepin.

90-9

-------


a,

88

u

100
90
80
70
60
50 E-

« 40

o

m 30
20
10
0

PromiSOx

o

:x'

D • / •

^ ™ TTnnrrwnni

••

Unpromoted

V

/

/

A

111 • 1111111111



i.

.0 0.5 1.0 1.5 2.0 2.5 3.0
Ca/S

Figure 6. Effects of Ca/S on sulfure capture.

1800

PromiSOx

Unprompted

2000	2200

Injection Temperature (°F)

2400

Figure 7. Effects of temperature on utilization.

90-10

-------
— Unpromoted I,in wood : Ca/S=1.75
	 PromiSOx: Ca/S=1.60

Hennepin Unit 1

Secondary Superheater
Sootblowing

Rch eater

' 		



Primary Superheater

50

100
Minutes

150

200

Figure 8.

Impacts of sorbent injection and sootblowing on convective
pass cleanliness factors.

HI Capital Costs
(3 Operating Costs

90% DeSOx

50% DeSOx

80% DeSOx

Wet FGD	FSI	PromiSOx™

Figure 9. Economic comparison.

-------
DEVELOPMENT OF PROCESS TO SIMULTANEOUSLY SCRUB N02 AND S02

FROM COAL-FIRED FLUE GAS

V.M. Zamansky, R.K. Lyon, A.B. Evans, J.N. Pont, W.R. Seeker
Energy and Environmental Research
18 Mason, Irvine, CA 92718

C.E. Schmidt

Department of Energy, Pittsburgh Energy Technology Center
P.O. Box 10940, Pittsburgh, PA 15236

Abstract

Energy and Environmental Research Corporation (EER) has conducted extensive
research under contract to DOE to further develop and test a new process capable of
simultaneously scrubbing SO2 and NO2. This process involves injection of methanol
into combustion flue gas to promote the oxidation of NO to NO2. Nitrogen dioxide can
subsequently be removed, along with the sulfur dioxide, in a liquor-modified wet-
limestone scrubber. The focus of this study was to determine a liquor composition and
a scrubbing method capable of removing both compounds effectively. This paper
presents the results of kinetic modeling, and bench and pilot scale experiments of the
scrubbing study.

Introduction

The CombiNQx process is being developed to provide a low cost method of controlling
the NOx emissions of coal-fired utility boilers to very low levels. This process
incorporates a family of NOx reduction technologies including staged combustion,
Advanced Rebuming (i.e. reburning combined with selective non-catalytic reduction)
and methanol injection to convert NO to NO2 which can then be removed by wet
scrubbing. While individually, these technologies are limited in their NOx reducing
capabilities, in combination they are capable of reducing NOx emissions to extremely
low levels at a fraction of the cost of selective catalytic reduction.

The methanol injection step, however, is subject to the limitation that one must have a
scrubber that will remove NO2. Hie removal of SO2 and NO2 by sodium based wet
scrubbers is a well-established technology, but the majority of wet scrubbers currently
in use are calcium based. Accordingly, this study was undertaken to determine
whether or not the chemistry which occurs in calcium based scrubbers could be
modified to allow removal of NO2 as well as SO2.

Bench-scale experiments were performed in conjunction with chemical computational
modeling to evaluate the effect of scrubbing solution composition on SO2 and NO2
scrubbing efficiency. In addition, potential byproducts were identified. Finally, larger

91-1

-------
pilot scale tests were performed with a packed tower scrubber to address scale-up
issues and confirm the bench-scale results.

Bench scale studies

The bench scale NO2 scrubbing apparatus is displayed in Figure 1. A simulated flue gas
containing variable amounts of NO2 and SO2 was flowed through a constant
temperature, 80 cubic centimeter, bubbler containing the scrubbing solution to be
evaluated. After passing through the "scrubber", the gas was analyzed for O2, NOx,
N20, and SO2 to determine removal efficiencies.

Figure 2 graphically summarizes the bench scale results. The numbers symbolize
various cases that were performed while varying slurry composition, the lines show the
resulting performance as a function of time. A Ca(OH)2 solution (2 percent Ca(OH>2 by
weight) achieved 99+ percent SO2 reduction, indicating that the scrubber provides good
mass transfer. However, only 50 percent of the NO2 was removed. Reference 1
indicates that the crucial reaction for NO2 scrubbing is:

N02 + SO32- = N02- + SO3-	(1)

With pure Ca(OH)2 solution, the necessary sulfite ion (SO32-) tends to precipitate out as
calcium sulfite, instead of reacting, as desired, with NO2-

Since sodium sulfite (Na2$03) can provide the necessary sulfite ions for N02 absorption,
a 2 percent solution was evaluated. 99+ percent SO2 and 89 percent NO2 capture was
obtained initially, however, this performance decreased after approximately 2 minutes
of run time. After these 2 minutes, SO 2 reduction became negative and N02 reduction
dropped to 17%. SO2 is captured via the reaction

SO2 + SO32- + H20 = 2HSO3-	(2)

and NO2 via reaction (1). Unfortunately, the SO3- generated by NO2 removal is a chain
carrier in the oxidation of sulfite and bisulfite ions to sulfate and bisulfate ions.
Oxidation of the bisulfite ion to the bisulfate ion acidifies the solution, forcing SO2 back
into the gas phase.

Adding Ca(OH)2 to the Na2S03 scrubbing solution eliminates the problem of SO2
rejection by keeping the solution basic, however the NO2 capture remains poor and
short lived. Replacing the highly soluble Ca(OH)2 with very low solubility CaCOs
increases the concentration of sulfite ion that stays in solution. This improves N02
absorption, but the sulfite ion quickly oxidizes to sulfate, hampering N02 removal.
Experiments were conducted under conditions which inhibited the oxidation of sulfite
to sulfate (Le. decreasing reaction temperature and/or flue gas oxygen content). Even
though these conditions can not be applied to a real application, they did show that if
sulfite ion stays in solution, N02 capture improves and can be sustained for a longer
period of time.

91-2

-------
Sodium thiosulfate (Na2S20g) has been proposed by Gorin et al.2 as a method of
inhibiting the oxidation of sulfite to sulfate. We found that by adding this compound to
the scrubbing solution (1% by weight), 95% NO2 capture was achieved.and this capture
was sustained for the same length of time that the solution captured SO2.

The resulting recommended slurry solution, capable of scrubbing 99+ percent SO2 and
95 percent NO2 at bench scale level is:

9% Na2C03
9% CaCOs
1% Na2S203.

This scrubbing mixture was studied using computer modeling with the mechanism
described by Chang et al.l For the batch experiments, the reaction was found to
proceed in three stages. During the first stage, carbonate ion concentration falls while
the concentrations of sulfate, bisulfate and sulfite ion increase. This first stage ends
when the ratio of carbonate to sulfite ion becomes so low that calcium carbonate starts
to dissolve while calcium sulfite precipitates. During this second stage the ratio of the
concentrations of carbonate and sulfite ion are constant and the pH remains fairly
steady. The concentration of sulfate ion increases until calcium sulfate starts to
precipitate. The second stage ends when the calcium carbonate is exhausted. When this
happens the pH begins to rise until it reaches a level at which SO2 absorption fails and
scrubbing solution is spent.

Hie model predicts that NO2 capture will remain effective through these three stages of
the process, but the fate of absorbed N02 changes. During the first and second stages
nearly all of the absorbed N02 will be present as nitrite ion, with only trace amounts of
the complex nitrogen sulfur ions being formed. During the third stage, however, the
nitrite ion is converted to complex nitrogen sulfur ions, chiefly the aminetrisulfonate
ion. Near the end of the third stage the aminetrisulfonate ions are hydrolyzed to
sulfate ion and sulfamic acid.

While the model contains reactions which are capable of forming N20 and nitrate ions,
these reactions are only significant at very low pfi During the bench scale experiments
samples were taken of the bubbler's exhaust and, consistent with the models
predictions, no significant N20 production was observed. Measurements were also
made with EM Quant test strips on spent scrubbing solution. As one would expect
from the model, nitrite ion was found in fresh solution but not in solution which had
been allowed to age. Nitrate ions were not detected in the spent solution.

Pilot scale studies

Scale-up effects were investigated in two different pilot-scale facilities corresponding
nominally to heat inputs of 0.35 MBtu/hr and 10 MBtu/hr. The small pilot-scale
scrubber tests were performed by Research-Cottrell using the facility illustrated in

Figure 3. The small pilot-scale scrubber consists of a propane combustor, absorber
tower, absorber feed tank, analytical train, and solid disposal system. To simulate a
coal-fired flue gas, variable amounts of S02 and N02 were doped into the exhaust

91-3

-------
upstream of the absorber tower. At the absorber tower exit, NOw SO2, CO, and O2
were measured. The absorber tower is a vertical, stainless steel, 16 inch diameter tube,
approximately 20 feet in height. Hie simulated flue gas enters the . tower from the
bottom, travels through five sections to the top, and exits to the gas sample
conditioning systems and analyzers. The first two sections can be packed with a light
packing material to provide improved gas/liquid contact. Sections of the tower may
also be removed, if desired, to reduce absorber tower residence time. The scrubber
slurry is continually being mixed with dry limestone, sodium salts, and water in the
absorber feed tank. From the 200-gallon feed tank, the slurry is pumped to the top of
the absorber tower and dispensed in counter flow to the flue gas with a single slurry
nozzle. The slurry solution is drained by gravity from the bottom of the tower back to
the feed tank.

The first tests were performed to verify that SO2 removal was possible on the pilot scale

scrubber. For a 6 percent limestone slurry, flue gas flow rates were varied between 127
- 140 cfm, and slurry flow rates were maintained at 12 gpm. Up to 99 percent SO2
removal was obtained indicating satisfactory mass transfer.

Simultaneous scrubbing of NO2 and SO 2 was evaluated using a mixture of solid
scrubbing salts consisting of 49.5 percent CaCOs, 49.5 percent Na2C03, 1 percent
Na2&zQ3. Note that this is approximately 1/5 of the concentration of Na2S203 utilized
in the bench-scale tests. The mixture of salts or pure limestone were fed to the
absorber tank with different feed rates. During this test series, the following
parameters were varied: liquid/gas ratio (liquid flow and gas flow were independently
varied), concentration of sodium carbonate in slurry, concentration of sodium
thiosulfate in slurry, and initial NO2 concentration. NO2 removal efficiency ranged
between 65 and 90 percent while maintaining 97 - 99 percent SO2 removal.

The ratio of slurry flow rate to flue gas flow rate is defined as the liquid to gas ratio
(L/G), and is expressed here in unite of (gallons of slurry)/(1000 cubic feet of gas). Hie
slurry flow arid gas flow were varied independently of one another. Figure 4
summarizes the effects of L/G ratio on NO2 scrubbing efficiency. As would be
expected, a larger L/G ratio results in higher NO2 removal. Even though not depicted
in the figure, data indicate that gas flow rate has a larger affect on the NO2 scrubbing
efficiency than slurry flow rate. By decreasing the gas flow rate by a small fraction
(from 135 to 115 cfm), efficiency increased from 77 to 84 percent However, when the
slurry flow rate was nearly doubled (11.4 to 20 gpm), the efficiency only increased by
approximately the same amount, 77 to 85 percent.

Experiments were performed to determine the effect of initial NO2 concentration on
NO2 scrubbing efficiency. Slurry flow rate remained approximately constant as inlet
N02 concentration was varied by adjusting the doping gas flow rate. Results are
displayed in Figure 5. The general trend shows that scrubbing effectiveness drops as
initial NO2 concentrations increase.

The effect of Na2C03 concentration on NO2 and SO2 scrubbing efficiency was evaluated
by diluting the scrubbing solution by a factor of two, while continuing to add limestone
to maintain pH and ion concentration. Figure 6 shows that even though scrubbing

91-4

-------
efficiency was initially hampered by the dilution that occurs 200 minutes into the test,
with time the NO2 removal efficiency rose again to approximately the same level as
before the dilution. Even after diluting the slurry a second time, the NO2 scrubbing
efficiency returned to almost the original value These data indicate that NO2 removal
efficiency is not very sensitive to Na2CC% concentration in this range.

The effect of sodium thiosulfate concentration was tested by adding an additional 3.8
mmol of sodium thiosulfate per liter of solution. Figure 7 shows that NO2 removal
efficiency jumped from 65 percent to 89-90 percent within 15 minutes, while SO2
removal efficiency remained at 99+ percent The thiosulfate inhibits oxidation of sulfite
to sulfate, sustaining the presence of sufficient sulfite ions for NO2 capture. Summary
of bench-scale and pilot-scale scrubbing results is shown in Figure 8. From 90 to 95%
NO2 reduction was achieved at 99+% SO2 removal.

Throughout the experiments discussed above, scrubbing solution composition
measurements were periodically taken. In general, these observations were consistent
with the model and the assumption that the scrubber was operating in the second stage
(see previous discussion). The model is, however, limited in its ability to account for
sulfite ion oxidation, therefore, not surprisingly, sulfite to sulfate conversion was much
higher than predicted. The pilot scale studies also show nitrate ion as a major product,
contradicting both model's prediction and the bench scale results.

A single test was performed in lER's large pilot-scale facility. The large pilot-scale
scrubber facility consists of a simple spray tower with a pad-type demister. The spray
tower is 6 ft in diameter and 16 ft high, with an array of 16 spray nozzles. The
scrubbing solution consisted of an original composition of 9 percent CaC03,9 percent
Na2CC>3 and 1 percent ^28203. This slurry was circulated through the scrubber while
the furnace was firing a high sulfur coal for three consecutive days of testing. This
circulation allowed accumulation of sulfites within the slurry which is necessary to
achieve NO2 removal. The actual scrubbing experiment took place using natural gas
fired flue gas doped with 74 ppm of NO2- SO2 was not added. The test was conducted
at a L/G ratio of approximately 30 gal/103acf. The test result is shown in Figure 4.
Much higher NO2 removal was achieved than in the small pilot-scale tests due to the
much higher concentration of sodium thiosulfate used in the large pilot-scale test. The
results do indicate that high NO2 removal efficiencies can be achieved even with a
relatively primitive scrubbing system operating at L/G ratios similar to the of large
commercial scrubbers.

Discussion

Hie primary goal of this research, demonstration of efficient NO2 and SO2 scrubbing in
a calcium based wet limestone scrubber, has been achieved. Two important questions,
however, remain with respect to the disposability of the products produced by this
modified scrubber. First, there is the question of whether or not using sodium
compounds in the scrubbing solution will result in unacceptable sodium contamination
of the calcium- sulfate/sulfite product Since wet scrubbers require considerable
amounts of makeup water, it is theoretically possible to solve this problem by washing

91-5

-------
the calcium sulfate/sulfite with the makeup water, but this option would require an
engineering design study which has not been done.

The second question of disposability regards the formation of nitrate ion seen in the
pilot scale experiments. This ion formation was not detected during bench scale or
computer modeling studies. One possible explanation for this discrepancy is that as the
scrubbing liquid passes downward through the absorber tower, it readies a point at
which its ability to absorb SO2 is completely exhausted. This low pH condition was not
considered in the modeling study, yet it may responsible for the increase in nitrate
concentration. Mechanisms within the model do, in fact, show nitrate ion formation at
low pH levels.

While possible formation of nitrate ion will require further research, there are a number
of ways in which this problem might be solved. Adjustment of scrubbing conditions so
that the solution is prevented from over-reacting may prevent nitrate formation.
Alternatively, if nitrate forms by oxidation of nitrite ion, the removal of nitrite ion by
reaction with NH2SO3H may prevent nitrate formation, or, all else failing, nitrate ion
could be removed by selective reduction with scrap aluminums. In a brief study, the
authors found that this method works quite well with shredded soda cans as the source
of aluminum.

Acknowledgments

This work was funded under DOE Contract No. DE-AC22-90PC90363, "Development of
Advanced NOx Control Concepts for Coal-Fired Utility Boilers". EER would like to
acknowledge Mr. Charles E. Schmidt, Department of Energy Program Coordinator of
the CombiNOx program. The small pilot-scale experimental scrubbing studies were
conducted by Research-Cottrell under subcontract to EER.

References

1.	S.C. Chang, D. Littlejohn and N.H. Lin, Flue Gas Desulfitrization, ACS Symposium

Series 188, American Chemical Society, p. 128 -152, (1982).

2.	E. Gorin, MD. Kulik, R.T. Struck, U.S. Patent 3,937,788 (1976).

3.	A.P. Murphy, Chemical Removal of Nitrate from Water, Nature 350,223-225 (1991).

91-6

-------
Figure 1. Experimental set-up for NO2/SO2 scrubbing experiments.

91-7

-------
1	2% Ca(OH)2

2	2^%Na2»33

3	2.2% NajSOg, 9% Ca(OH)2

4	2.5% NagSO,, 10% CaCOg

5	9% Na^C03,9% CaCOa

6	9% N^COg, 9% CaC03,1% N^S203

_L

X

X



-10 0 10 20 30 40 50 60 70 80 SO 100 110 120 130 140 150 160 170 180 190 200

Time (min)

Figure 2. Bench scale NC^ scrubbing results.

91-8

-------
Demister

• Outlet Gas





Sampling

Datms

1

/l\

ooo
ooo
ooo
ooo
ooo
ooo
ooo
ooo
ooo
ooo
ooo
ooo

ooo
ooo
ooo
ooo
ooo
ooo
ooo
ooo
ooo
ooo
ooo
ooo
ooo
ooo
ooo

AT

"L

TC

no2-»no
Converter

Analytical
Train

Gas
Conditioner



NCVNO*

CO

o,

NO2—- NO
Converter

	TC

Air

I

1

OR

-L

H2°

I

1
1
J

I
I
I

~

* _

Absorber
Feed Tank

a r

TC

—<^FeederJ

Pump

o-

AT -Abso iter Tower
MFM - Magnetic Flow Meter
M - Manometer
N - Nozzle

TC-Thermocouple
R- Rotameter
P- Packing Materia!
OR - 2.5" Orifice

Figure 3. Schematic of the pilot-scale scrubber.

91-9

-------
90

80

o

m
>

70

©
cc

CM

o
z

60

50

~ 2 MMBtu/hr Pilot Scale Tests

O 10 MMBtu/hr Pilot Scale Tests
j	j	!	

20

40

60

80	100

->3,

120

140

160

L/G, apm/1(r acfm
Figure 4. Effect of liquid to gas ratio on NC^ removal.

90

80

to
>

70

®
cc

CM

O

60

50

50

100	150

[N02J, ppm

200

250

Figure 5. Effect of initial N0_ concentration, on NO^ removal

91-10

-------
100

Slurry diluted by
1/2.

200	300	400

Time (min)

600



100



90



80

5*

70





O



c



o

60

a

lD



?

50

£!



s

40

&



o~

30

2





20



10



0

Figure 6. Effect of N^COg dilution on N02 scrubbing performance.

UG ratio = 83 gal/1000 cf

Addition of 3.8
mmol/L of
Thiosulfate

10	20	30	40

Time (min)

50

60

Figure 7. Effect of thiosulfate concentration on N02 scrubbing performance.

91-11

-------
100-/

80-*-

N02 Removal

S02 Removal

Slurry Comoostion:
9% Limestone
9% Sodium Carbonate
0.5-1% Sodium Thiosulfate

Bench-Scale

Pilot-Scale

Figure 8. Bench-scale vs. pilot-scale scrubbing results.

91-12

-------
The NOXSO Clean Coal Technology Project:
Commercial Plant Design

James B. Black
Mark C. Woods
John J. Friedrich
Clay A. Leonard

NOXSO Corporation
P.O. Box 469
Library, Pennsylvania 15129

Abstract

The NOXSO process is a dry, post-combustion flue gas treatment technology which
uses a regenerable sorbent to simultaneously adsorb sulfur dioxide (SOj) and nitrogen oxides
(NOj) from die flue gas of a coal-fired utility boiler. In the process, the SOj is converted to
a sulfur by-product and the NOx is reduced to nitrogen and oxygen. Based on pilot plant
results, the process can economically remove 90% of the acid rain precursor gases from the
flue gas stream in a retrofit or new facility. The NOXSO Clean Coal Technology Project
will demonstrate the NOXSO process on a commercial-scale. The project is co-funded by
the U.S. Department of Energy (DOE) under round III of the Clean Coal Technology
program. The DOE manages the project through the Pittsburgh Energy Technology Center
(PETC). The NOXSO process, the plant layout, and economics for a commercial-scale plant
are described in this paper.

92-1

-------
Introduction

The NOXSO process is a dry, post-combustion flue gas treatment technology which
uses a regenerable sorbent to simultaneously adsorb sulfur dioxide (S(X) and nitrogen oxides
(NO,) from the flue gas of a coal-fired utility boiler. In the process, the SQ> is converted to
a sulfur by-product (elemental sulfur, sulfuric acid, or liquid SOj) and the NO* is reduced to
nitrogen and oxygen. Based on pilot plant results,the process can economically remove 90%
of the acid rain precursor gases from the flue gas stream in a retrofit or new facility.

Process development began in 1979 starting with laboratory scale tests and
progressing to pre-pilot scale tests (3/4-MW) and a life cycle test. Each of these test
programs have provided data necessary for the process design. Tests of the NO, recycle
concept which is inherent to the NOXSO process have been conducted on small boilers at
PETC and the Babcock & Wilcox (B&W) Research Center in Alliance, Ohio.4 A 5 MW
Proof-of-Concept (POC) pilot plant test at Ohio Edison's Toronto Plant in Toronto, Ohio was
recently completed.5 The Clean Coal Project is currently in the project definition phase
incorporating recently obtained pilot plant data into a commercial-scale design.

The commercial demonstration of the NOXSO process will be cost shared by the
Department of Energy through the third round of the Clean Coal Technology program
through a cooperative agreement between DOE and NOXSO. The cooperative agreement is
currently being assigned (novated) to NOXSO by Morrison Knudsen Corporation - Ferguson
Group and will be formally executed shortly. NOXSO will provide overall management for
the project while MK-Ferguson will provide engineering and construction services. NOXSO
will conduct the operation phase of the project. W.R. Grace & Co.-Conn will be the sorbent
supplier. DOE will manage the demonstration project through the Pittsburgh Energy
Technology Center (PETC).

NOXSO Process Description

Flue gas is drawn from the power plant duct work either upstream or downstream of
the particulate collection device through the flue gas booster fan. Figure 1 shows a process
flow diagram with flue gas drawn from the particulate collection device discharge. Figure 1
shows single pieces of equipment, however multiples will be used as required to provide the
necessary capacity. Tail gas from the sulfur by-product plant is mixed with the flue gas at
the booster fan suction. The flue gas then passes through a two-stage, fluidized bed adsorber
where S02 and NOx are simultaneously removed using a high surface area ^-alumina sorbent
impregnated with an alkali material. Water sprays into the fluid beds maintain a 250°F
temperature by evaporative cooling. The cleaned flue gas passes through a particulate
separator and is returned to the power plant before it exits through the chimney. Sorbent
fines removed by the separator are directed to the dense phase transport system.

Sorbent is removed from the adsorber and transported to the sorbent heater by a dense
phase pneumatic conveying system. Make-up sorbent to maintain the sorbent inventory is
added downstream of the adsorber. The sorbent heater is a variable area five-stage fluidized
bed where a hot air stream is used to raise the sorbent temperature to 1150°F. During the

92-2

-------
Figure 1. Process Flow Diagram

-------
heating process, NOx and loosely bound S02 are desorbed and transported away in the
heating gas (NOx recycle) stream. This hot air stream at SOOT can be used to heat a slip
stream of the power plant's main condensate before being injected into the combustion air
system upstream of the combustion air preheater. Hie NO% recycle stream provides
approximately 30% of the required combustion air. Upon entering the boiler, a portion of
the recycled NO, is converted to nitrogen (Nj) and either carbon dioxide (CO2) or water
(H20) by reaction with free radicals in the reducing atmosphere of the combustion chamber.
NOx recycle studies were performed during a previous NOXSO test program (a 3/4 MW pre-
pilot scale test). More recently, NOx recycle studies were conducted using a scaled model
cyclone boiler.4

Once the sorbent reaches a regeneration temperature of 1150°F, it is transported by
means of a J-valve to the moving bed regenerator. In the regenerator, sorbent is contacted
with natural gas in a countercurrent maimer. The natural gas reduces sulfur compounds on
the sorbent (mainly sodium sulfate) to primarily SOj and hydrogen sulfide (H2S) with some
earbonyl sulfide (COS) also formed. Some of the sodium sulfate (N%S04) is reduced to
sodium sulfide (Na2S) which is subsequently hydrolyzed in a moving bed steam treatment
reactor which follows the regenerator. A concentrated stream of H2S is obtained from the
reaction of steam with Na2S. The offgases from the regenerator and steam treater are
combined and sent to a sulfur by-product plant which produces elemental sulfur, sulfuric
acid, or liquid S02. The tail gas stream from the sulfur by-product plant is recycled to the
suction of the flue gas booster fan.

From the steam treatment vessel, the sorbent is transported by means of a J-valve to
the sort>ent cooler. The cooler is a five-stage variable area ftoidized bed using ambient air to
cool the sorbent. The warm air exiting the cooler is further heated by a natural gas fired in-
duct heater before being used to heat the sorbent in the fluidized bed sorbent heater. The
sorbent temperature is reduced in the sorbent cooler to the adsorber temperature of 250°F.
Sorbent from the sorbent cooler is transported by means of a J-valve to a surge tank located
above the adsorber. The surge tank is used as a source and sink for sorbent to maintain
constant bed levels in the other process vessels. From the surge tank, sorbent flow to the
adsorber is regulated using an L-valve, thus completing one full cycle.

General Arrangement

Figure 2 shows a general arrangement for a nominal 100 MW NOXSO plant. The
major components will be identified by tracing the flow paths of the flue gas, the
heater/cooler gas, and the sorbent. This arrangement shows two adsorber trains. Flue gas
enters the lower of the two large ducts in the foreground, splits and flows through the flue
gas booster fans, adsorbers, and particulate separators before recombining and exiting the
NOXSO tower.

Ambient air for cooling the sorbent enters through two of three 50% capacity
heater/cooler fans. The air is preheated by the sorbent in the sorbent cooler (hidden behind
right adsorber) before flowing through the air heater (located below the sorbent heater) where
it is heated with natural gas. The high temperature air enters the bottom of the tapered

92-4

-------
Adsorber(2

NOx recycle t
power plant-

Flue gas
outlet

Flue gas
inlet

Sorbent heater

Jiegenerafcoi

Steam
Treater

Sorbent
Cooler

Flue gas booster
fan (2)

Heater/cooler
fans (3-)

Figure 2. NOXSO Process Tower

* 3^ *

92-5

-------
sorbent heater and exits from the top. This exit gas is the NO, recycle stream which goes to
the combustion air system of the power plant.

Sorbent is transported from the adsorbers to the sorbent heater. After being heated in
the sorbent heater, the sorbent is transported to the moving bed regenerator and then to the
steam treater. From the steam treater, the sorbent flows to the sorbent cooler where it is
cooled before being transported back to the adsorber completing the cycle.

Pilot Plant Test Program

NOXSO pilot plant tests on a 5 MW slip stream from Ohio Edison's Toronto power
plant in Toronto, Ohio have recently been completed. Results of these tests are reported in a
paper by Haslbeck, et al.s Highlights of the test program are discussed below.

After the startup aid shakedown phases of the test program were completed, a
parametric test phase was begun. During the period of time from April through December
of 1992 while the parametric tests were being conducted, the pilot plant availability averaged
76%. This high availability for a pilot plant undergoing frequent operating changes is a
testimony to the simplicity and reliability of the process.

SO2 and NOx removal efficiencies were quantified as a function of flue gas flow rate,
solids circulation rate, adsorber solid inventory, adsorber temperature, inlet S02
concentration, and inlet NO, concentration. A modification to the original project scope
included installation of a second bed in the adsorber and adding in-bed water sprays to lower
the adsorber temperature. With these modification, S02 removals of 95% and NO* removals
of 90% were easily attained at economical operating conditions.

Sorbent attrition was determined by measuring the initial and final sorbent inventory
and the quantity of sorbent added to the system to maintain the sorbent inventory. After
3,232 hours of testing the attrition rate was 3.0 lb/hr. This represents 0.011% of the total
sorbent inventory per hour which is less than the 0.016% lb/hr previously estimated.

Economics

Data from the pilot plant have been incorporated into the design of a commercial size
NOXSO plant. Using this commercial plant design, an economic analysis was performed for
the NOXSO process. The basis for the analysis and cost information are included in Table
1. The analysis was conducted for a 500 MW power plant burning 3% sulfur coal and
emitting 0.6 lb NOx/MMBtu.

Since the NOXSO process is a combined S02/NO, removal process, it is not possible
to separate the cost of removing S02 from the cost of removing NO,. Consequently, an
assumption is made that the cost of removing NOx is 3.0 times higher than the cost of
removing SOz. The value of 3.0 represents a reasonable average for the relationship between
die cost of NOx and SO2 removal based on published economic studies of separate high

92-6

-------
Table 1.

NOXSO PROCESS ECONOMIC ANALYSIS (1)

POWER PLANT PARAMETERS

GROSS CAPACITY
CAPACITY FACTOR
HEAT RATE
COAL HEATING VALUE
COAL SULFUR
NOz EMISSIONS

500 MW

mo %

10,000 Btu/kWh
12,000 Btu/lb
3.0 %

0.6 tb/MMBm

ECONOMIC PARAMETERS

ELECTRICITY	S0.03 /kWh

NATURAL GAS	%USQMsd

SORBENT	$3.40/lb

NET SULFUR VALUE	$50 Aon

S02 ALLOWANCE VALUE	J30Q

FIXED CHARGE RATE (2)	10.6 %

REMOVAL COSTNOx/REMOVAL COSTS02	3.0

NOXSO PROCESS REMOVAL EFFICIENCIES

SOZ	95%

NO*	80%

EMISSIONS DATA

UNCONTROLLED S02
CONTROLLED S02
PHASE IS02 LIMIT

S02 ALLOWANCES GENERATED

UNCONTROLLED NOx
CONTROLLED NOx

POLLUTANT REMOVAL EFFICIENCY

76.550 toasfyear
3,833 tons/year
38325 umsfytat
34,493 tons/year

9.158 tons/year
1,840 tons/year

93,4 %

OPERATING AND MAINTENANCE COSTS

FIXED (3)
VARIABLE (4)
NATURAL GAS
SORBENT
ELECTRICITY
TOTAL

$5,714,000
$129,030
$5,131,000
$10 J 12,000
13,642,000
$24,728,000

CAPITAL COST

REVENUES

S12&5OO.OO0

$257 flcW

S02 ALLOWANCES
SULFUR VALUE
TOTAL

$10347,750
$1320,438
$12468.188

NET LEVEUZED COST

S26J80,813/year

8.5 railk/kWh
S276/ton~S02
$828/ton—NOx

(1)	1993 dollars.

(2)	Based on 30 year book life. 20 year tax life, 38% composite federal and state tax.
and 2.0% for property taxes and insurance.

(3)	Includes operating labor. fringes, and supervision; maintenance labor and equipment:
and general and administrative expenses.

(4)	Includes process water and Clans plant catalyst.

92-7

-------
efficiency technologies. This value does not affect the overall economics, however it does
affect the relative cost of S02 and NOx removal.

Emissions date are also listed in Table 1. The "Phase I S02 Limit" is calculated
based on allowable emissions of 2.5 lb S02/MMBtu. It is appropriate to consider over
compliance since the high removal efficiency of the NOXSO process will alow a utility to
generate S02 allowances which can be sold to partially offset the operating cost. A value of
$300 has been assumed for S02 allowances. Beginning in the year 2000, the number of
allowances generated will decrease, however it is also likely that the value of allowances will
be significantly higher offsetting to some degree the reduction in the number of allowances
generated.

The annual operating and maintenance cost is $24.7 million with the cost of sorbent at
$10.1 million representing 41% of the total. The capital cost of $257/kw is based on a
recent EPRI study.6

Revenues for the process will be generated by the sale of the sulfur by-product and
the S02 allowances. The sulfur by-product can be elemental sulfur, sulfuric acid, or liquid
S02. The choice of sulfur by-product will be influenced significantly by the local demand
for the specific product. Since the market for sulfur is larger than the other two, sulfur is
used in this analysis. If a local market exists for sulfuric acid or liquid S02, either would be
a more economical choice since the revenue from sulfuric acid would be approximately three
times more than sulfur and liquid S02 would be six to eight times more. Making sulfuric
acid or liquid S02 would also result in minor increases in capital and operating costs.

The net levelized cost for the process is presented from three points of view. The
cost of buying, operating, and maintaining the plant will be $26.2 million dollars per year.
This translates to 8.5 mills/kwh of electricity produced. On a pollutant removal basis, it cost
$276 to remove each ton of S02 and $828 to remove each ton of NOx.

Acknowledgement

NOXSO Corporation wishes to thank the following for their contributions to the
project: U.S. Department of Energy's Pittsburgh Energy Technology Center, Ohio Coal
Development Office, W.R. Grace Co-Conn., and Moirison-Knudsen Corporation - Ferguson
Group.

References

1.	Haslbeck, J.L., C.J. Wang, L.G. Neal, H.P. Tseng, and J.D. Tucker, "Evaluation of
the NOXSO Combined NO^/SO;, Flue Gas Treatment Process", U.S. Department of
Energy Contract NO. DE-AC22-FE60148, November 1984.

2.	Haslbeck, I.L., W.T. Ma, and L.G. Neal, "A Pilot-Scale Test of the NOXSO Flue
Gas Treatment Process", U.S. Department of Energy Contract No. DE-FC22-
85PC81503, June 1988.

92-8

-------
3.	Ma, W.T., J.L. Haslbeck, and L.G. Neal, "Life Cycle Test of the NOXSO Process",
U.S. Department of Energy Contract NO. DE-FC22-85PC81503, May 1990.

4.	Zhou, Q., J.L. Haslbeck, L.G. Neal, "An Experimental Study of NOx Recycle in the
NOXSO Flue Gas Cleanup Process", U.S. Department of Energy Contract No. DE-
AC22-91C91337, March 1993.

5.	Haslbeck, J.L., M.C. Woods, W.T. Ma, S.M. Harkins, J.B. Black, "NOXSO
SO^/NO* Flue Gas Treatment Process: Proof-of-Concept Test", presented at the S02
Symposium, Boston, MA (August 1993).

6.	J.E. Cichanowicz, C.E. Dene and W. DePriest, et al., "Engineering Evaluation of
Combined N0j/S02 Controls for Utility Application", Electric Power Research
Institute, Palo Alto, CA, (December 1991)

92-9

-------
INVESTIGATION OF SORBENT REGENERATION KINETICS
IN THE COPPER OXIDE PROCESS

Joanna M. Markussen
Henry W. Pennline
U.S. Department of Energy
Pittsburgh Energy Technology Center
P.O. Box 10940
Pittsburgh, PA 15236

Introduction

The Copper Oxide Process is a dry, regenerable process using a copper (Cu)-
impregnated alumina sorbent for the simultaneous removal of sulfur dioxide (S02)
and nitrogen oxides (NOx) from flue gas. Testing erf the Huidized-Bed Copper Oxide
Process has ranged from laboratory experiments to process-developmental-scale
integrated operation1"4. The Moving-Bed Copper Oxide Process is scheduled for
future testing in the life-Cycle Test System at the Pittsburgh Energy Technology
Center (PETC)5. The fluidized-bed and moving-bed processes use different gas-solid
contacting designs for absorption, but both processes use a moving-bed regenerator
for sorbent regeneration. Most of the kinetic information in the literature has focused
on the absorption process. Kinetic data dealing with the effect of changing gas
composition on the regeneration step (as would be found in a moving-bed
regenerator) are very limited. Such information is needed for a better understanding
of the regeneration step and for potential improvements in the design of the moving-
bed regenerator used in both copper oxide process configurations.

In the Copper Oxide Process, copper oxide (CuO) is converted to copper sulfate
(CuS04) by reaction with S02 and oxygen (O2) in the absorber at approximately
400°C.

CuO * S02 + 1/2 Oz - CuS04	(1)

Catalytic reduction of NOx to nitrogen (N2) is achieved with the injection of ammonia
(NH3) into the absorber. Both CuS04 and CuO act as catalysts in the reduction
reaction.

93-1

-------
4 NO + 4 NH2 + 02 - 4 Nz + 6 if20

(2)

2 N0Z * 4 iW3 + 02 ¦* 3 2\^ + 6 HzO

(3)

The sulfated sorbent is then regenerated at 450-500°C with methane (CH4) in a
moving-bed reactor, where the reducing gas flows countercurrent to the sorbent The
overall reaction is

Note that as the sulfated sorbent is reduced, the gas composition along the length of
the reactor changes due to the release of SOy C02, and H20. For every mole of CH4
consumed, five moles of gas are produced.

The regenerated sorbent is then cooled and returned to the absorber, where the Cu is
rapidly oxidized to CuO by 02 in the flue gas.

It is essential to get nearly complete regeneration of the spent sorbent so that 95-99%
S02 removal can be achieved in the absorber with only a slight excess in the Cu/S
stoichiometric ratio.

The Fluidized-Bed Copper Oxide Process was tested in an integrated mode in the
PETC Life-Cycle Test Unit using a fluidized-bed absorber and a continuous
moving-bed regenerator3'4. The regenerator had a solids residence time of 1-2 hours,
depending upon the solids flow rate, and the average regenerator temperature was
approximately 450°C. Only 20-50 percent of the CuS04 fed to the regenerator was
reduced, and this incomplete regeneration resulted in a decrease in the S02 removal
in the absorber. To get 90% S02 removal, solid circulation rates of 4-6 times the
theoretical minimum had to be used. In a commercial process, high solid circulation
rates would increase the operating costs through higher sorbent attrition losses and
higher process energy costs. Thus, for future process development, there is a strong
incentive to improve the regeneration efficiency.

The kinetics of S02 removal have been studied using fixed beds and different size
fluidized beds, but there is little information on the kinetics of regeneration. Early
tests by McCrea et al.6 in a packed-bed reactor showed slow regeneration with CH4 at
400°C, but iwarly complete sulfur removal was obtained in 30 minutes at 450°C or in
5 minutes at 500°C The sorbent retained about 1% sulfur even after 1 hour exposure
to CH4 at 500°C, but the sulfur was believed to be bound to the alumina support and
did not affect the capacity of the CuO for subsequent SO, removal. Tests by Yeh et
al.2 using a microbalance reactor showed 80% reduction of CuS04 with CH4 in 10
minutes at 450°C or in 25 minutes at 400°C. These rates of reaction were higher than

CuSOt + 1/2 CHi - Cu * SOz + Kp + l/2 COz

(4)

Cu + 1/2 Oz - CuO

(5)

93-2

-------
those reported by McCrea, probably because a larger excess of CH4 was used. Also,
since McCrea and coworkers used a 4-in.-deep sorbent bed, any inhibition of
regeneration by the product gases may have affected the observed reaction rate.

Harriott8 suggested that the incomplete reduction in the moving-bed regenerator may
have been due to low excess methane concentration, high product concentrations, or
a combination of these factors. In a commercial operation, a slight excess erf methane
will be used, such as 10% excess CH4. However, even when 10% excess CH4 is fed to
the regenerator, the exit gas concentration is reduced to 2% due to dilution by the
product gases (see equation 4). The sorbent near the gas exit of the regenerator is
actually exposed to a higher concentration of S02 than CH4. Thus, to aid in
regenerator scaleup and design studies, kinetic data are needed over a wide range of
gas concentrations.

As a result, a series of tests was conducted at PETC using a microbalance reactor
exploring the kinetics of regeneration9. Samples from a batch of sulfated sorbent
were regenerated at 450°, 480°, and 510°C using mixtures of CH^ SO^, COj, and N2.
Tests showed that the regeneration reaction is strongly inhibited by S02 in the
product gases. It is also postulated that unsulfated CuO entering the regenerator is
first sulfated before being reduced to Cu with CH4. The reactions of copper sulfate
and sulfite with CH4 do not go to completion when S02 is present, but not enough
data is available to dearly define the equilibrium limits. More laboratory tests are
needed to develop kinetic correlations that can be used to help interpret pilot-plant
data and determine the best conditions for sorbent regeneration.

Experimental

A series of regeneration tests is currently being carried out in a Cahn Electrobalance
(Model 1000) at PETC. A schematic of the microbalance system is shown in Figure 1.
Gases are blended from certified cylinders to generate the desired regeneration gas
mixtures. Methane concentrations range from 10-100%. During regeneration tests
with CH^SO-j-Nj mixtures, S02 concentration is varied from 10-40%. Total gas flow
rate to the microbalance reactor tube are maintained at about 2000 cm3/min (STP).
A series of regeneration tests in the microbalance at varying gas flow rates from 1000
to 3200 cmVmin (STP) showed that operation at 2000 em'/min results in negligible
gas mass-transfer resistance in the system. The temperature of the reactor is
measured with a thermocouple placed 0.5 cm below the sample basket Regeneration
kinetics are being studied at reactor temperatures of 450°C, 480°C, and 510°C.

The copper oxide sorbent, fabricated by Universal Oil Products, Inc. (UOP SOX-3), is
l/16-in.-diameter spheres of gamma alumina impregnated with 6.8 wt% Cu. A large
batch (15.0 g) of the sulfated sorbent being used for microbalance testing was
prepared in a packed-bed reactor unit. The CuO sorbent was exposed to two cycles
of S02 sorption and CH4 regeneration followed by a final sulfation step. Sulfation
was carried out at 400°C, with a simulated flue gas mixture for the final sulfation step

93-3

-------
of 3990 ppm SCX 2.72% Oy and a balance of N2. Regeneration was performed at
450°C using pure CH4. The total gas flow rate through the packed-bed reactor was
1600 cm3/min (STP).

In the microbalance regeneration tests, samples of sulfated sorbent weighing
approximately 50 mg are spread over a layer of quartz wool placed in a quartz
sample basket The basket is positioned inside the quartz reactor tube in the furnace
and suspended on a wire connected to the microbalance. The sample is then heated
to the reaction temperature using pure N2. The gas stream enters the microbalance at
tiie base of the reactor tube and passes through a section containing ceramic beads,
which are put in the tube to reduce the amount of dead gas volume. Steady-state
temperatures are reached in about 1-1.5 hours.

After the heat-up period, a regeneration gas mixture is introduced into the reactor by
switching the air-actuated four-way valves shown in Figure 1. The flow rates of the
two gas streams are matched to minimize shock buoyancy forces that can cause
inaccuracies in the sample weight measurements during the initial time period. In a
typical CH4-S02-N2 regeneration test, the sulfated sorbent is exposed to a CH4-S02-N2
gas mixture until no further change in weight is observed, which typically occurs
within 60 minutes. After the regeneration has reached completion, the S02 gas is shut
off and the N2 gas flow is increased. The sample is then exposed to a CH4-N2 gas
mixture or pure CH4 to determine if the sorbent can be further reduced. The reaction
system typically reaches steady-state conditions within an additional 30 minutes.
Upon completion of the regeneration, the sample is cooled to room temperature using
N2 gas. From the microbalance weight loss curves, fractional conversion data as a
function of time are obtained.

Discussion

A study of the kinetics erf sorbent regeneration occurring in the Copper Oxide Process
is being conducted using a laboratory microbalance unit This study is designed to
complement and expand upon the work of Harriott and Markussen9,10. The ultimate
goal of the research is to provide a greater understanding of the kinetics of sorbent
regeneration, such as the combined effect of temperature and gas composition on the
reduction reaction, in order to facilitate improvements in the design of the moving-
bed regenerator in the Copper Oxide Process.

The first series of tests in the microbalance unit involves the regeneration of fully
sulfated copper oxide sorbent at varying temperatures and gas compositions. Tests at
temperatures of 450°, 480°, and 510°C are being performed with CH4-S02-N2
compositions containing up to 40% S02. The tests with 40% S02 approximate the gas
concentration expected near the gas exit of a counterflow regenerator fed with a
mixture of CuO and CuS04 sorbent and a small excess of CH4. Previous laboratory
tests were mainly conducted at 450°C or 480°C9. However, because of the strong

93-4

-------
negative effects of SO, on the regeneration rate, thereby limiting conversion, some
additional tests at higher temperatures are being explored in this test program.

A second series of microbaknee tests is being conducted to measure the reaction
rates and limiting conversions for unsulfated sorbent, which is representative of
sorbent that is not sulfated in the absorber. The solid will be exposed to mixtures of
gases similar to those predicted at the gas exit of a counterflow moving-bed
regenerator. These tests will help determine if CuO is first sulfated prior to reduction
with CH4. If the sulfation reaction occurs before regeneration, the sorbent is expected
to gain weight rapidly as SO, combines with the solid and then slowly lose weight as
the reaction with CH4 takes place. Analytical determination of the sulfite species on
the solid will be attempted using X-ray diffraction, scanning electron microscopy, and
X-ray photoelectron spectroscopy. Sensitivity modeling studies of the process show
that the possible sulfation of CuO in the regenerator is potentially important to the
overall process performance and economics".

The laboratory tests with fully sulfated and unsulfated sorbent may not completely
simulate conditions in the moving-bed regenerator, because particles fed to the
regenerator may be partially sulfated and interactions between the solid phases
present may be possible. Therefore, a few regeneration tests with sorbent that is 50
to 70% sulfated are planned. Even though the individual reaction rates cannot be
determined, the overall weight loss can be compared with that predicted from the
kinetics of the separate reactions.

Most of the earlier work9 concentrated on the inhibition effect of SO^. Since C02 and
H20 are products of the regeneration reaction, COz and HzO are also expected to
have some effect on the regeneration conversion. A few tests conducted with CQ,
gave a slightly lower final conversion than when CO, was not present9. More tests

with C02 in the gas stream will be performed to better quantify its effect on the
regeneration reaction rate.

The possible inhibition effect of H20 on the regeneration reaction has not yet been
explored. To study this effect, tests will be conducted in a packed-bed reactor unit
designed for gases containing relatively high moisture content. A number of CuS04
regeneration tests using a thin sorbent bed are planned with 10-15% HjO in the feed
gas along with a slight excess of CH4. By following the S02 concentration in the
packed-bed unit, the amount of CuS04 reduced to Cu can be determined. After the
sorbent has reached steady state with the gas mixture, the sorbent will be exposed to
pure CH4 to determine the amount of sorbent not reduced by the mixture.

Combining the experimental data from this study with the results obtained in
previous microbalance tests will allow better correlation of the reduction kinetics
occurring in the moving-bed regenerator. Equations correlating the kinetic data will
be useful in interpreting pilot-plant data and will be instrumental in determining the
best conditions for sorbent regeneration.

93-5

-------
Summary

A study of the kinetics of sorbent regeneration occurring in the Copper Oxide Process
is being conducted using a laboratory microbalance unit at FETC. Regeneration tests
at varying reaction temperature and gas composition are being performed with fully
sulfated, partially sulfated, and unsulfated copper oxide sorbent. The laboratory tests
will provide data needed to develop kinetic correlations that will aid in the design of
the moving-bed regenerator, will help determine the best conditions for sorbent
regeneration, and will help interpret pilot-plant data.

References

1.	Demski, R.J., S.J. Gasior, E.R. Bauer, Jr., J.T. Yeh, JJP. Strakey, and J.I. Joubert,
"Simultaneous Removal of S02 and NOx from Hue Gas Employing a Fluidized-
Bed Copper Oxide Process," presented at the American Institute of Chemical
Engineers Summer National Meeting, Cleveland, OH (August 29-Sept. 1,1982).

2.	Yeh, J.T., R.J. Demski, J.P. Strakey, and JX Joubert "Combined SOz/NOx
Removal from Hue Gas." Environmental Progress. Vol. 4, No. 4, p. 223 (1985).

3.	Plantz, A.R., C.J. Drummond, S.W. Hedges, and FJST. Gromicko, "Performance

of the Huidized-Bed Copper Oxide Process in an Integrated Test Facility,"
presented at the 79th Annual Air Pollution Control Association Meeting,
Minneapolis, MN (June 22-27,1986).

4.	Williamson, R.R., J.A. Moriri, and T.L. LaCosse, "Phase I - Sorbent Life-Cycle
Testing Huidized-Bed Copper Oxide Process," UOP Project Report to C.J.
Drummond, Pittsburgh Energy Technology Center, Pittsburgh, PA, 1988. DOE
Contract DE-AC22-85PC81004.

5.	Yeh, J.T., J.S. Hoffman, and H.W. PennJine, "Design of a Moving-Bed Copper
Oxide Process for Simultaneous S02 and NOx Removal," Preprint 93-TA-
173.06P, presented at the 86th Annual Air and Waste Management Association
Meeting, Denver, CO (June 13-18,1993).

6.	McCrea, D.H., A.J. Forney, and J.G. Myers, "Recovery of Sulfur from Hue
Gases Using a Copper Oxide Absorbent." JAPCA. Vol. 20, p. 819 (1970).

7.	Yeh, J.T., J.P. Strakey, and J.I. Joubert, "SOz Absorption and Regeneration
Kinetics Employing Supported Copper Oxide," Unpublished Paper, Pittsburgh
Energy Technology Center, Pittsburgh, PA, 1987.

8.	Harriott, P., "Kinetics of Sorbent Regeneration in the Copper Oxide Process for
Hue Gas Cleanup," Report to C.J. Drummond, Pittsburgh Energy Technology
Center, Pittsburgh, PA, August 1989.

93-6

-------
9.	Harriott, P., and J.M. Markussen, "Kinetics of Sorbent Regeneration in the
Copper Oxide Process for Hue Gas Cleanup." Ind. Eng. Chem. Res. Vol. 31, p.
373 (1992).

10.	Harriott, P., "Studies of Sorbent Regeneration in the Copper Oxide Process for
SO2 Removal/' Report to CJ. Drummond, Pittsburgh Energy Technology
Center, Pittsburgh, PA, April 1990.

11.	Frey, H.C., "Performance Model of the Ruidized Bed Copper Oxide Process for
S02/NOx Control," Preprint 93-WA-79.01, presented at the 86th Annual Air
and Waste Management Association Meeting, Denver, CO (June 13-18,1993).

93-7

-------
M1CR0BALANCE

N,

Air

2-c^h^hxi— 3

-an

REGENERATION GASES

iXb

Sample
Basket
Thermocouple
Furnace

Ceramic Beads

Reactor
Tube

a



^ Regulating Valve

Micro Valve

Pressure Control
Valve

Four-Way Valve

Filter

Pleasure
Indicator

?

Balance
Mechanism

Nt Purge

» Vent

Exit
Qas

Water Trap

Indicator

Figure 1, Schematic of the mlcrobalance unit.

-------
Removal of Particulate Matters and Air Toxics

together with S02
in Chiyoda Thoroughbred 121 FGD Systems

Hiroshi Yanagioka
Minora U chid a

Chiyoda Corporation
2-12-1, Tsurumi-chuou, Tsurumi-ku, Y okohama, 230 Japan

Wet FGD systems remove particulate matters and air toxics together with
S02 from flue gas. Intimate contact between gas and liquid is prerequisite
for such ancillary capabilities. The Jet Bubbling mechanism is suited for
efficient gas/liquid contact and the use of the CT-121 FGD systems is
more advantageous for the total emission control of air pollutants than just
switching to low-sulfur fuel.

1. INTRODUCTION

In Japan, the use of sulfur-containing heavy petroleum fuel during the period of rapid
economic growth in the late 1960s caused serious acidic-gas pollution problems. Efforts
were made to develop flue gas desulfurization(FGD) processes, which have been widely
applied for flue gases generated from heavy oil, coal and even Orimulsion*.

It has become known that the FGD systems are also very effective for removing any
impurities in flue gas to provide more cleaner air in industrial areas than expected. This is
because of simultaneous removal of sub-micron particulate matters and other impurities

such as Cf, F", and Hg compounds.

Although electrostatic precipitators (ESP) are in general equipped upstream of FGD
systems for particulate matters removal, there are some industrial boilers without an ESP
or with an ESP de-energized for power saving.

1. Chiyoda Thoroughbred 121 FGD Systems

Fig. 1 illustrates the Chiyoda Thoroughbred 121 process, which is characterized by size
compactness and operation reliability. The Jet Bubbling Reactor(JBR), the heart of the
process effectively integrates all chemical reactions involving removal of S02 and

94-1

-------
formation of gypsum. Internal liquid circulatiorf which is induced by gas bubbling to
absorbent liquid through sparger pipes is utilized to eliminate external slurry circulation
using large pumps.

Fig. 2 shows a comparison of gas /liquid contacting devices in typical wet FGD
scrubbers. The JBR has a higher density of liquid phase even without external liquid
circulation by pumps. In this apparatus, the height of the absorption section involving gas
bubbling or the jet bubbling layer ranges only 1* to 3'. This low height of absorption
results in low energy consumption which is only about 1 % of the power output of a
power plant

The reliability of the system has been increased by suitable operation conditions and
careful design tailored to individual plant requirements. The absorbent having a low pH of
4 to 5 is advantageous for oxidation of sulfite to sulfate by injected air as well
neutralization by limestone with virtually 100 % utilization. Agitation by air bubbling and a
mechanical mixer maintains adequate amounts of suspended gypsum crystals in the lower
part of the JBR to allow crystals to grow to the desired size of about 100 micron. Efficient
washing systems are provided to remove gypsum deposits if any in such zones as a mist
eliminator through which flue gas passes.

The main body of JBR can be fabricated using 317 L type stainless steel to withhold
corrosion because of a thin protective passive film formed on the metal surface in an
oxidative environment. Flake glass reinforced plastic,whether solid or lined, is also
acceptable for this purpose.

A prescrubber, which is usually necessary when particulate matters are detrimental to
gypsum as a salable product, can be omitted if the amount of soot is small or the by-
product gypsum is left stacked at the site. A single length of ductwork with spraying
gypsum slurry from the lower part of JBR works for particulate removal as well as gas
cooling and humidification.

Table 1 shows a list of Chiyoda Thoroughbred 121 systems constructed to date in
different countries such as U.S.A.,Germany, Malaysia, Australia in addition to Japan,
The size of the system started from 23 MW and has recently reached an equivalent
capacity of 1,000 MW.

3. Particulate Removal Efficiency

When flue gas is contacted with scrubbing liquor, particulate matters collide with the
liquor by inertia and diffusion and are subsequently separated from the main stream of
gas. Thus, removal efficiency of particulate matters can be expressed as a function of their
diameters and properties.

94-2

-------
When particulate matters are composed mainly of solid spherical fine panicles that are
insoluble in water as encountered in coal-burned flue gas, collection efficiency is
expressed in terms of particle diameters only. Fig. 3 shows one of the collection
efficiency curves obtained from CT-121FGD systems. This curve indicates that a
collection efficiency of 90 % or more is expected for coal-fired flue gas where particle
diameters are generally greater than 1 micron meter. This has been confirmed in several
installations where outlet dust loadings are below 10 mg/Nm3 with an ESP upstream of
the FGD system.

On the other hand, when particulate matters are soluble in water or scrubbing liquor,
collection efficiency is largely dependent on their chemical properties rather than their
physical sizes or diameters. Particulate matters from oil-fired flue gas contain fine
panicles of ammonium sulfate((NH4)2$04, for example, which dissolves quickly into the
scrubbing liquor, resulting in virtually 100 % removal of this substance at the outlet of the
FGD system. However, carbon dust, which is rather sticky and water-repellent, tends
to slip from the contacting surface between gas and liquid, resulting in less collection
efficiency than expected basal on the particle diameter. Because of such differences in
chemical and physical properties, heavy oil or re-run oil burned flue gas shows collection
efficiencies ranging 2 to 43 mg/Nm from plants to plants.

This holds true in Orimulsion-fired flue gas, where 70% or more of the particulate matters
are composed of such chemically active substances as VOS04 and MgS 04,though finer
particles less than 0.6 micron meters are present. Thus, actual collection efficiency will be
much greater than predicted by the data obtained from coal-fired flue gas.

4. Mercury Removal Efficiency

Various coals contain toxic heavy metals such as Mn, Cr, Ni. As, Hg, and Be depending
on their mining sites though their amounts are considered negligible at present. Literatures
indicate that Hg is present in coal at concentrations of 0.02 to 0.09 ppm by weight.

Hg, when released from coal in boilers at high temperatures, can take various forms such
as Hg vapor, HgCl2, HgO and Hg2Cl2. These compounds significantly differ in
solubilitiy in water. While HgCl2 dissolves in water up to a concentration of 74,000 mg/1,
the solubility of Hg2Cl2 is only 2.1 mg/1 at a temperature of 250 C, indicating that Kg (I) is
difficult to be trapped even in wet FGD systems.

In a coal-fired boiler equipped with a CT-121 FGD plant, analyses were made to
determine whereabouts of Hg and the following results were obtained: 2 % in bottom ash,
12 % in fly ash, 3 % in gypsum, 41 % in waste water and balance in gas to the stack The
data suggest that Hg compounds generated by burning coal are relatively volatile and the

94-3

-------
large portion is water-soluble.

5. Conclusion

In addition to the intended SO2. the Chiyoda Thoroughbred 121 FGD system can
effectively remove particulate matters. Coal-derived particulate matters in particular
provide consistent removal efficiency as a function of particle diameter.

Particulate matters contained in heavy oil -fired flue gas are different from coal-derived fly
ash that can be assumed to be solid, insoluble spheres with a distribution of diameter.
Surface properties and solubility to the scrubbing liquor also play a role in removal
efficiency.

Hg, one of air toxics was found to be removed in the FGD systems, but more
investigations should be made. This added performance of wet FGD systems is certainly
more advantageous than switching to low-sulfur fuel in view point of total air quality.

Table 1 HISTORY OF CHIYQDA'S FGD PLANT

CT-101

14 plants started up during 1972 to 1975.

CT-121

21 plants awarded to date.

PLANT OWNER

LOCATION

PLANT CAPACITY

(MW)	FUEL

START UP

Gulf Power

Mitubishi Petrochemical
Nippon Mining
Toyama Kyodo Power
Kashima North Power
Hokuriku Power
Japan Oil Shale
State of Illinoi
Wieland Werke AG
Hokuriku Power
Tioxide Group PLC
Georgia Power
Chubu Power
Hokuriku Power
Chemical Co.

Okinawa Power
Western Mining
0. Power
T. Power
K. Power

Florida USA
Vokkaichi Japan
Chita Japan
Toyama Japan
Kashima Japan
Toyama Japan
Kitakyushyu Japan
Illinois USA
Voehringen Germany
Fukui Japan
Terer.gganu Malaysia
Georgia USA
Hekinan Japan
Ishikawa Japan
Japan

Gushikawa Japan

Australia

Japan

Japan

Japan

20x2	Cu Smelter gas

156	Coal

1000	Coal

900x2	Coal

110	Coal

700	Coal

500	Coal

85	Coal

156	Coal

40 Coal
12 Coal
350 Oil

225 Olimuision
250 Oil
2 Oil Shale

23	Coal

85	Oil

75	Oil

500	Coal

14 Calcine r Off Gas

1978

1982

1983

1984

1985
1987

1987

1988
1990

1990

1991

1992

1993
1995

1993

1994
1993

1995
1998
1997

Several plants expected in Japan before 2000.

94-4

-------
©	BOOSTER FAN

©	STACK

-------
Efficiency(I)







\ i I i i \ \ © \ i















































(























n \ii



















! j | | |





































't : : ¦





















! I





























*1	o. s 1	s 10 dp ( fi m)

Figure 3. Collection Efficiency of Particulate Matters

94-6

-------
A FINAL REPORT ON THE MOVING-BED LIMESTONE EMISSION CONTROL (LEC)

PROCESS PILOT PLANT PROGRAM

M. E. Prudich
Ohio Coal Research Center
College of Engineering and Technology
Ohio University
Athens, OH 45701

K. W. Appell
J. D. McKenna
ETS, Inc.
1401 Municipal Road
Roanoke, VA 24012

Abstract

ETS, Inc., a pollution consulting firm with headquarters in Roanoke, VA, has developed a
dry limestone-based flue gas desulfurization (FGD) system. This S02 removal system, called
Limestone Emission Control (LEC), can be designed for installation on either new or
existing coal-fired boilers. A nominal 5,000 acfm LEC plot plant has been designed,
fabricated and installed on the slip stream of a 70,000 pph stoker boiler providing steam to
Ohio University's Athens, Ohio campus. The SO2 in the flue gas reacts with wetted granular
limestone that is contained in a moving bed. A surface layer of principally calcium sulfate
(CaS04) is formed on the limestone. Periodic removal of this surface layer by mechanical
agitation allows high utilization of the limestone granules.

Hie primary goal of the current study is the demonstration of the techno/economic capability
of the LEC system as a post-combustion FGD process capable of use in both existing and
future coal-fired boiler facilities burning high-sulfur coal. Based on the results of the
moving-bed pilot program, an in-house economic study has shown the LEC to have definite
economic advantages (both in capital and operating costs) when compared to a wet limestone
scrubber.

A total of over 90 experimental trials have been performed using the pilot-scale moving-bed
LEC dry scrubber with run times ranging up to a high of 125 hours. The primary
processing variables studied were: superficial flue gas velocity (0.5 to 1.6 ft/s), limestone
bed velocity (4.4 to 21.6 ft/hr), inlet S02 concentration (400 to > 3000 ppmdv), and water
addition rate (0 to 2 gallons/min - includes both over-bed sprays and flue gas humidification
water).

SQz removal efficiencies as high as 99.9% were achievable for all experimental conditions
studied during which sufficient humidification was added to the LEC bed.

95-1

-------
Technical Discussion
LEC Basics

The limestone Emission Control (LEC) process is a unique system employing standard
quarry-sized limestone to remove SCX from coal-fired boiler flue gases. In the LEC process,
hot flue gases (<350°F) are contacted with a bed of 1/32 to 1/4 inch limestone granules
covered with a thin film of water. Sulfur dioxide is absorbed from the flue gas into the
water film where it subsequently reacts with dissolved limestone. A layer of reaction
products, primarily calcium sulfate and calcium sulfite, forms on the surface of the limestone
granules as the reaction proceeds.

History of LEC Development

In the early 1970s, the U.S. Department of Energy initiated bench-scale testing of dry
limestone for flue gas desulfurization. Shale and Stewart1 reported that SO2 removal
efficiencies of greater than 90% for extended periods of time were possible, and that the
limestone granules were regenerated easily with mild agitation. This early work resulted in
U.S. Patent No. 3,976,747s entitled "Modified Dry limestone Process for Control of Sulfur
Dioxide Emissions."

ETS, Inc. began development work on the present LEC system in mid-1982. In June 1983,
ETS and the U.S. EPA agreed that the LEC was potentially a simple, low cost FGD system.
It was agreed to proceed along three paths: (1) ETS performed a series of bench-scale slip
stream tests to verify the basic SOj emission reductions. Results showed greater than 90%
S02 removal; (2) EPA/IERL requested that TV A do a preliminary economic analysis of the
LEC system. The resulting analysis was in agreement with preliminary ETS economic
studies showing that the LEC process offers a significant economic advantage over both
spray drying and wet scrubbing systems; and (3) EPA/IERL initiated an in-house laboratory
bench-scale study of the LEC process aimed at verification of the ETS dip stream tests.
Although Step 3, EPA/IERL's in-house work, was not formally published, ETS arrived at
two significant conclusions based on the EPA bench-scale testing program. First, EPA's
bench-scale tests using simulated flue gases adequately substantiated ETS' slip stream test
results and Shale's early work at DOE. And second, although regeneration of the spent
limestone was demonstrated, the regeneration was not accomplished as easily as that reported
by Shale and Stewart, i.e., more vigorous grinding was required.

In the fall of 1986 the Ohio Coal Development Office awarded a research grant to Ohio
University to design, install and operate a small, semi-batch fixed-bed LEC pilot plant (400
acfm) on a slip stream of flue gas from the 70,(XX) lb/hr stoker boiler providing steam to the
Ohio University campus. Figure 1 shows a simplified flow schematic of the fixed-bed LEC
pilot unit. The flue gas first entered a spray chamber where it was conditioned to the desired
temperature and humidity. The conditioning was achieved by adding water via an atomizing
nozzle and/or by injecting live steam into the chamber. The conditioned gas then entered the
LEC reactor, passing downward through a fixed-bed of limestone. The reactor contained a
removable basket filled with limestone. The basket was 22 inches square and permitted the

95-2

-------
TO
STACr

FLUE
GASES

\	 ATOMIZING

NOZZLE

OVERHEAD
SPRAYS

±±±

STEAK
INJECTORS

ID FAN

.LIMESTONE
BED

REACTOR
DRAIN

Figure 1. Simplified schematic of the fixed-bed LEC process.

limestone bed depth (the bed dimension in the direction of flue gas flow) to be varied
between 0 and 24 inches. The scrubbed flue gases were then drawn out from the bottom of
the limestone bed by an induced draft fan and returned to the duct leading to the stack.

Over 100 experimental trials were conducted using the LEC fixed-bed pilot unit between
May 28, 1987 and December 18, 19873. Three different Ohio limestones covering the full
range of Ohio limestone compositions were used in these tests. The program studied the
effects of varying limestone bed depth (from 6 to 18 inches), superficial flue gas velocity
(from 1.0 to 2.0 feet per second), flue gas humidity, Met flue gas SOj concentration, and
inlet flue gas temperature. All three of the limestones used in the study were able to achieve
greater than 90% S02 removal in the LEC process. Many tests showed removals of up to
99%.

The fixed-bed LEC unit was shown to perform effectively for inlet S02 concentrations in the
flue gas ranging from 500 ppm to 3500 ppm. At an inlet SO2 concentration of 700 ppm, a 6
inch deep LEC bed was able to sustain SO2 removals of over 90% for 16 hours and removals
of over 50% for more than 24 hours (Figure 2). Even at 3500 ppm SO2, the 6 inch LEC

95-3

-------
100-

87121%

1 u *** * L

0 \ * *

wioja I



c 80-



Q.

60-

\

\ ^



I

Removal

O

. ! > . .

k

ft

\

^871217C

V
%



V

CM

0

01	20-















C

1 1 1 1 1 i 1 1 1 j 1 1 i 1 1 j i i 1 j 1 1 1 1 1

4 8 12 16 20 4 24

Elapsed Run Time,

1 1 1 1 1 1 1 i

28 32

hours

I 1 J i "1

36 40

Figure 2. Deactivation of an LEC bed due to surface blinding. [Run data from

Reference 3 with a limestone bed depth of 6 inches and a superficial flue gas
velocity of 1.0 ft/s.)

bed (the smallest bed tested) was able to sustain an SO2 removal efficiency of greater than
90% for about 1 hour and a removal efficiency of greater than 50% for more than 2 hours.

Two factors were identified which limit LEC operation. These were (1) drying of the
limestone surface and (2) build-up of the reaction product on the reactive limestone surface.
A dry limestone surface is effectively non-reactive under LEC processing conditions. The
build-up of reaction product on the limestone surface results in a decrease in the reaction rate
and eventually renders the limestone inactive.

Two operational modes were used to overcome deactivation by drying. Operating the
process with a saturated flue gas would halt evaporation and thereby prevent the limestone
bed from drying out. An over-bed water spray could also be used to keep the bed wet. In
order to keep the bed from drying out, the spray rate would have to be equal to or greater
than the rate of evaporation.

Deactivation due to surface blinding cannot be avoided but it can be overcome through the
removal of the reaction products. After the limestone became non-reactive due to the
blinding of the surface, the solids still retained a core of unreacted limestone. The layer of
reaction products could be removed by mild attrition. In the small plot plant studies the

95-4

-------
reaction products were removal from the reacted limestone through off-line attrition in a
stirred ball milL After attritting, the stone was returned to service. When the surface of the
stone was again blinded this process could be repeated until all of the limestone had been
utilized. Hie reactivated stone showed the same S02 removal activity as the virgin
limestone.

Three methods were used to supply water to the LEC process. One method was to prewet
the bed of limestone before starting the process. This was accomplished by pouring water
onto the bed before sealing the reactor. The life of the bed was determined by the amount of
conditioning of the incoming flue gas. The higher the relative humidity of the gas the longer
the time before the limestone bed became dry and therefore unreactive. The second method
was by condensation. Normally the limestone bed was cooler than the incoming flue gas.
As the gas contacted the cooler bed, water would condense from the gas phase onto the
limestone. The amount of water deposited by condensation would depend upon the
conditioning of the flue gas. As the relative humidity increased the amount of water
deposited onto the limestone would increase. Condensation would slow and eventually cease
as the bed was heated by the incoming flue gas. After condensation stopped the life of the
bed was once again determined by the condition of the incoming flue gas. The final method
of water introduction was by spraying water directly onto the limestone bed. As shown in
Figure 1, spray nozzles were positioned directly over the bed. The life of the bed was
determined by the rate of the water spray as well as the condition of the flue gas.

Figure 3 shows experimental data taken from the LEC pilot unit that illustrates both
deactivation due to drying and surface blinding. Runs 870714A', 870716A, 870716B,
870717A, 870720A and 870720B represent consecutive runs using the same 12 inch deep bed
of limestone. The superficial gas velocity remained constant from run to run at 1.5 ft/s but
the SO, concentrations ranged from 620 to 870 ppm. The inlet flue gas relative humidity
was approximately 50% but varied slightly from run to run. The variation of SOa
concentrations and gas humidities was due to the fluctuating conditions of the incoming flue
gas. Run 870714A' started with a prewet bed of limestone. After a little over 1.5 hours the
bed became unreactive. The same bed of limestone was then rewet before starting run
870716A. The bed regained its initial SOz removal efficiency before once again losing its
reactivity due to drying. The process of rewetting the bed before starting the next run
continued for the remaining four runs. Starting with run 870717A, rewetting the bed failed
to return its activity to its original level. This decrease in activity can be attributed to the
blinding of the surface due to the build-up of the reaction products on the reactive surface.

Deactivation due solely to surface blinding is illustrated in Figure 2. These data show the
performance of three different 6 inch deep beds of limestone. The inlet flue gas was
saturated and therefore drying of the bed was not a factor in the decline of the SOj removal
efficiency of the limestone. Run 871217C is a continuation of 871216B and uses the same
bed of limestone. Run 871105A continued using the same bed of limestone started by run
871103A. As can be seen from Figure 2 each bed lost activity at different times. The life
of the bed was inversely proportional to the product of the SQ, concentration of the incoming
flue gas and the flue gas flow rate. Run 871103A/871105A with an inlet S02 concentration
of 344 ppm lasted the longest. Next was run 871214B with an SQz concentration of 702

95-5

-------
ELAPSED RUN TIME (hr)

Figure 3. Deactivation of an LEC bed due to surface drying and surface blinding. [Run
data from Reference 3 with a limestone ted depth of 12 inches and a
superficial gas velocity of 1.5 ft/s.]

ppm. Run 871216B/871217C with an Met S02 concentration of 736 ppm had the shortest
life. The Met flue gas flow rate was held constant for all three sets of runs.

Besides the difference in operating conditions, the limestone used for the runs shown in
Figure 2 differs from the limestone used for the runs shown in Figure 3.

Moving Bed Process Description

A continuous moving-bed LEC pilot plant sized to handle 5000 ACFM has been installed on
the dip stream of the 70,(XX) Ib/hr stoker boiler providing steam to the Ohio University
campus. A amplified flow scheme of the moving-bed LEC process is shown in Figure 4. A
detailed schematic of the pilot plant is given in Figure 5. Limestone is added to the bed via
a 32 foot bucket elevator which is supplied with limestone from the feed hopper, the make-
up hopper, the vibrating screen, or the recycle chute by means of a feed screw. The make-
up hopper is used to compensate for limestone lost due to reaction with S02 and its
subsequent removal as reaction product. The limestone bed itself has dimensions of 14
inches X 36 inches X 128 inches (Figure 6). The depth of the bed in the direction of the

95-6

-------
flue gas flow is 14 inches. The LEC
reactor internals consist of inlet and outlet
louvers, overhead bed sprays and outlet
screening. The inlet and outlet louvers (as
originally designed) are four inches long
and are inclined at a 75° angle with a 3/8
inch overlap between louvers. The flue
gas inlet and outlet plenums on the sides
of the reactor are sized such that there is
uniform distribution of the gas across the
whole bed. The flue gas is drawn from
the duct leading to the stack from the
electrostatic precipitator. The superficial
velocity of the flue gas through the LEC
reactor can be varied from 0.3 to 1.6 feet
per second. The live hopper at the top of
the bed is provided so as to prevent the
leakage of ambient air into the outlet gas
plenum. An induced draft fen is used to

draw the processed flue gas from the	Figure 4. Simplified schematic of the LEC process,

outlet plenum. The processed flue gas is
directed to the stack.

WET
LIKESTONE
BED

rim

GA



DRY
LIMESTONE

BED

DESULFTOIZED
FLUE
GAS

Figure 6. Limestone bed geometry for
moving-bed LEC pilot plant.

limestone moves vertically downward
through the bed with a velocity sufficient to
prevent bridging of the limestone which
may occur when the flue gas dries up the
bed. The limestone removed from the
bottom of the bed is sent to the discharge
screw via a 19 foot bucket elevator. The
discharge screw has three outlets, any one
of which may be opened. The first outlet
directs the limestone to the recycle chute
which leads directly to the feed screw and
back to the reactor. The second outlet
leads to the reactivation system. The build-
up of reaction products on the limestone
surface tends to retard the rate of reaction
per unit granule surface area. To prevent
this surface blinding from rendering the
limestone inactive, the limestone granules
are passed through a reactivation device that
removes the reaction product layer by either
diy or wet abrasion. This abrasion exposes
a fresh layer of reactive limestone. This
mixture of reactivated limestone and

95-7

-------
LEGEND

1-Feed	Hopper

2-Waste	Hopper

3-Make-up	Hopper

4-Screen

5-Fluidizer

6-Bucket	Elevator

7-Bucket	Elevator

8-Inlet	Gas Plenum

9-Outlet	Gas Plenum

10-Reactor

11-ID	Fan

12-Bed	Twin Screws

13-Livo	Hopper

14-Inlet	Gas

15-Outlet	Gas

16-Reactor	Screw
17 Discharge Screw

18-Feed	Screw

19-¥aste	Screw

-------
abraded material is then passed through a vibrating screen were the abraded material is
removal as a waste. The reactivated limestone is then directed to the feed screw and
returned to the top of the LEC bed. The third discharge screw outlet is directed to the waste
hopper. Normal LEC operation directs the limestone through the reactivation system.

The S02 concentration in the inlet flue gas can be varied from 500 to 3000 ppm. Early
experimental runs using this pilot plant have shown that the LEC process has a capability of
removing more than 98% of the SOj contained in the inlet flue gas. The LEC process is
considered to be potentially attractive for enabling coal-fired power plants to meet stringent
S02 emission standards.

Discussion of Results

Initial Moving-Bed Experimental Results

A total of 30 experimental runs were conducted between May 29, 1991 and July 25, 1991
using the LEC moving-bed pilot unit. The length of each experimental run ranged from 30
minutes to 290 minutes. The primary variables studied were:

•	Superficial flue gas velocity - 0.5, 1.1, and 1.6 ft/s.

•	Limestone bed velocity - 4,4 ft/hr, 13.7 ft/hr, and 21.6 ft/hr
(these limestone bed velocities correspond to in-bed limestone
residence times of 2.4 hr, 0.8 hr, and 0.5 hr).

•	Inlet SQz concentration - 500 to 1100 ppm.

•	Water addition rate - 0 to 2.0 gallons/minute (includes both over-bed sprays
and flue gas humidification water).

As mentioned earlier, the SC^/CaCOs reaction is a liquid-phase reaction and therefore a thin
layer of water must be present on the limestone granules to allow significant SOb removal to
occur. The run philosophy to date has been to operate the LEC bed so that the limestone
exiting the bed is essentially dry. This is accomplished by controlling the position of the
drying front that exists in the limestone bed. As the limestone approaches the bottom of the
bed, the drying front moves closer to the outlet side of the reactor. This results in there
being less available wet limestone for S02 removal. Although this results in relatively poor
S02 removal values in the lower portion of the bed, the overall average bed removal remains
high because of the excellent performance of the top and middle portions of the bed.

The results obtained from the initial experimental trials have demonstrated SO2 removals in
excess of 90% and, often, SO, removals in excess of 99%. Performance plots for a
representative experimental trial (910723A) are shown in Figure 7. This figure also shows
the influence of die mechanical difficulties that inhibited long-term operation during the
initial period of operations. The outlet design of the LEC moving-bed pilot plant relied only

95-9

-------
~

/

/

s g*Ko 9

CM :

O 60-

tn :

Run No. 910723A

50

I | I I I I | I 1 1 I i ! I 1 1 J ! (l I I I O	*-0

50	100 150 200 250

Elapsed Run Time, min

Figure 7. LEC moving-bed experimental results. Trial number 910723A. AASHTO
No. 9 limestone. Met SQ,: 1100 - 1200 ppmdv.

on the outlet louver to retain toe Emestone in the moving bed. As it has turned out, the
original outlet louver design did not accomplish this purpose. During operations, limestone
was drawn from the bed into the outlet plenum and subsequently into the I.D. fen. In order
to prevent this limestone loss from the bed and to protect the I.D. fan, a screen was attached
to the inside (limestone bed side) of the outlet louvers. The screen was used to keep the
limestone in the bed. However, the presence of the screen introduced another operational
problem. Damp fines were deposited on the screen, plugging the screen and significantly
increasing the pressure drop across the limestone ted. The blinding of the screen occurred
most rapidly at the top of the bed (normally the most active part of the bed). Due to the
screen blinding at the top of the bed, the flue gas would begin to flow preferentially through
the bottom portions of the bed (normally the least active part of the bed). This action would
result in a simultaneous increase in pressure drop across the bed, decrease in SO2 removal
performance, and decrease in flue gas flow rate. This behavior can be observed in Figure 7.

After the initial run series, the outlet louvers and outlet planum wore redesigned. The
redesign modifications for these systems were made during a September/December 1991
shutdown. The new outlet louvers were eleven inches long with a 3-1/2 inch overlap
between louvers. The new and old outlet louver designs are shown in Figure 8. Results
from shakedown runs performed in February 1992 have indicated that limestone loss from
the reactor bed into the outlet plenum was reduced by several orders of magnitude.

It has been observed that significant limestone surface abrasion occurs within the LEC bed

95-10

-------
Figure 8. Old and new outlet louver designs used in the LEC moving-bed reactor.

itself. 3ii traveling downward through the LEC bed and in transport by the screw conveyors,
the limestone particles seem to provide most of the granule on granule abrasion necessary to
reactivate the limestone. In cases where this effect may be fully exploited, the external
reactivating attritor (equipment item #5 in Figure 5) may be eliminated from the process flow
sheet.

Enlarged Louver Design Experimental Results

A total of 60 experimental runs were conducted from May 5, 1992 through November 24,
1992. The duration of these experimental runs ranged up to 134 hours. The most important
variables studied, primarily with respect to overall SO2 removal efficiency, bed pressure drop
and stone carry-over, were:

•	mean granule diameter.

•	superficial gas velocity.

•	inlet S02 concentration.

Effect of Mean Granule Diameter. The four granule diameters that were studied in the
LEC moving-bed were:

•	AASHTO No.9 (Dp = 0.10 inches).

•	AASHTO N0.8 (Dp = 0.28 inches).

•	50/SO^otum- basis) No.8/No.9 mixture (Dp = 0.20 inches).

•	AASHTO No.9 pre-screened with a N0.8 mesh (Dp = 0.12 inches).

95-11

-------
100-

c
o

tv

75-

O

£

CO 50-

B

25-

O

0-
0.1

Particle Size, mm

10

Figure 9. Typical particle size distributions for the
AASHTO No. 8 and No. 9 limestones
used in LEC processing.

Typical particle size distributions for the two major stone sizes [AASHTO No. 8 and No. 9]
used in LEC processing are shown in Figure 9. The No.9 limestone has the smallest mean
diameter of the four types used and therefore has the largest surface area available for
reaction. Unfortunately, the diameter of a limestone granule is inversely related to its
minimum fluidization velocity. At smaller particle diameters this can result in a portion of
the limestone becoming fluidized by the flue gas stream and, subsequently, exiting the bed
through the outlet louvers. At superficial gas velocities greater than 0.75 ft/s, excessive
limestone carry-over was observed to occur. In order to reduce the fluidization potential of
the stone, limestone sorbent granules with increased mean diameters were studied.

The next larger commercially available size stone is AASHTO No.8 (Dp = 0.28 inches).
With the relatively large- particle diameter of the No.8 stone, the carry-over problem was
greatly reduced, even at superficial flue gas velocities in excess of 1.2 ft/s. The improved
material handling characteristics of the No.8 limestone enabled the first extended trial of the
moving- bed system (Figure 10). Run 920430 lasted for over 134 hours at an average
superficial gas velocity of 1.1 ft/s and achieved SO2 removal efficiencies of up to 85 percent.
However, subsequent experimental trials (LEC Runs 920601 and 920602), also using No. 8
limestone, that ware operated at Iowa1 superficial gas velocities (0.3 ft/s) with a 100 percent
bed moisture zone yielded S02 removals of only 40% (average). With die enlarged mean
diameter of the No.8 stone, the loss in available reaction surface area resulted in substantially
lower S02 removals when compared to No. 9 limestone. The low S02 removals resulting

95-12

A

-------
100

0

\> • i» \i
Run No. 920430

i—i—i—i—|—i—i—r

0	25

—j—r—r-T—i—j—

50	75

'I 1 1 1 1 I 1

100 125

.CM

o

CM
X

-oo

-to

Elapsed Run Time, hours

Q.

o

0

15

m

CD

-CN

o
,o

ro

.§
o

K>

o
.o
in

o

¦8

CM

o
.©
m

o

.o
o

o
¦o
m

"-o

U
D

i

o

Ll.

W
O
O

Q)

c

Figure 10. Long-term LEC operation using AASHTO No. 8 limestone. Run No. 920430.
Met S02: 800 - 1400 ppmdv.

from the use of No. 8 stone are unacceptable in terms of operating a viable commercial
system. This low removal more than negates the No. 8 stone carry-over and material
handling advantages. Increasing the LEC bed depth beyond 14 inches would add reactive
surface area and might compensate for the disadvantages of the No. 8 stone.

The next stone size used in an attempt to reduce stone carry-over and still achieve +90%
S02 removal was a 50/50 mixture by volume of No.8 and No.9 limestones. This mixture
had an average particle diameter of 0.20 inches (LEC Runs 920604 and 920605). The
resulting S02 removal was approximately 10% greater than with No.8 stone alone, but the
limestone carry-over was still excessive. Therefore, the combination of the two stones
yielded a mixture that exhibited the worst traits of both the No.9 (high stone carry-over) and
No. 8 (low S02 removal) stone sizes.

The last

method investigated in order to reduce the stone carry-over while maintaining high

95-13

-------
S02 removals consisted of scalping (removing the fines from) the No.9 stone with a No. 8
mesh screen (0.094 inches). This was done because the theoretical fluidization velocity of a
size distribution is weighted towards the higher individual fluidization velocities of the
smaller particles within the distribution. Thus, by scalping off the smallest 40% of the No.9
limestone granules, the fluidization velocity of the remaining stone would be increased while
at the same time maintaining a larger reaction surface area than that available from the No. 8
limestone. The resulting mean particle diameter of the screened stone was 0.12 inches. The
screened stone was used in LEC Runs 920810, 920811 and 920812. These runs yielded poor
results since there was no decrease in limestone carry-over when compared to No.9 stone
(Dp = 0.10 inches) although S02 removals were generally reduced to less than 90%.

The overall results of varying the mean particle diameter in order to achieve an optimum
stone size tot would limit the amount of limestone carry-over into the outlet plenum while
maintaining high SO2 removals was unsuccessful. The carry-over problem must be solved
through changes in the mechanical design.

Effect of Superficial Flue Gas Velocity. The superficial flue gas velocity affects three
dependent variables of the LEC system:

(1)	Reaction Theory.

(2)	Stone Carry-Over.

(3)	Bed Pressure Drop.

The reaction kinetics theory with respect to superficial flue gas velocity of the LEC system
has already been extensively dealt with by both Visneski4 and Reddy5 and their general
conclusions have been validated by the moving-bed project.

Stone Carry-Over, It has been empirically determined that flue gas superficial velocities of
greater than 0.75 ft/s result in excessive limestone carry-over when using No.9 stone (Dp -
0.10 inches). Fluidization of the limestone particles is believed to take place in fee 75° outlet
louver slot since the superficial flue gas velocity is more than doubled as it passes through
the outlet louvers. The excellent S02 capture performance of an experimental trial (920613)
operated processing No. 9 stone and at a superficial flue gas velocity of 0.31 ft/s is shown in
Figure 11. Hie use of No.8 stone with an average granule diameter of 0.28 inches allows
superficial flue gas velocities of up to 1.2 ft/s without excessive stone carry-over at the
expense of the loss of reactive surface area (at constant bed depth).

Bed Pressure Drop. Bed pressure drop is related to the square of the superficial flue gas
velocity and can be quantified through the Ergon equation. This is illustrated by LEC RUN
921020, during which the bed pressure drop was observed to be less than 2 in.H20 at a
superficial flue gas velocity of 0.31 ft/s. After approximately 26 hours into the trial, the
superficial flue gas velocity was increased to 0.47 ft/s which resulted in an increase in
pressure drop to approximately 7 in.H20.

Effect of Inlet S02 Concentration. All of the LEC experimental trials were conducted
using flue gas obtained from the slip stream of the Ohio University heating plant. The

95-14

-------
100

c


Q


O

E

Q)
CK

90

80

70'

CM

O 60-
CO

50-

S02 Removal

Inlet Gos Flow

Run No. 920613

Delta P

0

i—i—i—i—|—i—i—i—i—|—r

50	100

150

t—i—r

.CM

o

CN|

¦oo

•o

200

Elapsed Run Time, min

Q.

o

s—

0

L_

W
CD
95% of the SQ, entering the scrubber. However, whenever the
moisture (damp limestone) zone was reduced in order to allow the exiting stone to remain
dry enough for effective regeneration via the vibrating screen (between hours 4 to 5, 6.5 to
8.5, and 11.5 to 13.5), the SQa removal efficiency dropped dramatically down to
approximately 70%. In terms of sorption theory, this would indicate that at these high inlet
SO2 concentrations the SOj breakthrough curve resides within (but only barely) the fully
wetted 14 inch bed. This means that the S02 breakthrough curve resides close to the outlet
side of the bed even at low superficial gas velocities (0.3 ft/s). Therefore, careM water
addition is critical in order to ensure a maximum possible moisture zone within the bed when

95-15

-------
100-

c

a>
o

u_

a>

Cl

o
>
o

£
CD

cr

04

o

if)

80-

60'

40-

20-

S02 Remcrvol

O

¦coO
CM

-<£>



/ Inlet Gas Flow

Oelta P

Run Nos. 920716/920718/920719

"T	1	1	i	1	1	1	1	]	1	J	T

5	10

Eiapsed Run Time, hrs

-CM

15

Cl
O

k_

O

-------
The LEG process and conventional limestone scrubbing have been compared on an equatable
basis using flue gas conditions that would be expected at the outlet of the electrostatic
precipitator (ESP) on a 500 MW coal-fixed power plant. The equipment list for the LEC
case was obtained by conceptually scaling up the current LEC moving-bed pilot plant. The
equipment list for the wet limestone scrubber case was obtained from an EPRI technical
report®. Equipment costs for both the LEC and the wet limestone scrubber were generally
obtained using empirical estimating charts.

The LEC was found to have a definite economic advantage in both direct capital costs and
operating costs. The LEC equipment capital cost of $12,290,000 represented a 48% savings
when compared to the wet limestone scrubber process equipment cost of $23,534,000. On
an absorption tower basis, a capital cost advantage of 46% for the LEC occurs due to its
simpler mechanical design. Operating costs are also lower for the LEC process since a
slurry is not required, resulting in a lower water consumption. The solids material handling
associated with the LEC requires less electricity than that of the slurry handling system
required by the wet limestone scrubber.

The economic results quoted above are in general agreement with a 1989 study7 that
compared the LEC with wet scrubbing, a stand-alone dry scrubbing system, and retrofit
LIMB (with humidifieation). Hie results of this study are given in Table 1.

Table 1. Economic comparison of LEC and other scrubbing

technologies7.

Technology

SO2 removal,
%

Capital costs,
$/kW

SO2 removal
costs, $/ton

LEC

90+

168

406

Wet Scrubber

90+

220

680

Dry Scrubber

70

186

740

LIMB

50

74

608

Based on the success and findings of the present project, the next step will be a full-scale
commercial demonstration unit. Sales presentations to utility and industrial sites for a
demonstration unit have already been initiated. The commercial demonstration for the LEC
is currently under consideration in Taiwan.

95-17

-------
Licensing agreements with two major air pollution control suppliers have already been
negotiated. The first was with Procedaire Industries, which is the North American subsidiary
of the highly respected industrial furnace manufacturer, Stein Heurtey. Procedaire Industries
was granted the right to sell the LEG system to the North American industrial market. The
second agreement was granted to the U-Tah Industrial Company, Ltd., Taiwan, a major
supplier of air pollution control equipment, for the exclusive rights to sell the LEG process
throughout Taiwan, R.O.C. Taiwan's FGD market is expected to reach $2.8 billion U.S. by
1997.

Conclusions

The LEG moving-bed pilot plant project has demonstrated its capability as a high-
performance SO2 removal system. S02 removals as high as 99.9% have been achieved for
extended periods at inlet flue gas flow rate of up to 3200 acfm. Met flue gas SOj
concentrations ranging up to >3000 ppm have been successfully tested. This shows that the
chemistry behind this road-grade limestone-based technology is valid.

Mechanical difficulties, i.e., limestone blow-over and exit screen plugging initially made
long-term LEC operations difficult Outlet louver modifications made over the
September/December 1991 shutdown enabled long-term testing.

The next step in the commercialization of the LEC process is the demonstration of the
technology on an electric utility boiler. A conceptual schematic of a generic compact
moving-bed LEC is shown in Figure 13.

95-18

-------
wmmsmtm



(a.) Top View

/\:»:/\:::::
/Nh^x/NA

/N::::::::;XN:il

11

/N;::::::::><\

(b.) Face View

Figure 13. Schematic of conceptual compact LEC bed. (a) Top view, (b) Cross-
sectional face view. Arrows indicate flue gas flows. Shading indicates
limestone bed, Rue gas is distributed and collected through louvered
plenums.

Acknowledgements

This study was funded in part by the Ohio Coal Development Office under grant
GDO/D-88-49.

95-19

-------
References

(1)	C. C. Shale, U.S. Patent 3,976,747 (1976).

(2)	C. C. Shale and G. W. Stewart, "A New Technique for Dry Removal of SOj,"
presented at the Second Symposium on the Transfer and Utilization of Particulate
Control Technology, Denver, CO (1979).

(3)	M. E. Prudich, K. W. Appell, M. J. Visneski, J. D. McKenna, D. A. Furlong, I. C.
Mycock, J. F. Szalay, and J. E. Wright. Small Pilot Plant Demonstration ofETS'
Limestone Emission Control System, Vols. 1 & 2. Columbus, Ohio: Ohio Coal
Development Office, 1988. CDO/R-86-24.

(4)	M. J. Visneski. Modeling of the Low Temperature Reaction of Sulfur Dioxide and
Limestone Using a Three Resistance Film Theory Instantaneous Reaction Model,
Ph.D. Dissertation, Ohio University, 1991.

(5)	S. N. Reddy. A Mathematical Model ofETS' Limestone Emission Control Process
Using a Moving Bed Configuration, M.S. Thesis, Ohio University, 1991.

(6)	R. J. Keeth, M. J. KrajewsM, and P. A. Ireland. Economic Evaluation ofFGD
Systems, Volume 5; The NOKSO and SOXAL Sodiwn-Based Processes and Four
Additional Calcium-Based Processes. Palo Alto, Calif.: Electric Power Research
Institute, 1986. CS-3342-V5.

(7)	Ohio Coal Development Office. Proposal CDO/89-64, 1989.

95-20

-------
TUNG FGD PILOT PLANT PERFORMANCE

Peter Strangway
Niagara Mohawk Power Corporation
300 Erie Blvd. West
Syracuse, New York

Shao Tung
Raycon Research & Development
91 Blake Rd, Brookiine, MA

Robert J. Keeth
Raytheon Engineers & Constructors Inc.
5555 Greenwood Plaza Blvd.
Englewood, Colorado

Abstract

Raycon Research & Development Inc. (Raycon) contracted with Raytheon Engineers &
Constructors (Raytheon) to design, construct, and operate a nominal 2 MW pilot
facility. This test program will demonstrate the performance of the Tung FGD Process
when treating flue gas from a medium-to-high-sulfur, coal-fired utility boiler at Niagara
Mohawk Power Corporation's Dunkirk Station.

The Tung FGD Process pilot plant is designed to remove more than 90 percent of the
S02 from the flue gas in a conventional prescrubber/sodium sulfite packed tower
absorber system. The resulting bisulfite solution is extracted with an organic solvent,
thus regenerating the sodium sulfite. The organic solvent loaded with sulfite is then
steam stripped in a stripper column, producing a concentrated S02 gas stream. The
organic is returned to the extraction system for reuse.

This paper will present the project history, design considerations, and available
performance data generated during the parametric testing program, currently
scheduled for July-September, 1993. Integrated system operation will be evaluated to
determine the optimum operating conditions for the long-term reliability of the system.
Available data summarizing the parametric test performance to-date will be included in
the presentation.

96-1

-------
Introduction

Development of the Tung FGD Process began in the early 1980's and progressed
through a laboratory development phase to the point where a small pilot plant test
program was necessary to demonstrate the performance of the process on the flue
gas from a coal-fired utility boiler. Laboratory testing continues to enhance the
performance of various subsystems.

The program is funded by the Department of Energy (DOE), Niagara Mohawk Power
Corporation (Niagara Mohawk), the Empire State Electric Energy Research
Corporation (ESEERCO), and the New York State Energy Research and Development
Authority (NYSERDA). Raytheon Engineers & Constructors (Raytheon), Englewood,
Colorado, was contracted to provide engineering support to design, construct, and
operate a nominal 2 MW equivalent pilot plant. The pilot facility will demonstrate the
performance of the Tung Flue Gas Desulfurization (FGD) Process when treating flue
gas from a medium-to-high-suifur, coal-fired utility boiler. The pilot facility is installed at
Unit No. 4 of Niagara Mohawk Power Corporation's Dunkirk Station at Dunkirk, New
York.

Tung FGD Process Summary

The Tung Flue Gas Desulfurization Process utilizes an innovative liquid/liquid
extraction system to regenerate a sodium-based aqueous scrubbing liquor. Steam
stripping of the solvent recovers S02 as a byproduct, and requires only a fraction of
the energy needed to steam strip S02 from the aqueous solution itself.

A simplified process schematic developed for the Tung FGD Process is shown in
Figure 1. The process consists of three primary subsystems:

•	Scrubbing	S02 in the flue gas is scrubbed with a sodium sulfite

solution yielding a sodium bisulfite solution:

S02 + Na2S03 (aq) + H20 -~ 2NaHS03 (aq)

•	Extraction	S02 in the scrubbing bisulfite solution is transferred to an

organic solvent. The regenerated sulfite solution can be
reused for scrubbing:

2NaHS03 (aq) + Org. -* loaded Org. + Na2S03 (aq)

•	Steam Stripping The ioaded organic is steam stripped to separate S02

gas from the organic. The regenerated organic can be
reused for extraction.

loaded Org. -* Org. + S02

96-2

-------
The process also involves several minor steps such as sulfate formation and removal,
solvent regeneration, etc. They will not be discussed in this paper.

Major advantages of the Tung FGD Process include:

•	Relatively low steam consumption - about one pound of steam per pound of
S02 recovered is projected.

•	Relatively low scrubbing chemical consumption - the scrubbing chemical is
regenerated within the process. Only the sodium removed as sodium sulfate
waste or byproduct needs to be replaced.

•	Relatively low volume of waste products - sodium sulfate is produced in
moderate amounts and can potentially be marketed as a byproduct. The
prescrubber blowdown stream is neutralized and sent to disposal. There may
also be some small amount of organic bled off from the process. This organic
waste from a commercial plant can be incinerated off-site or perhaps mixed in
with the utility plant fuel and burned at the site.

Pilot Plant Description

The pilot plant is being installed at Niagara Mohawk's Dunkirk Steam Station located
on the south shore of Lake Erie about 45 miles southwest of Buffalo, New York. The
Dunkirk Steam Station Unit No. 4 (nominally 200 MW gross capacity) provides the flue
gas source for the pilot plant. The coal burned at the Dunkirk Steam Station is
bituminous coal from a variety of mines in western and northern Pennsylvania and
West Virginia. This coal has a sulfur content ranging from 2-3% by weight.
Approximately 70 percent of the coal used at the Station is washed and the remainder
is raw coal. The net heat rate for the Unit averages 9590 Btu/kWh. An ESP operates
upstream of the flue gas extraction point, providing an inlet particulate concentration of
about 0.04 Ib/MBtu.

The Tung FGD Process pilot plant is designed to remove more than 90 percent of the
S02 contained in a nominal 2 MW flue gas slipstream extracted from the Unit No. 4
ductwork downstream of the ESP and I.D. fan. The pilot plant is installed beneath the
existing Unit No. 4 precipitator. The following material describes the equipment
components currently incorporated into the system design. Figure 2 is a somewhat
simplified process flow diagram for the Tung system.

S02 Removal

The Niagara Mohawk power plant provides all utilities supplied to the pilot plant,
including fresh water, plant and instrument air, cooling water, steam, deionized water,
and flue gas. A 3 MW equivalent slipstream of the flue gas is taken from the ductwork
downstream from the existing plant I.D. fan. Two-thirds of this gas (approximately
6,500 ACFM) flows into a venturi scrubber/saturator constructed of rubber-lined
carbon steel with a Hastelloy venturi throat. The other third of the flue gas bypasses
the absorbers and serves as reheat for the scrubbed gases prior to their return to the
utility ductwork.

96-3

-------
In the venturi scrubber, the flue gas is contacted with recycled water. Fly ash,
hydrogen chloride/fluoride and a small amount of S02 are removed in this quench
operation. A purge stream from the venturi liquor recycle loop is pumped to the
quench biowdown/neutralization tank where lime is added for pH adjustment The
purge stream is then combined with organic-contaminated waste water and sent to the
local city sewer system.

Saturated flue gas exiting the venturi will flow through a cyclonic separator to remove
residual acidic droplets, then flow into and up through a packed tower absorber where
the flue gas is contacted with the recycled scrubber liquor to remove S02. The
packed tower absorber is constructed of 316L stainless steel. Two, 100% capacity,
recycle pumps are installed to supply both the absorber and the venturi/quencher,
allowing continuous operation of both units. A portion of the absorber recycle liquor is
diverted through a bleed filter for residual solids removal, prior to sending it to the
sodium bisulfite liquor tank. The bisulfite liquor leaves the scrubber at approximately
52°C (125°F). This solution is cooled to 24°C (75°F) by a heat exchanger and is
then pumped to the sulfite extraction system.

The scrubbed flue gas, with at least 90% of the S02 removed, exits the absorber
packing and then passes through a mesh pad mist eliminator before finally leaving the
scrubber. Scrubbed flue gas mixes with the by-passed flue gas for reheat, and then
passes through the pilot plant I.D. fan. All pilot plant ductwork is constructed of
rubber-lined carbon steel. The concentrated regenerated S02 from the stripper off-gas
condenser is combined with the reheated gas, and then returned to the Niagara
Mohawk Unit No. 4 flue gas duct system.

Sulfite Extraction

The bisulflte-rich spent scrubber liquor is pumped to a five-stage mixer-settler
extraction system where it mixes with an organic solvent. The solvent extracts sulfite
from the bisulfite solution, thus regenerating the sodium sulfite scrubbing liquid.
Following the sulfite extraction from the solvent, the regenerated sodium sulfite solution
is separated from the organic solvent by a centrifuge. This sulfite solution then serves
as the regenerated feed for the S02 scrubber. Any sodium losses are made up by
direct feed of soda ash solution.

The loaded organic stream exiting the mixer-settlers also passes through a centrifuge
to remove any entrained aqueous solution. The loaded organic flows to a storage
tank, and the solvent is then preheated to 205°F and pumped to the top of the steam
stripper. A nitrogen blanket is maintained in all solvent systems to prevent oxidation of
the organic reagent and sulfite solution.

96-4

-------
Steam Stripping

Steam is fed into the bottom of the stripping column to strip S02 from the suifite-
loaded solvent stream. The lean organic solvent stream exiting from the bottom of the
stripper is first separated from any entrained aqueous fraction in a horizontal settling
tank. The regenerated solvent is then cooled and returned to the lean solvent storage
tank. The stripping column is a 30-tray, 10-inch diameter column constructed of alloy
steel. Multiple inlets are provided in the column to allow evaluation of performance for
various sizes of columns.

The S02/steam stream exiting the stripper passes through the S02 off-gas heat
exchanger, which condenses the water in the S02 gas stream. The S02 gas is
returned downstream of the pilot plant l.D. fan and is mixed with the reheated flue gas.
In a commercial plant, the S02 would be processed to provide a saleable byproduct.

A general arrangement drawing is presented in Figure 3 showing the locations of
various subsystems on the pilot plant site.

Pilot Plant Design

While all major process steps use only standard chemical process equipment
(absorber, extractors, steam stripper), design of these equipment components for the
pilot plant was not always straight forward. There were specific problems that
emerged either due to the nature of the process or the size of the pilot plant. This will
be illustrated based on the design of the steam stripper.

Design Data

Vapor-liquid equilibrium data, chemical reaction rate data, and relevant hydraulic data
were gathered before starting the steam stripper design. These data are presented

below:

1. Vapor-liquid equilibrium data were assembled by Badger Engineers, Inc. (now
Raytheon Engineers and Constructors). The data, which is presented in Figure
4a, shows that the loaded organic is relatively easy to strip. At 70% sulfur
loading, the vapor pressure at 100°C is about 260 mm Hg.

96-5

-------
2. A Thermogravametric Analysis (TGA) rate study of the loaded organic

decomposition was conducted aid the results are summarized in Figure 4b.
This figure shows that at 100°C, complete decomposition of the 100% loaded
organic can be achieved in about 4.9 minutes. The variation of rate constant
with temperature is shown in Figure 5. From these data, a rate equation for
decomposition was developed:

JL» = k (i_a) = 0.019 (1 -a) at 100°C
at

Where « = fraction of decomposition

The significance of this data is that in the stripper column design, the residence
time of the organic liquid in the column must be longer than 4.9 minutes. To
provide a reasonable operating margin, 5.5 minutes was used as the residence
time in tine design specification.

3. Using a two-inch diameter Oldershaw column, a series of laboratory-scale

steam stripping tests with the loaded organic were completed. The results are
summarized in Figures 6a and 6b and Table 1. These data show the
importance of turbulence when stripping the organic liquid. In particular, Table
1 shows that to achieve effective stripping, the F value must be larger than 0.24.
The F value, defined as vVp, where v is the superficial steam velocity and p is
the vapor density, has been regarded by many as a mixing index. When it
drops below 0.24, the turbulence is not sufficiently intense, and the stripping
performance, as indicated in Table 1 as the percent stripped, decreased.

Design Considerations

1. Column Type - Three different stripper column designs were considered for
stripping the loaded organic: packed column, valve tray and sieve tray. The
type of column selected is a judgemental decision., In the test program
planning, a mathematical model was constructed to guide in the selection.
Unfortunately, for logistical reasons, the model computations were not
completed prior to the design selection; therefore, the decision was made
without the benefit of the model's guidance.

Although a packed column, under certain circumstances, can give a high F
value, it was decided that the F value in this instance was not a good criterion.
In a packed tower, mass transfer is predominantly the across-the-film mass
transfer, and that is not the type that would best meet the needs of the process.
Therefore, in spite of the fact that F value could be large, a packed column was
not suitable for this application.

96-6

-------
The decision of whether a vaive tray or a sieve tray should be used was a more
difficult one. It was concluded:

•	The valve tray column has an advantage over the sieve tray column in
providing larger turndown (because slots in the valves can be reduced
as the gas loading decreases).

•	For large industrial columns, the stripping efficiencies of these two types
of columns are comparable. However, for a 12 inch (or smaller)
diameter valve tray column, data were not available to confirm that
performance of the column would not suffer because only a small
number of valves can be positioned on a tray. Furthermore, most valve
fray vendors do not want to supply these trays if the column diameter is
less than 12 inches.

Therefore, selection of a valve tray versus a sieve tray depends primarily on the
column diameter. If the column diameter is larger than 14 inches, a valve tray
column should be favored. On the other hand, if the column diameter is less
than 12 inches, a sieve tray should be selected. The fact that the laboratory-
scale Oldershaw column, which used a sieve tray, worked well also favored the
selection of the sieve tray.

Column Diameter - Four pounds of steam per pound of S02 released, which
had been successfully demonstrated in Oldershaw column, was selected as the
design point. To be suitable for this application, the column must satisfy the
following hydraulic and kinetic conditions:

•	The F valve must be greater than 0.24

•	The liquid residence time must be greater than 5.5 minutes.

The F value calculation shows that for columns having 12, 10, 9, 8, and 7 inch
diameters, the F values are 0.32, 0.48, 0.57, 0.72, and 0.95, respectively. Thus,
a 12 inch column cannot demonstrate a steam use rate of two pounds of steam
per pound of S02 released, because at that steam rate, the F value will be
0.32/2 = 0.16, which is less than 0.24. A 10 inch column could possibly
demonstrate two pounds of steam/pound of S02, because for that column, the
F value at that steam rate is 0.48/2 = 0.24, which is barely sufficient to achieve
effective stripping. The column diameter would have to be reduced to 7 inches
in order to demonstrate one pound of steam/pound of S02 since 0.95/4 s
0.24.

For a 20 tray stripping column, the liquid residence time of a 10" column with a
weir height of 2 inches is estimated to be 5.7 minutes. This is sufficient time for
loaded organic to achieve 100% decomposition at 100°C according to the
kinetic data. The smaller the column selected, the higher the weir height should
be in order to keep the liquid residence time above 5.5 seconds.

96-7

-------
Since the column diameter is less than 12 inches, a sieve tray rather than a
valve tray was selected.

3.	Column Tray Design - The layout of a standard sieve tray consists of several
areas as shown in Figure 7 Area for down flow pipe positioning from the plate
above and to the plate below, calm zones between the downflow pipes and the
active areas, and active areas where the sieve holes are placed and the
liquid/gas interaction occurs must all be taken into consideration. As the
column gets smaller in size, the areas reserved to accommodate downflow
pipes as well as the calm zones cannot be reduced proportionally. As a
consequence, the active area where liquid/gas interaction (or stripping) actually
occurs becomes smaller and smaller. Actual layout design showed that for
column less than 10 inches in diameter, an effective design cannot be attained
for a standard sieve tray layout (downcomers on either side of the tray). In
addition, most vendors have little experience in making trays smaller than 12
inches in diameter. Some companies were willing to build a 10 inch column
tray, but they could not offer guarantees that the trays would maintain
satisfactory hydraulic performance.

It was decided that in order to build a smaller than 12 inch diameter column,
conventional tray design would have to be abandoned in favor of a special
design called "orbital flow." The orbital flow design is usually selected for
systems having a low liquid flow rate. Although the Dunkirk pilot plant will have
a very low liquid flow rate, the primary reason that the orbital flow design was
selected is that it has less wasted area. As shown in Figure 7, the orbital
design puts two down flow pipes on the same side. By so doing, the wasted
area is reduced and the active area is increased. Although a 7 inch column
was possible with orbital flow design, it was decided to remain conservative and
a 10 inch column was selected instead. By deciding to build a 10 inch column
rather than a 7 inch column, the possibility of demonstrating one pound of
steam/pound of S02 was eliminated. A demonstration of a steam feed rate of
two pounds of steam/pound of S02 will be adequate for the current pilot test.
Demonstration of one pound of steam/pound of S02 may occur in a larger-
scale demonstration.

4.	Column Height - Even though 20 trays were found to be sufficient in the
laboratory-scale test with the Oldershaw column, it was decided to build a 30
tray column in order to determine, among other things, whether higher
residence time would make any difference in column performance.

Project Schedule

Preliminary engineering and development of a definitive cost estimate by Raytheon
was started in November 1991 and was completed on January 31, 1992. Following
DOE approval of the definitive cost estimate, Raytheon proceeded with final equipment
and facility design and procurement.

96-8

-------
Based on a competitive bidding process, Quackenbush Inc. of Buffalo, New York was
selected to construct the pilot plant and to install the necessary equipment.
Quackenbush initiated site work in October 1992. Pilot plant construction started
shortly thereafter. Construction and equipment installation were completed in March
1993. Figures 8, 9,10, and 11 show various views of the complete pilot plant.

A Dunkirk Station Unit No. 4 annual outage occurred between March and April 1993.
This outage, along with DOE and Niagara Mohawk approval of the pilot plant Safety,
Failure Mode Analysis, and Operating Manuals, resulted in a delay in pilot plant startup
until July 1993. As is shown in Figure 12, the pilot plant test program schedule
consists of a one-month shakedown test during July 1993, followed by approximately
three months of parametric optimization testing. Following completion of the
parametric testing, a five-month long-term reliability test program will take place which
should be completed the end of March 1994. After completion of the test program,
the site will be restored to its original condition.

Test Program

The primary objectives of the pilot plant demonstration are to establish the
performance and costs of the Tung FGD Process when treating the flue gas from a
medium-to-high-sulfur coal-fired utility boiler. The test program is divided into two

phases:

•	Parametric testing - to establish the capacity, performance, and turndown
capability for each component in the system.

•	Reliability testing - to demonstrate the process performance after extended
periods of continuous operation.

Parametric Testing

The parametric testing immediately following the completion of the system startup
program. Testing consists of performance analyses for the following subsystems and
operating parameters as listed below:

•	S02 Absorption - establish the performance of the venturi/absorber system and
confirm that system removes more than 90% of the S02 from the inlet flue gas.
These tests are being done at 1 and 2 MW equivalent flue gas flow rates.

•	Extraction - evaluate performance of 1, 3, and 5 stages of extraction at 1 and 2
MW equivalent flue gas flow conditions. Study aqueous and organic
entrainment over a range of aqueous/organic phase ratios.

•	S02 Steam Stripping - evaluate stripper performance as it relates to the number
of trays in the stripping column and the steam feed rate.

•	Organic Regeneration - analyze performance of proprietary components at
various operating conditions.

The results of the parametric test program will be used to establish the optimum
operation conditions for the long-term reliability test program.

96-9

-------
Reliability Testing

The long-term reliability testing will begin as soon as the parametric testing has been
completed. The purpose of this test period is to demonstrate the performance
capability of the process during continuous, 24-hour/day operation. Equipment
reliability and chemical regeneration efficiency will be analyzed during this five-month
program. Test conditions will be changed periodically to gain additional information
regarding the impact of different equipment configuration impacts on performance over
extended operating periods. Various steam stripping and regeneration configurations
and flow rates will also be evaluated during this test period.

Laboratory Analyses

An on-site laboratory is providing analytical date on organic loading, process stream
compositions, organic viscosity, and stripping and extraction efficiency. Analytical
equipment includes a gas chromatograph, titration equipment, a viscosity meter, pH
meters, and other chemical testing equipment The date generated in the laboratory is
added to the data acquisition system data base to provide system performance results
for all operating conditions. A consistent sampling protocol is being used to allow the
laboratory data to be accurately combined with process operating conditions in order
to produce an overall performance evaluation for each set of test conditions.

Test Results

At the time when this paper had to be submitted for publication (July 2, 1993), the pilot
plant was in the initial stages of startup. Test data generated during July and August
of 1993 will be presented at the FGD Symposium.

Conclusion

To date, the Tung FGD Process has been successfully demonstrated in laboratory-
and bench-scale tests. It is currently being demonstrated at a nominal 2 MW pilot
plant-scale using a slipstream of actual flue gas from a coal-fired utility boiler.

Ultimately it is hoped that a full-scale facility will be built based on the anticipated
successful results from the pilot plant testing that is now in progress.

96-10

-------
Cleaned	Spent	Loaded

Flue Gas	Scrubber	Organic

Liquor	Solvent

Liquor	Solvent

Figure 1. Tung Gas Desulfurization Schematic

96-11

-------
-------
r

to

0>
CO

Figure 3. Tung Pilot Plant General Arrangement Drawing

-------
SULFUR LOADING-MOLES SULFITE PER MOLE ORGANIC

Figure 4a. Equilibrium Vapor Pressure of SO2 Over Loaded Organic

TEMPERATURE (deg C)

Figure 4b. Loaded Organic Decomposition Kinetics
Conversion Time Versus Temperature

96-14

-------
1M6

3J> 2.1 22 Z3 Z* £6 2.8 Z7 £8 2M 9X> 3.1 &2 3.3 &4
1/TEMPERATURE {lOOOMefl K)

Figure 5. Rate Constant Versus 1/Temperature

Figure 6a. Stripping Column SO2 Removal Versus L/G

Figure 6b. Steam Stripper Test - Steam Consumption

96-15

-------
ORBIT FLOW DESIGN
Figure 7. Sieve Tray Design Alternatives

98-16

A

-------
Figure 8. Process Storage Tanks

Figure 9. Absorber Recycle Pumps

96-17

-------
Figure 10. Stainless Steel Absorber, Bleed Filter, and Motor Control Center

Figure 11. Insulated Process Heat Exchangers

96-18

-------
Uffll "" "• • 			i KJOBlMWiro		""'•"X'-lBBmBBBWrWWm			

7UNBR3DTE5T FACILITY-PHASE il j

RATCON RESEARCH & DEVELOPMENT, INC. i
ti&AKEROAO ¦*- aawmaiatTt
SHOOKUNE.IMCK46 I

7" tLtMKT 1*

eoas i tuvor

« 0WT1M ^ I ,	 1

wtia ( fffi J f¥M

14 wwann
conpucm

1

II | 9 U j F j U A m 4 4 A i a j 0 | K 1 S J f m I A It j J | J -tiwjojujDijjF |m .*>

a. |t.
mm 1 4C*ttU.

tvommjm ^

¦*"** miwT

puu ryiw -7Tf«i

1



i







ntfj . hmm'

, ,UTUM

T7







j







:ST







1







TASK 4

watfer>ujuaB»Ai6orrf

HtlESTatE STATUS REPORT

ELBfEHT

CODE

PUNNED PLANT DESCRIPTION



ELEHEHT
CODE

PLANNED PLANT DESCRIPTION



Raveon Task 1 started



iC



_ _

Issue Phase I! Work Plan



40

Issue Contractor Soscifi cation for Bid

„„

Issue Phase II Mark Plan fRev. 1)



4E

Start Construction

1A

Task 1 - Raytheon Subcontract Hark Started



4F

Start Shakedown (Begin C/Q & S-U
Activity]

!3

Pre'pet. Coot 1 eteri



4S

Carolete Construction

ic

Issued Phase II Final Renort



4H

Finish Shakedown (Mechanical CcmnlRtionl

2A

laboratory Work Corns! eted



41

Test Facility Removed

3A

"¦"ask S - Paci J itv Oesfon Standard



SA



3E

Definitive Eaiiipment/Frograni Budget
Completed and Safin Engineering
(1st Ravcon Oesien Review)



53

Test Completed

3C

?nd Ra*con Desian ievipw



-------
TABLE 1

LABORATORY SCALE STRIPPING DATA

Steam Consumption

F Value = v vp

% of S02 Stripped

6.3 lb/lb S02

0.32

93%

4.0 lb/lb S02

0.23

97%

1.8 lb/lb S02

0.11

65%

1.0 lb/lb S02

0.06

42%

96-20

-------
INTEGRATED FLUE GAS TREATMENT BY WET FGD
OPERATING IN A WATER-CONDENSING MODE

J. P. Heaphy
J. C. Carbonara
Consolidated Edison Company
of New York
4 Irving Place
New York, New York 10003

W. Ellison
Ellison Consultants
4966 Tall Oaks Drive
Monrovia, Maryland 21770

Abstract

Results are presented for electric-utility, field testing of novel, heat-recovery type, wet
flue gas desulfurization facilities of ultra-high efficiency in removal of S02, S03 and
solid particulate matter. The focus of this technology development and
demonstration-scale, continuous performance testing is an upward-gas-flow, water-
cooled, condensing heat exchanger equipped with teflon-covered tubes and tubesheets
that also provides a basis for simultaneous NOx absorption in an oxidation/absorption
FGD process-mode while serving to reduce trace metal emissions in keeping with
objectives of new Clean Air Act, Title 1H. Hie mechanical design of this advanced
flue-gas cooling/scrubbing equipment is based on more than twelve years of
commercial application of such units for energy recovery by boiler make-up water
preheating in residual-oil-fired boiler operation. This Integrated Flue Gas Treatment
(IFGT) technology is additionally proposed for use in coal-fired boiler service,
typically positioned downstream of a high-efficiency electrostatic precipitator as in
conventional wet FGD practice.

97-1

-------
Introduction

The technical development by Con Edison described herein has the unique aim of
demonstrating a new technological approach fen* power plant flue gas heat recovery
and stack emissions reduction. This technology employs equipment known as a
condensing heat exchanger, and the first condensing heat exchanger demonstration at
Con Edison began in February, 1992. A condensing heat exchanger extracts generally
unavailable energy from the flue gas stream by lowering the temperature of flue gas
well below the sulfuric acid dewpoint temperature. Conventional flue gas heat
recovery equipment foregoes use of "waste teat" in the flue gas to avoid sulfuric acid
formation and corrosion that results when the flue gas temperature is lowered below
this acid dewpoint. These limitations do not exist with the condensing heat exchanger
design now being demonstrated because all gas-side tube and shell surfaces are
covered with a layer of teflon resin. Therefore, ultra-low flue gas exit temperature
can be achieved by utilizing such a condensing heat exchanger. Considering that the
heat lost through discharge from the stack is one of the major contributors to boiler
inefficiency, substantial stack temperature reduction and the attendant heat recoveiy
can reduce fuel consumption considerably. In addition, environmental benefits are
significant since a condensing heat exchanger system can be designed to serve as a
"scrubber" in that a substantial amount of water vapor in the flue gas may be
condensed to provide a unique counter-current interaction with the gas. By successful
demonstration during this Con Edison project, condensing heat exchangers can be
shown to have the potential to become standard boiler equipment for improving plant
efficiency while serving as the key component of a stack gas treatment system for
integrated flue gas cleanup.

Prior Use of Teflon-Covered

Condensing Heat Exchangers in Heat Rate Improvement

The mechanical design of this advanced flue-gas processing equipment is based on
more than twelve years of commercial application of single-stage units in downward-
gas-flow design/operation solely for enhanced energy recovery, e.g. in preheating of
boiler makeup water or combustion air in residual-oil and natural-gas fired service.
Con Edison has now pioneered commercial, electric-utility scale application of the
process.

industrial Boilers
Overview.

The Teflon-coated condensing heat exchanger, although new to electric utility service,
has been applied in the industrial sector for more than twelve years. There are
currently more than one hundred such condensing heat exchangers that have been

97-2

-------
installed in the U.S. recovering energy from boiler flue gases and other stack gases.
The use of corrosion resistant, Teflon-covered condensing heat exchangers allows
previously wasted energy to be utilized increasing plant efficiency. Companies such
as IBM, Exxon, Union Carbide, ADM, Bristol Meyers, and Anheuser Buseh have
installed Teflon-covered condensing heat exchangers to improve the efficiency of their
steam-generating operations. The corrosion proof design of the condensing heat
exchanger system is achieved by covering all gas side surfaces with DuPont Teflon.
Teflon has been used in chemical service for over forty years, its main application
being Teflon-lined pipe. Just as the Teflon lining protects chemical plant piping form
acid attack, the Teflon covering on the outside of the heat exchanger tubes protects
them from corrosion. The conventional criteria for minimum allowable metal
temperatures for common, exposed-metal, heat recovery equipment designs therefore
do not apply. Inside the shell and tube heat exchanger, the tubes are covered with a
0.015 inch (0.38mm) thick extrusion of DuPont Teflon on the outer surface. The
shell is protected with sheet Teflon 0.060 inch (1.52mm) thick. The tube to tube-
sheet seal is a dynamic Teflon-to-Teflon seal. The outlet plenum and discharge duct
are typically fabricated of fiberglass reinforced plastic (FRF), which, like Teflon, is
inert to add attack (1).

Tube Preparation.

The fabrication process for a Teflon covered heat exchanger tube is similar to that for
insulating a copper wire. Clean copper/nickel tubing is fed through an extruder
crosshead where molten Teflon flows evenly around the tube's circumference under
high temperature and high pressure. The thickness of the coverings, typically 0.015
inch (0.38mm) is precisely controlled by varying extruder speed relative to the tube
feed rate. After the Teflon is in place, the tubes are cooled in a water bath to set the
material. Next, a doughnut shaped, sine-wave spark tester is passed over the tube to
check for flaws in the covering. If flaws are found, the tube is rejected and the metal
tubing is recovered. If not, it is installed directly into a heat exchanger being
assembled to avoid possible damage during storage. The end result of this painstaking
manufacturing process is a covering that is far superior to conventional sprayed-on or
dipped coatings. The Teflon covering has a recommended upper temperature limit of
400F (204°C) but temperatures up to 500F (260°C) can be tolerated as long as there is
fluid flowing through the tubes to limit the Teflon temperature. Studies indicate that a
thin Teflon covering typically reduces the heat-transfer ratio of the tubing by less than
10%. Unlike exposed metal tubes, Teflon does not become fouled, nor does it retain
a continuous surface layer of water. Thus, the overall impact on heat transfer rate is
negligible. (2)

97-3

-------
Electric Utility Boilers

Heat Recovery.

Favorable heat recovery performance has been achieved for more than a year in an
initial electric utility application of this technology at Con Edison's 74th Street
Cogeneration Plant with 0.3% sulfur residual oil firing (3). Two, in-parallel, single-
stage exchangers with a total flue gas flow capacity of 320,000 lb/hour, one third of
average plant flow, reduce gas temperature to less than 100°F (38°C). By extracting
additional energy from the flue gas, boiler makeup water is preheated approximately
40°F (22°C), with the same fuel input. By doing so, an additional 23,OCX) lb/hour of
steam is made available for district supply as a result of reduction in the amount of
steam required in the boiler deaerator to heat the water to its boiling point. The
cooled gas is intermixed with the balance of boiler gas yielding a stack temperature of
250°F (121°C). Performance attained verifies the potential for integrating the
condensing heat exchanger into various power plant cycles. A realistic goal for heat
rate reduction (4) when boiler condensate is heated is:

•	Combined cycle plants operating at greater than 40% cycle efficiency:
approximately \xk% heat rate reduction

•	Reheat boilers at approximately 35% cycle efficiency: approximately 4%

•	Nonreheat boilers at approximately 30% cycle efficiency: approximately 6%

•	MSW (municipal solid waste) plants at 20% cycle efficiency: approximately
11%.

Flue Gas Cleaning Effects.

For the most efficient design for heat transfer, downward-gas-flow heat exchanger
arrangements, such as at the 74th Street Plant, have long been applied at industrial
plants. In all such applications in residual oil service significant removal of flue gas
pollutants has been observed through the substantial dissolved and suspended solids
content of condensate wastewater. However, experience with upward-gas-flow
installations indicates significantly greater flue gas cleaning effect with this latter
physical arrangement, the result of the effect of counter-currency of flow, gas vs.
descending liquid condensate, on the shell side. Such field observations have now
pointed to an improved integrated flue gas treatment (IFGT) technology in which;

•	A maximum degree of heat recovery and of flue gas cleanup may be achieved

•	A downward-gas-flow, first-stage heat exchanger, used primarily for flue-gas
sensible heat recovery and conditioning of flue-gas pollutants, is followed by an
upward-gas-flow, second-stage heat exchanger serving to recover latent heat
awl as a wet scrubber that treats water-saturated inlet flue gas.

97-4

-------
Innovative, Multi-Stage Use of Indirect

Heat Exchange for Simultaneous Flue Gas Cleaning

Process Rationale

The focus of this integrated flue gas treatment (IFGT) technology development is an
upward-gas-flow, indirectly water-cooled, condensing heat exchanger (the second of
two exchanger stages) fitted with acid-proof, teflon-covered tubes and tubesheets. It
provides a unique condensing (non-evaporative) wet-scrubbing mode to address air
toxics control objectives of new Clean Air Act, Title in. Moreover, it is also capable
of simultaneous S02 and S03 removal. Advantageous trace-metal
condensation/nucleation/agglomeration along with substantially enhance boiler
efficiency is accomplished in the IFGT system by use of boiler makeup water or an
alternative heat side. Boiler flue gas is indirectly cooled to a near-ambient-
temperature, low-absolute-humidity, water-saturated state. Applied in an optimal
manner via the two-stage arrangement, (EFGT), as per Figure 1, condensing heat
exchangers possess unique potential to remove S02 S03 HC1, sub-micron flyash and
trace heavy metals, volatile organic compounds and other toxics that may soon be
regulated under the new Clean Air Act (5). Thus, with beneficial use of the recovered
heat, this may be the first stand alone scrubbing process with simultaneous removal of
multiple flue-gas pollutants that actually improves the energy efficiency of a scrubbed
power plant.

System Design/Operation

The crux of this new technology (IFGT) is the inclusion of two stages of condensing
heat exchange designed to recover useful energy and to cool the flue gas to near-
ambient temperature. Flue gas is passed first through a downflow Teflon-covered
condensing heat exchanger that transfers energy to condensate from the hotwell pump
discharge in a closed cycle power system, to boiler makeup water, process water, or
some other sink for recoverable heat. The flue gas is cooled in the downflow heat
exchanger to a temperature slightly above the water vapor dewpoint, principally
removing sensible heat. The flue gas leaving the first heat exchanger is brought to
water saturation by sprays in a fiberglass transition chamber. These sprays cool the
flue gas to its wet bulb temperature, remove any coarse particulate matter and further
condition the gas for efficient removal of S02, S03, and other pollutants. The water
saturated flue gas then enters the upward-gas-flow Teflon-covered heat exchanger. As
the saturated flue gas contact the cold tube-surfaces of the heat exchanger, additional
energy is recovered and substantial condensation of water vapor and condensable trace
metals in the gas stream occurs.

97-5

-------
Flue-Gas Pretreatment.

In the case of applications with high ash-content fuel such as coal, the use of an
electrostatic precipitator, fabric filter, or mechanical collector is used to avoid a
massive solid particulate loading entering the EFGT system. In high-sulfur fuel
service, additional pretreatment of the flue gas prior to the JFGT may include flue-gas
NO to N02 conversion, e.g. in the rear boiler cavity, to accomplish IFGT operation
that achieves removal/absorption of NOx in conjunction with S02 removal.

Dry Bulb Temperature Reduction by Indirect Heat Transfer.

The first IFGT stage is a downflow Teflon-covered heat exchanger that cools the gas
by removing sensible heat. It also conditions the gas for subsequent effective removal
of raw-gas S03. The flue gas leaving the first stage is cooled and saturated by the
action of water sprays in a fiberglass transition chamber. These sprays not only cool
the flue gas to its wet bulb temperature, but also removes coarse particulate matter
and further conditions the gas for efficient removal of fine particulates, SQ2 and S03.

Spray Fogging and Water Saturation.

Hie water saturation of the flue gas in the transition zone from the first to second
stage provides an opportunity to apply fogging sprays serving to augment
nucleation/removal of small/sub-micron particulate.

Counter-Current Wet Scrubbing With Condensation.

As the water saturated inlet flue gas contacts the cold tube-surfaces
of the second stage heat exchanger, additional energy is recovered and substantial
condensation of water vapor occurs. This subcooling of the flue gas, accompanied by
water vapor condensation, creates a large amount of fine water droplets that form on
both pre-existing and newly-formed submicron solid particles and sulforic-acid
droplets. This supports a unique condensation-scrubbing action that makes possible
substantial collection of fine particulate matter. It is captured by impaction on large-
droplet and wetted-tube surfaces with nucleation of ultra-fine droplets by continuous
gas temperature reduction in the second stage heat exchanger. These sub-micron
droplets thereby coalesce and agglomerate to larger droplets that are removed by
gravity, by impaction on tube surfaces and by mist elimination. Hie addition of an
alkaline reagent with recirculation of the IFGT liquor allows the particulate removal to
be accompanied by simultaneous flue gas desulfurization. Soda ash provides a water-
soluble, clear-liquor, sulfite/bisulfite scrubbing medium compatible with the exchanger
construction and with capability for high S02 removal efficiency. When used in
conjunction with NO to N02 gas conversion technology, IFGT has the capability for
simultaneous absoiption of S02 and NOx. In both low and high-sulfur service,

97-6

-------
important potential exists for S02 removal without chemical reagent through the effect
of catalytic oxidation.

Mist Elimination.

The mist eliminator, which provides final cleanup of gasborne wetted particulate,
removes the small droplets (down to 10 micron size and smaller) from the gas exiting
the exchanger for return to the process liquid inventory. The inertial mist eliminators
used in large-capacity wet scrubbing systems for boiler flue gas operate on the basis of
gravitational (dynamic) forces generated through change in gas flow direction. For
maximum efficiency a horizontal gas flow arrangement, employed for the final
eliminator stage, uses elevated gas velocity through vertical, vaned elements. Phase
separation chambers on fee trailing edges of the vanes extending counter to gas flow
direction afford drainage paths external to the gas-flow regime for removal of the
collected liquid without its reentrainment by the high-velocity gas flow between the
vanes.

Stack Discharge.

It is foreseen that unique innocuous exit-gas properties will allow atmospheric
discharge, as is, without complication in wet stack system design and operation
encountered with conventional high-gas-humidity, wet scrubber operations. Exit-gas
humidity is expected to be very low, (approximately 0.03 lb water vapor per lb dry
gas at 85°F i.e. 29°C). This alleviates problems usually associated with condensation
of moisture on stack internals that often require costly gas reheat equipment. An
effective mist-eliminator system will further serve to minimize presence of liquid in
the stack and thereby preclude downpitching of the plume due to the thermodynamic
effect of internally-created, plume chilling (6). Additionally, because of the
anticipated efficient removal of principal flue gas pollutants, adverse impact on
ground-level air quality is expected to be avoided despite no significant plume rise.

Waste Management

In a low-sulfur applications, wet chemical treatment of the excess-condensate purge
stream (for outfall discharges) by lime generates waste gypsum particles that enhance
flocculation in the wastewater treatment operation and reduce outfall-liquid trace metal
concentrations to acceptable levels. Collected trace metals are rendered in a low-
volume, trace metal/gypsum sludge that would be disposed of in a secure landfill.

Based on extensive, favorable experience in the U. S., calcium sulfite cake from a
high sulfur application, intermixed with the dry fly-ash catch from upstream of the
heat exchanger system (and a small proportion of hydrated high-calcium time additive)
advantageously forms a pozzolanic fixation product that can be used to build a

97-7

-------
structural landfill. Such a disposal-mass structure has permeability so low that, in
perpetuity, it remains non-water-saturated and thereby cannot be the source of
significant outflow of leachate to any underlying groundwater body. This is of
particular significance and a major environmental benefit in exploiting the gas cleaning
performance of the heat exchanger system since:

•	Heavy metals, chlorides, etc., collected in the condensing heat exchanger and
contained in the solid and liquid phases of the sulfite cake waste are
permanently encapsulated/isolated in the structural landfill mass.

•	A substantial portion of excess condensate yield and its dissolved solids is
assimilated as surface moisture in the sulfite cake waste.

•	Sludges formed in treatment of the net excess condensate yield may be
conveniently discarded by blending into (interspersing with) the fixed landfill
mass during the cake-waste/fly-ash mixing step.

Alternatively, high-sulfur applications may use the Thioclear FGD process which
employs a recirculating, clear-liquor, magnesium hydroxide medium and yields
gypsum byproduct usable in large-volume gypsum wallboard manufacture. As in die
low-sulfur mode, above, FGD liquid effluent outfall is chemically treated to render
hazardous air pollutants in concentrated form as a low-volume, disposable sludge
solids and achieve a permitted discharge to an available receiving stream.

Performance Objectives

Simultaneous Pollutant Removal,

Ultra-low emission of diverse pollutants is foreseen via IFGT use, even with poorest
quality, residual fuel stocks. Although not generally recognized as such, residual fuel
oil is a waste product, the bottom-of-the barrel remnant of gasoline and distillate fuel
production from diverse petroleum resources. Much international residual production
is available only in medium/high sulfur form and is increasingly being restricted in use
by worldwide initiatives for S02 emission reduction. Depressed price of low-grade
heavy oil, both petroleum and bitumen based, makes it a uniquely attractive energy
source in operation of utility boilers equipped with advanced, state-of-the-art, flue gas
cleaning facilities. Thus, gas cleaning performance in IFGT use is aimed at gaining
futuristic, low, pollutant-emission levels in firing of available, minimum-cost, heavy
oil fuel.

Flue-Gas Heat Recovery.

The Teflon-covered condensing heat exchanger is an indirect-contact heat exchanger in
which both latent heat and sensible heat are transferred form the flue gas to a heat
sink. Typical heat sinks available at sufficiently low temperature are boiler makeup

97-8

-------
water, process water return, condenser condensate-water and combustion air. In a
typical application for the condensing heat exchanger in industrial or commercial
service, cold water is heated with flue gas to reduce steam consumption. For
example, preheating boiler makeup water prior to the deaerator decreases the steam
requirements associated with deaeration. By reducing this steam use, overall
efficiency of the steam generation operation is increased. Ideal use of IFGT exploits
available heat sinks to gain energy cost savings in the most cost-effective manner.
The most favorable heat-recovery application of a condensing heat exchanger system
is for a steam send-out (district steam supply) power system that uses a high rate of
boiler makeup water (e.g. raw water at 60°F to 70°F, i.e. 15.5 to 21°C), as well as
high heat-rate power systems such as those employing non-reheat boiler units. For
example, a condensing heat exchanger system installed at the Rochester District
Heating Co-op in Rochester, New York, heats 100 percent of the makeup water used
for steam sendout, increasing plant energy efficiency by 7 to 10 percent with an
investment payback within approximately 2 years.

Field Pilot Plant Testing to Date of

Gas Cleaning by Integrated Flue Gas Treatment

Field test work has been carried out in residual oil fired service at two plants, one an
industrial site, (Morgan Linen Company), the other a steam sendout facility of Con
Edison at its Ravenswood Generating Plant in Queens, New York.

Morgan Linen Company

Test Objectives.

Condensing Heat Company Coiporation has carried out a program for source emission
testing on a factory-pilot-scale. This Integrated Flue Gas Treatment facility has also
been installed at Morgan Linen Company, Menands, New York. Testing was
performed to determine emissions of solid particulate matter, sulfur dioxide (SOa) and
sulfur trioxide (S03). The primary purpose was to generate emissions data to provide
preliminary quantification of the gas cleaning efficiency of the unit in medium/high
sulfur fuel service.

Field Test Operations.

Morgan Linen Company has a 14,000 pph rated boiler producing an average of 7,000-
8,000 pph of steam at lOOpsig. The fuel burned in the boiler is 1.5% sulfur, No.6
residual fuel oil. This pilot IFGT System is of a capacity that treats a 5060 pph
slipsteam of this boiler exhaust flue gas at 320°F, i.e. 160°C (approximately 65% of
total boiler flue gas). The 5060 pph of flue gas was extracted from the existing stack
by the IFGT system's induced draft fan. The downflowing flue gas then passed

97-9

-------
through the first Teflon-covered condensing heat exchanger, transferring heat to the
water in the tubes. Hie flue gas was cooled from 320°F to 113°F, which is near the
water dewpoint temperature of the gas. Upon exiting the downflow heat exchanger,
the 113°F (45°C) flue gas entered the fiberglass reinforced plastic (HIP) spray
chamber/plenum where it was contacted by sprays of a reagent solution. These sprays
quench and water saturate the flue gas at 105T (40.5°C). The sprays promote liquid
to gas contact allowing dissolved reagent (sodium-base alkali) to begin the process of
S02 removal, resulting in a water-soluble reaction product. The saturated flue gas
enters the second (upflow) Teflon-covered condensing heat company and is cooled to
88°F (31°C) as sodium liquor recirculation completes the S02 removal operation. As
the saturated flue gas contacts the cold tube surfaces of the heat exchanger, water
vapor condenses. The condensation creates a mist of nucleating water droplets that
contribute to capture of particulate. Scrubbing liquid flow due to condensation in the
upflow teat company is augmented by the recirculating-liquor spray nozzles positioned
above the tubes. The gas passes through a mesh type mist eliminator of high
efficiency prior to discharge to the FRP stack.

Overview of Test Results.

Results of the test program for this particular boiler indicate that the IFGT system
removal efficiency for total particulates averaged 89%. Hie average system removal
efficiency for S02 and S03 during tests was 99.1 and 99.3%, respectively. The
efficiency of removal of various metals is believed to be very high and will be
quantified in follow-up testing at Ravenswood. A perfunctory test was performed with
methanol injection in the rear cavity of the furnace to convert NO to N02 for removal
by the alkaline IFGT liquor. Confirmation of simultaneous SO2/N0x removal was
achieved. Further technical development and field testing work will be carried out to
achieve higher particulate removal efficiency and to quantify metals removal
efficiency.

Ravenswood Steam Sendout Plant
Overview.

Preliminary Ravenswood test work was carried out in early 1993 in anticipation of a
late 1993 formal field program at the residual-oil-fired, Ravenswood *A"-House
Steam Sendout Plant. This program is to address the field operation and testing of a
recently installed Integrated Flue Gas Treatment (IFGT) system supplied by
Condensing Heat Exchanger Corporation with capacity to treat 25,000 lb/hour of
boiler flue gas. This system will ultimately serve as a proof-of-concept,
demonstration facility for heavy-oil fired service achieving substantial heat recovery as
well as integrated stack gas cleaning functions. Testing will establish environmental
pollution control performance with respect to principal toxic chemical (hazardous air

97-10

-------
pollutant) emissions, including nickel, in oil firing both so far as stack gas cleaning as
well management of collected wastes, liquid and solid. Other pollutant species
targeted for removal from flue gas include total particulates, (condensable in addition
to non-condensable), sulfur dioxide, sulfur trioxide (NO and NOJ, as well as
hydrogen chloride.

Focus of Long-Range Test Program.

This will demonstrate:

•	Heat recovery benefits and efficiency improvements significantly decreasing
plant rate

•	Potential to burn less costly, high-sulfur fuel (in lieu of low-sulfur) without
increase in SO* and S03, emissions

•	Marked reduction in air toxics (hazardous air pollutants) and S02 emission

•	Potential major reduction in emission of sulfuric acid mist (S03/H2S04)

•	Elimination of visible acid-mist discharge and, except for winter months,
absence of visible steam-plume discharge

•	Avoidance of visible plume occurrence in highly populated localities during
warm-weather months

•	Capability to use existing stacks for flue gas discharge, avoiding need to erect
new stacks for discharge of the low-humidity, water-saturated flue gas

•	Lessening of NOx emission.

Collaboration with Brookhaven National Laboratories.

Field test activity is being augmented by pilot plant testing at Brookhaven National
Laboratories in conjunction with a test combustor of capacity less than 1 million
Btu/hour. This collaboration provides a versatile means of assessing a broad range of
operating conditions including enhanced condensate flux generation per unit gas flow
volume with the aim of evaluating anticipated improvement in sub-micron particulate
collection.

Preliminary Field Test Results.

In low-sulfur oil fired operation flue gas testing has shown that:

•	Cooling, scrubbing and nucleation effects achieve, without benefit of
optimization of system design/operation, stack particulate loadings as low as
0,005 lb/MM Btu, one-sixth of the 1979-enacted New Source Performance
Standard (NSPS) for coal-fired boilers.

•	Mercury removal efficiency exceeds that for dry-bottom coal-fired boilers
served by a high-efficiency ESP and prescrubber-equipped wet FGD system,

97-11

-------
shown by field testing to be approximately 50%.

• Removal of S03 and sulfuric add mist averaged 66%, substantially greater than
the 30 to 50% removal level typically achieved in conventional, utility wet-
FGD scrubbing operations. IMs is expected to increase to over 90% with
alkaline reagent use as in the Morgan Linen Company tests reported above.

Influence of Key Variables in System Design /Operation

Pre-Coo!ing

The initial indirect cooling of the raw, low-dust, flue gas in the first exchanger stage
takes place in a protracted manner that provides advantageous conditioning of S03 gas
by its slow cooling down to and below the sulfuric acid dewpoint temperature. Such
gas its treatment provides an advantageous means of promoting growth of acid mist
particles formed in advance of small-micron particulate collection in the condensing
regime provided by the second stage exchanger.

Atomized Sprays

Use of fogging sprays generates ten micron water droplets immediately downstream of
the first-stage exchanger. It is believed to provide a significant mechanism for
improvement of vapor droplet ami particle agglomerating effects available within the
condensing regime of the second stage exchanger. Fogging sprays are augmented by
conventional coarse-particle sprays to ensure total water saturation of flue gas
upstream of the second-stage exchanger.

Flux Rate of Water Condensation From Flue Gas

Hie heart and essence of the technology as it is expected to be broadly/universally
applied is the ultra-low-temperature, final condensing heat exchanger stage. Favored
with water-saturated inlet flue gas to which to apply the condensation scrubbing, it
provides a mechanism for ultra-low emission of sub-micron particulate and of
condensing gaseous contaminants that produce the sub-micron particulate. Note that
sub-micron control means i.e. polishing superior in performance and/or attractiveness
to conventional scrubbing methods, is afforded even when the heat-sink (cooling water
supply) temperature is as high as 80 or 90°F, (27 or 32°C), or more. Subcooling of
the flue gas In the second-stage exchanger accompanied by water vapor condensation,
serving as a unique scrubbing "flux", creates a large amount of condensate that forms
on both pre-existing and newly formed submicron solid particulate matter. This
droplet formation and accompanying gravity flow of Equid in a counter direction to
the gas flow create a unique condensation-scrubbing action that utilizes intimate
dropwise and tube surface, gas-to-liquid contact, making possible substantial collection
of fine particulate matter. Capture is enhanced by impaction on droplet and tube

97-12

-------
surfaces with benefit of condensation caused by gas temperature reduction in the
second stage heat exchanger. After droplet formation on sub-micron particulate,
resultant matter coalesces and agglomerates to large droplets that are removed by
impact and through separation in the mist eliminator.

The principal means of increasing design condensation-flux rate is by upward
adjustment of the gas wet bulb temperature at the exit of the first-stage exchanger,
thereby providing increased absolute humidity in the water-saturated gas stream
entering the second stage exchanger.

Con Edison Late-1993 Test Program at Ravenswood Station
Test Objectives

The field work to be done at Ravenswood will specifically assess the degree to which
Con Edison's purposes ami objectives in implementing this technical development have
been fulfilled.

•	Verify that addition of this treatment system to utility-scale oil fired boilers will
significantly improve the economics of steam and electric power generation

•	Characterize the performance of Integrated Flue Gas Treatment equipment of
Condensing Heat Exchanger Corporation in improving plant energy efficiency
and in minimizing air quality impact of the boiler

•	Devise and demonstrate optimal design and operation of the process in cost-
effectively increasing energy recovery, and maximizing flue gas cleaning

•	Establish gross and net pollutant emission levels and determine the fate of
pollutants collected

•	Evaluate boiler fueling alternatives that include replacement of high-priced,
imported, 0.3% sulfur residual fuel oil, commonly fired by Con Edison, with
reduced cost, higher-sulfur, heavy oil fuelstock available from within the
Western Hemisphere

•	Collect data for use in utility and other commercial-scale system designs

•	Report the results of the testing and demonstration program to the electric
utility industry.

Test Program Overview

Con Edison is continuing to closely monitor and scrutinize the results of ongoing field
operations and testing of the IFGT system installed at Ravenswood. The pilot plant
scale demonstration unit will be fully evaluated for heat recovery and stack gas
cleaning functions in oil fired service. It will assess thermal heat recovery
characteristics, including parasitic energy and consumables use, as well as
environmental management aspects, including stack gas cleaning and management of

97-13

-------
collected wastes. Detailed test evaluations will be made utilizing EPRI PISCES air
toxics program test protocol. This includes commercially-representative system
operation reflecting performance in a high-sulfur oil firing mode with the objective of
simultaneous removal of particulates including sulfuric acid mist, trace metals and
condensables; as well as applicable acid gases: S02, N02» N0X) and HC1.

The field work will be conducted so as to optimize the results and benefits of the
IFGT technology for use in the electric utility industry.

Test results will be evaluated to ascertain:

•	Heat recovery benefits and system efficiency improvements as measured by the
extent of net overall plant heat rate reduction

•	Extent of achievable reduction in the emissions of air toxics, S02,N0X and
particulates

•	Ability of the equipment to suppress both sulfuric acid mist emission and stack
opacity

•	Potential for use of existing stacks for flue gas discharge

•	Degree of avoidance of a visible steam-plume in discharge of the water-
saturated exit gas.

Differentiation Between High and Low Sulfur Fuel Service

The test work will pursue two alternative IFGT process/design strategies preliminarily
defined as follows:

•	The high-sulfur mode, e.g. in firing of 2.5% sulfur heavy fuel oil or other
high-sulfur, low-grade residuum, with the purpose of:

—	Greatest fuel cost savings through use of IFGT

—	Utilizing high-sulfur flue-gas chemistry, i.e. moderate flue-gas (yS02
concentration ratio, permitting practical use of either of two "clear liquor"
wet scrubbing means (employing oxidation inhibition) for S02 removal: (a)
Thioclear, magnesia-buffered lime scrubbing process (Dravo Lime Co.)
yielding usable gypsum, or (b) sodium alkali scrubbing with conversion of
the S02-catch to usable byproduct, (Aquatech or Wellman Lord Process,
etc.)

—	Exploiting of high flue-gas SOj/NOx concentration ratio, permitting
simultaneous NOx removal (using augmental NO to N02 gas conversion).

•	The low-sulfur mode, e.g. in firing of 0.3% sulfur residual fuel oil:

—	Significant S02 removal possible in a zero-alkali-reagent type of scrubbing
(using recirculated scrubbing liquor, if justified) by virtue of trace metal
catalysis, accompanied by effective S03 gas removal in a low-pH, water-
condensing, process regime, forming dilute sulfuric acid waste

97-14

-------
—	Favorable trace metal removal at low (e.g. 1.0) scrubber pH

—	lime treatment of the excess-condensate purge stream (for outfall discharge)
generating suspended gypsum particles that enhance flocculation in
wastewater treatment operation for reduction of outfall trace metal
concentrations to permissible levels

—	NOx removal achieved by furnace injection of ammonia (using urea reagent)
with excess ammonia (ammonia slip) emission curbed by furnace exit
injection of methanol, or efficiently absorbed in the low-pH IFGT scrubbing
step

—	Minimal byproduct volume: A low-volume, toxic, tract-metal/gypsum
sludge, (directly comparable to that formed in treatment of excess-
condensate generated in the high-sulfur mode), is the only waste product.

Application to Coal-Fired Boilers

IFGT technology can be applied in coal-fired boiler service with or without provision
for recirculating an alkaline scrubbing medium in the second-stage heat exchanger.
Scrubbing liquor recirculation affords flue gas desulfurization (FGD) and deNO, (by
N02 absorption where applicable). The IFGT system is positioned downstream of a
high-efficiency electrostatic precipitator as in conventional wet scrubbing practice.

Overall System Integration

The most technically and economically attractive means of applying/adapting this
scrubbing/heat recovery equipment for use at coal-fired emission sources is by
combining FGD performance with its unique trace/heavy metal removal capability,
(particularly with regard to improved collection of mercury). Current or improved
means/technologies for control of diverse, targeted, acid-gas pollutants including S02
and NGX, can be designed into the second-stage heat exchanger operation. Note that
in the case of volatile particulate emissions such as mercury and selenium, the
pronounced reduction of flue gas temperature in the condensing exchanger step results
in substantial condensation of water contributing significantly to the anticipated high
efficiency of removal of these trace metals.

Plant Heat Rate Improvement

Successful achievement of energy cost savings objectives in condensing heat exchanger
use is tied to:

Nature of Available Heat Sinks.

Hie type of heat sink, the amount of energy that can be transferred to the sink, and
the variations of sink capacity with time are important elements in establishing both

97-15

-------
the technical and economical Justification for condensing heat exchanger use.
Examples of possible heat sinks include:

—	Makeup water, which offers the best choice for heat exchanger effectiveness,
with a combination of the lowest incoming temperature and the highest teat
capacity available

—	Condenser hot-well condensate that can be preheated upstream of the first LP
(low pressure) feedwater heater: This will decrease the extraction steam flow
to the LP feedwater heaters and increase the LP tuibine-generator output.

There will be an increase in heat loss in the condenser. However, the
recovered energy from the flue gas exceeds this and there is a net energy gain
in the cycle.

—	Combustion air, which of these three offers the poorest heat-transfer sink: In
general, utility-size boilers are equipped with air preheaters, and further
preheating is generally not as attractive as preheating boEer condensate or
makeup water. The disadvantages of gas-to-air condensing heat exchangers
include smaller heat-transfer coefficients resulting in less-effective heat transfer;
larger equipment to accommodate both the heat-transfer surface requirements
and the large volume of combustion air; larger and often lengthy duct work; or,
alternatively, paired heat exchangers, which lower ducting costs but decrease
heat-transfer effectiveness and raise equipment cost. Such combustion air
preheat increases the temperature of air entering the air preheater, increasing
the plate temperature and decreasing air preheater duty and its recovery of
energy.

Heat Availability.

The flue-gas flow rate, temperature and composition define the energy available for
recovery and flow, along with heat sink conditions, are the primary variables
influencing the condensing heat exchanger design.

Value of Recovered Heat.

The value of energy-use being avoided by heat recovery determines the value of the
heat recovery. The steam characteristics of the steam, use of which is being reduced
by a condensing heat exchanger installation, has a direct bearing on fuel savings. For
example, if 200 psig steam that could be "sent to the street" has been used to heat
makeup water in the deaerator, reduction in use of that steam is more valuable than if
water heating had hypothetically been by 5 psig exhaust steam.

Medium/High Sulfur Service
Adaptability for Use as FCD Means.

97-16

-------
The condensing heat exchanger facilities can be employed in medium/high sulfur,
bituminous-coal applications requiring enhanced collection of volatile and non-volatile
particulate together with simultaneous, high-efficiency S02 removal. SOj emission
control is achieved by augmental use of principal existing/emerging flue gas
desulfurization (FGD) process means incorporated into the second stage heat
exchanger system. Reduction of stack gas wet bulb temperature by unique indirect
gas cooling advantageously reduces the gas flow volume for FGD design process.

Use of Throwaway Waste Design.

Because of the substantial presence of coal-based toxic substances in the pollutant
catch of the heat exchanger system, FGD design, as ultimately influenced by CAA
Title HI, can be expected to have greatest overall environmental control effectiveness
if it is predicted on throwaway-waste system operation that assimilates all collected
wastes, solid and liquid, in an all-in-one solid waste used to construct a structural
landfill. This, of course, calls for calcium alkali, i.e. lime/limestone, use so as to
yield a comparatively insoluble chemical reaction product, calcium sulfite and/or
calcium sulfate. This is desirable because it can be landfilled as a solid waste as per
common electric utility practice in the U.S. Because of the very substantial internal
equipment surface-area present within the second-stage exchanger in which a major
portion of S02 mass transfer is accomplished, calcium sulfate (gypsum) surface-scale
control/prevention is highly critical. This greatly favors use of magnesia-buffered
lime (calcium oxide containing 2 to 4% active magnesium oxide) reagent with
elemental sulfur oxidation-inhibitor additive to ensure calcium-sulfate unsaturated
mode operation. (Common wet limestone scrubbing provides a calcium-sulfate
supersaturated mode of operation and is best applied through use of open spray tower
scrubbers devoid of internals that may accumulate gypsum scale without
foreshortening boiler operating campaigns). By this means, high, e.g. 98%, S02
removal efficiency can be foreseen with production of a predominantly calcium-sulfite,
fine-grained precipitate that will dewater by thickening and filtration to cake solids
containing a high proportion (approximately 50%) of surface moisture.

Low Sulfur Service

Focus on Air Toxics Removal.

As in the case of low-sulfur residual oil firing, Con Edison's present liquid fuel mode,
the IFGT system lends itself to meeting new Title III objectives in control of air toxics
emissions in low sulfur coal fired applications. This system would be installed (along
with an upstream-ESP), downstream of the air preheater with or without use of FGD
provisions. In either case, hydrogen chloride, also included in the new Title HI
listings as a toxic substance, is removal at very high efficiency.

97-17

-------
IFCT Operation Without Alkaline Reagent Feed.

Absent use of alkaline reagent supply, such as when formalized FGD operation is not
mandated (including some medium/high sulfur applications), substantial S02 removal
is nonetheless foreseen in IFGT operation. Particularly in an augmented oxidation,
wet-scrubbing mode with recirculation of condensate/liquor, trace-metal catalysis,
including oxidation of flue-gas S02 to S03 and uniquely high S03 removal efficiency,
even at low pH, by tihe water-condensing regime fixes a significant portion of the flue-
gas-S02 as dilute sulfuric acid byproduct-waste. Moreover, low-pH scrubbing
operation is expected to provide superior trace metal collection from flue gas (than
IFGT system operation in the alternative intermediate-pH mode, as in nominal
medium/high sulfur service.)

Management of Collected Wastes.

Lime treatment of the excess-condensate purge stream, (for outfall discharge),
generating waste gypsum particles form die waste add, will enhance fiocculation in
wastewater treatment operation designed to reduce outfall trace-metal concentrations to
permissible levels. But management of major volumes of sulfurous waste/byproduct
volumes is avoided. A low-volume toxic, traee-metal/gypsum sludge, (directly
comparable to that formed in the treatment of wastewater discharge form a
medium/high-sulfur operating mode,) is the only solid waste product.

References

1.	1. Heaphy, J. Caibonara, A. Kressner, J. F. Carrigan, and W. Ellison,

"Integrated Flue Gas Treatment System for Simultaneous Emission Control and
Heat Rate Improvement - Demonstration Project at Ravenswood", presented at the
"International Symposium on Improved Technology for Fossil Power Plants - New
and Retrofit Application", Washington, DC (March 1993).

2.	"Utility seeks to integrate heat recovery flue-gas treatment". POWER, May 1993,
pp. 65,66, and 68.

3.	"Boilers, Combustion Systems, and Their Auxiliaries". POWER, Special Report,
June 1992, p. 82.

4.	S. M. Cho, D. Dietz, M. Kandis, J. Carbonara, J. Heaphy, A. Kressner and J.
Carrigan. "Heat Rate Improvement with Condensing Heat Exchangers", presented
at the Fifth International Power Generation Exhibition and Conference, Power-
Gen '92, Orlando, FL, (November 1992), Figure 6.

5.	"Developments to Watch." POWER, June 1992, pp 103 -104.

6.	R. S. Scorer. Air Pollution. Oxford. U.K.: Pergamon Press Ltd., 1968, pp
86-106.

97-18

-------
RA /% t	I •	f IPH/^Nf11 A	|

gure 1. Schematic of IFGT System

-------
A

-------
10-MW DEMONSTRATION OF THE ADVACATE
FLUE GAS DESULFURIZATION PROCESS

L. R. Lepovitz
C, A. Brown
Radian Corporation
8501 North MoPac Blvd.
Austin, TX 78759

T. E. Pearson
J. F. Boyer
ABB Environmental Systems

1400 Centerpoint Blvd.
Knoxville, TN 37932-1966

T. A. Burnett
V. M. Norwood
E. J. Puschaver
Tennessee Valley Authority
17A Chestnut Street Tower
1101 Market Street
Chattanooga, TN 37402

C. B. Sedman
U.S. Environmental Protection Agency
Air and Energy Engineering Research Laboratory
Research Triangle Park, NC 27711

B. Toole-O'Neil
Electric Power Research Institute
3412 Hillview Avenue
Palo Alto, CA 94304

98-1

-------
This paper has been reviewed in accordance with the U.S. Environmental
Protection Agency's peer and administrative review policies and approved for
presentation and publication.

Abstract

A 10-MW ADVAnced siliCATE (ADVACATE) pilot plant was built and operated in
1992 under a cooperative agreement between the Tennessee Valley Authority (TVA)
and ABB Environmental Systems (ABBES), and supported by the U.S. Environmental
Protection Agency (EPA) and the Electric Power Research Institute (EPRI). The
ADVACATE pilot plant was incorporated into TVA's National Center for Emissions
Research (NCER) located at the Shawnee Fossil Plant in Paducah, KY. Flue gas was
diverted from the existing pilot spray dryer inlet into a new duct and then returned to
the inlet of the existing pilot electrostatic precipitator (ESP). The ADVACATE process
blended an enhanced sorbent with reaction products and additional water and fed this
solids mixture into the flue gas where cooling of the flue gas, reaction with sulfur dioxide
(SQz), pneumatic conveying, and flash drying of the solids occurred.

The ADVACATE process was studied at stoichiometric ratios of 1.0 to 1.6 and at an
approach-to-adiabatic-saturation temperature of 20° Pn while treating flue gas from the
combustion of a 3.0 % sulfur coal. The sorbent preparation results, SC^ removal
efficiencies, and particulate control performance for the process are presented and
discussed.

Introduction

ADVACATE, a sorbent for removing SQ, from flue gases, was first developed and
patented by the University of Texas, Acurex, and EPA in 1986 *, In 1987 ABBES
developed an *in-duct" flue gas desulfurization process followed by a pulse-jet baghouse
called Moist Dust Injection (MDI) at its 0.3 MW pilot plant located at the University of
Tennessee. The results from a parametric test program conducted in 1988 and 1989 led
to the conclusion that an enhancement to the sorbent was needed to improve calcium
utilization. ABBES tested the ADVACATE sorbent during the parametric test program
and found a substantial performance enhancement over the use of lime2. This finding
was consistent with the trends in performance reported by EPA and the University of
Texas at laboratory scale. Consequently, ABBES obtained an exclusive license for
ADVACATE, Subsequent ADVACATE patents3,4 were issued in 1991 and 1992 and
encompassed slurry preparation, humidification, and gas/solid contact including the MDI
process. Therefore, the remaining text will discuss ADVACATE as a process as well as
a sorbent

This paper focuses on the first pilot-scale, continuous process testing of the ADVACATE
process in an MDI configuration with the solids mixer located on the flue gas duct at the
NCER from May through September 1992. The NCER site was ideally suited for the
ADVACATE and MDI process research needs because the site was of sufficient scale-up
size; utilized an ESP; was currently conducting flue gas desulfurization (FGD) research;

(I) Conversion factors to convert to metric units are included at the end of this paper.

98-2

-------
and had all of the necessary analytical, numerical, and control capabilities. Previous
spray dryer and ESP research performance data provided an excellent data base for
comparison. The site was folly staffed to support continuous ADVACATE operations
and testing.

The ADVACATE process performance results for SO? removal and particulate control
performance with an ESP are presented for the 5 months of actual testing. The results
from ADVACATE slurry preparation tests are also presented.

ADVACATE Process Description

The ADVACATE process employs the reaction of lime and fly ash in water at elevated
temperatures to produce calcium silicate compounds. These compounds show increased
reactivity toward SOj and acid gas species over that of lime alone. ADVACATE solids
are dried by blending the ADVACATE slurry in a pug-mill-type mixer with additional
dry reaction products to create a mixture which is free-flowing with relatively high
moisture content In the pilot plant configuration, the mixer feeds this blended material
into the flue gas.

The flue gas conveys the moist solids material through the ductwork to the particulate
collection device. While the solids are being conveyed, flash drying occurs and the flue
gas temperature is decreased. The flue gas SOj is simultaneously absorbed by the
reaction with the moist ADVACATE sorbent

The reaction products and fresh fly ash are collected from the flue gas in a particulate
control device and stored in a silo. A fraction of this material serves as the fly ash
(silica) supply to prepare ADVACATE slurry. The major fraction of the stored solids is
blended with the ADVACATE slurry to form a free-flowing, high-moisture-carrymg
material. The remaining fraction of the reaction products is discharged as waste.

Pilot Plant Description

The 10-MW ADVACATE pilot plant is located at TVA's NCER near Paducah, KY, and
is adjacent to TVA's Shawnee Fossil Plant. A slipstream from the Shawnee Fossil
Plant's Unit 9 boiler provides the flue gas for testing at the NCER. Unit 9 is a front-
fired Babcock & Wilcox boiler with a nameplate rating of 175 MW. Unit 9 fired a
medium sulfur (4.0 to 5.0 lb SO^/MMBtu), low chloride (0.03%), eastern bituminous
coal during the ADVACATE test program. The coal was supplied by the Peabody Coal
Company from their Martwick mine in western Kentucky. Typical coal and coal ash
compositions are shown in Table 1.

The flue gas slipstream from Unit 9 is withdrawn downstream of the boiler air heater
and multidone system. The multiclones on Unit 9 remove approximately 70% of the fly
ash from the flue gas exiting the boiler such that the flue gas particulate concentration is

98-3

-------
approximately 03 grains per actual cubic foot (gr/acf). During most of the
AD VACATE tests, a fly ash injection system was used to compensate for the fly ash
removed in the multiclones. Typically, the pilot plant inlet particulate concentration
after the fly ash injection was approximately of 2.0 gr/acf.

The flue gas slipstream from Unit 9 pases through a 40-inch-diameter duct to the pilot
plant The slipstream is routed through a venturi located at the pilot plant inlet to
measure the flue gas flow rate. Downstream of the venturi is a preheater/precooler heat
exchanger, shown in the Figure 1 process flow diagram, which is used to control the flue
gas temperature. The flue gas then bypasses the existing spray dryer to the ESP via the
duct constructed for AD VAC ATE evaluation.

A mixer, located on top of the duct, blends ADVACATE slurry, additional dry reaction
products, and water before discharging the mixture into the duct The flue gas residence
time between the AD VACATE sorbent injection location and the inlet to the ESP is
approximately 22 seconds. This residence time is site specific and may exceed that
needed in a full-scale application.

The NCER ESP is designed with four electrical sections and eight gas flow passages. An
empty fifth field is not equipped with a hopper or ash collection system. The collecting
plates are 23 feet high by 92 feet long and are spaced 10 inches apart. This resulted in
an ESP specific collection area (SCA) of approximately 440 square feet per thousand
actual cubic feet of flue gas per minute (tf /kacfm), with all four fields energized at
humidified gas conditions during ADVACATE testing.

The corona electrodes are stainless steel spiral wires mounted in a rigid-frame
configuration. Both the collecting and discharge electrodes are rapped by tumbling
hammers mounted on a rotating shaft The ESP power supplies are rated at 55 kV and
200 mA in each field.

The pilot system is equipped with flue gas analyzers to monitor SOj and oxygen (00
concentrations at the pilot unit inlet, ESP inlet and ESP outlet. Flue gas static pressures
and temperatures are also measured at these locations. The ESP outlet flue gas is
monitored for opacity. The inlet flue gas wet bulb temperature is measured manually
using a wetted wick technique approximately once every 2 hours during testing. The
measured wet bulb temperature plus the desired approach-to-adiabatic-saturation
temperature are used to calculate the set point for the ESP inlet diy bulb temperature.

The solids collected in the ESP are transported via screw conveyors to a bucket elevator.
The bucket elevator feeds a solids storage silo. The silo is equipped with a fluidized
dual-pant-leg cone bottom. One of the lep feeds the dry reaction product material
through a rotaiy valve to a series of screw conveyors that transport the solids to the pug-
mill-type mixer. The other leg feeds the solids to either a reaction product slurry mix
tank or a waste slurry mix tank.

98-4

-------
The ADVACATE sorbent is prepared by mixing quicklime from a 100-ton capacity silo
with a reaction product shiny in the ADVACATE mix tank. This slurry is pumped to
the bottom of an attritor mill. The slurry overflows from the attritor to the
ADVACATE feed tank. Hie slurry then either overflows back into the ADVACATE
mix tank, completing a recirculation loop, or is pumped to the pug-mill-type mixer. A
progressive cavity-type pump is used to feed nozzles located on top of the mixer where
the slurry is sprayed and mixed with dry reaction product material and trim water prior
to feeding ADVACATE sorbent directly into the duct.

The ADVACATE feed and mix tanks are covered, insulated tanks with steam heat
exchangers located in the bottom to control the slurry temperature. Both tanks are
equipped with dual-blade agitators to keep the slurry well mixed. Another ADVACATE
feed tank was added later in the test program to provide additional slurry residence time.
Although this tank was insulated, it did not have a supplemental steam heating system
for slurry temperature control. The agitator in this tank was marginal in maintaining a
well-mixed slurry.

Results

The following sections present the SQ, removal results and the ESP particulate control
performance during the ADVACATE test program. Also included is a section
describing the results from ADVACATE sorbent preparation tests via batch and semi-
batch techniques at different slurry residence times.

S02 Removal Performance

The SOj removal results and operating conditions for all of the ADVACATE tests
conducted during the test program are presented in Table 2. Each data point represents
the average value for the duration of a test at specific operating conditions, which
includes 4 to 8 hours of test data following 12 to 36 hours of operation to achieve steady-
state conditions. Test data were collected in 15-minute average increments. The range
of 15-minute average SO^ removal data for each test is provided in Table 2.

The major operating variables were stoichiometric ratio (moles of Ca(OH)2 in the fresh
lime feed per mole of SOj in the inlet flue gas), ADVACATE slurry residence time,
ADVACATE slurry temperature, mixer solids moisture, and fly ash injection rate. All
tests were conducted with an inlet flue gas temperature of 320° F, an ESP inlet flue gas
approach-to-adiabatic-saturation temperature of 20° F, and a duct residence time of 2.1
to 23 seconds. Since only nine ADVACATE tests were completed during the test
program, statistical modelling of the data was not attempted. Therefore, these results
are compared and discussed in general terms as a function of the major operating
variables.

98-5

-------
The total system SOjj removal performance is a combination of the SO, removals across
the duct and the ESP. The data plotted in Figure 2 present the contributions from the
duct and from the ESP to total SOj removal performance. The ESP SC^ removal is
based oil the system Met SO? concentration, so the duct SO, removal and the ESP SO>
removal can be added together to obtain the total system SO* removal. In Figure 2, the
duct SOj removal ranges from 49% to 84%, while the ESP contribution to SOfe removal
is lower and ranges from 3% to 19%. Generally, the higher ESP removals correspond
with the lower duct removals. As removal in the duct increases, there is less SOj
available for removal in the ESP.

Stoichiometric Ratio. The total system SOjj removal performance results, as a function
of lime stoichiometric ratio, are presented in Figure 3. The average SO^ removal ranged
from 66% to 89% at stoichiometric ratios of 0.99 and 1.47, respectively.

ADVACATE Slurry Residence Time. Most test runs were made with an ADVACATE
slurry residence time of 2.8 to 3.5 hours. Two tests (l-AS-12 and l-AS-13) were
completed at higher residence times of 10 to 13 hours. Test l-AS-13 produced nearly
the same SQ> removal (80%) as test l-AS-3 (82% at the same stoichiometric ratio of
L25) and test 1-AS-l (80% at a stoichiometric ratio of 136). The comparison of test
l-AS-13 and l-AS-3 is confounded by a difference in solids moisture content, which
results from differing amounts of reaction product blending. Test l-AS-12 was run at a
stoichiometric ratio of 0.99, and there are no comparable tests at a low residence time at
this stoichiometric ratio. When plotted in Figure 3, the results show nearly the same or
lower SO* removal results for the two tests at long slurry residence times as other tests.
The Bnmauer, Emmett, and Teller (BET) surface areas measured for the solids in the
ADVACATE slurry feed at the longer residence times ranged from 13 to 19 meters
squared per gram (nf/g), whereas the slurry solids surface area ranged from 22 to 40
nf/g for the 2.8- to 35-hour residence time tests.

ADVACATE Slurry Temperature. Several tests were performed at different average
ADVACATE slurry temperatures with all other test conditions at the same level (20° F
approach-to-adiabatic-saturation temperature, 1.45 stoichiometric ratio, and 3-hour slurry
residence time). Hie results plotted in Figure 4 show that lower SQ> removal
performance is achieved at lower slurry temperatures. At an average slurry temperature
of 191° F, the average SOj removal performance was 89%. There was a significant
decrease in performance to 73% SOj removal at a slurry temperature of 169° F.

In the laboratory, elevated temperatures are required to promote the reactions that form
calcium silicate compoundsI'3*4. Therefore, the observed increase in SO2 removal
performance with increasing slurry temperature was expected.

Mixed Solids Moisture Content. The moisture content of the blended ADVACATE
solids from the mixer is controlled by the rate of reaction product solids addition to the
mixer. Increasing the solids moisture content might be expected to enhance SOj removal
performance, although the effect may be offset by decreasing the rate of solids injection
into the duct. Due to the limited number of tests, a definitive conclusion cannot be

98-6

-------
reached regarding the net effect Comparison of test l-AS-3 at a solids moisture content
of 7.0% and test l-AS-13 at a moisture content of 22.6% produced similar SOj removal
performance (82% and 79% total SGj removal, respectively).

Fiy Ash Injection. Fly ash is the source of silica in the reaction to form calcium silicate
compounds. Fly ash injection was used during most of the ADVACATE test runs,
because the source of the flue gas slipstream was downstream of multiclones that
removed about 70% of the fly ash from the flue gas. Initially, there was a concern that
low amounts of fly ash would limit silica dissolution and the production of calcium
silicate compounds, resulting in decreased SO^ removal performance. Therefore, during
most tests, fly ash was transferred from the multiclones to the pilot plant and
pneumatically injected into the flue gas at a rate that increased the fly ash loading to
approximately 2.0 gr/acf.

To determine if fly ash injection affected SO? removal performance, fly ash injection was
not used during two of the ADVACATE test runs. The test results plotted in Figure 3
show that fly ash injection did not have a significant effect on SOj removal performance.
More test data would aid in confirming this effect and develop an accurate explanation.

Comparison with Spray Dryer/ESP Pilot Plant. Prior to testing the ADVACATE
process at the NCER, a spray dryer/ESP test program was completed using much of the
same pilot plant equipment5. This provided a unique opportunity to compare the
performance of two processes using data from the same pilot plant while Unit 9 was
firing the same coal. The processes, however, are in different stages of commercial
development The spray dryer process is commercially installed on low- and medium-
sulfur coal FGD applications and has been tested on high-sulfur coals in several pilot
programs, while the ADVACATE process as tested at the NCER was the first pilot-scale
demonstration of this new FGD technology.

Results from the previous spray dryer tests and the ADVACATE tests are plotted in
Figure 5. Spray dryer test results at approach-to-adiabatic-saturation temperatures of
18° F and 23° F are plotted to bracket the ADVACATE results, which were obtained at
an approach temperature of 20° F. Comparison of the results shows nearly identical SQ
removal performance for the two processes at stoichiometric ratios above 1.2. The
results at a stoichiometric ratio of 1.0 show a greater difference, which could be due to a
number of factors which cannot be explained due to the limited amount of data.

ADVACA TE Sorbent Preparation

The ADVACATE sorbent preparation process evaluated at the NCER site was designed
to expand on prior laboratory and bench-scale research at the EPA's laboratories in
Research Triangle Park and at the University of Texas in Austin. ADVACATE sorbent
is prepared by reacting lime slurry and fly ash, which is collected with reaction products
in the particulate collection device, at elevated temperatures to produce highly reactive
calcium silicate compounds. During previous research efforts, the ADVACATE sorbent

98-7

-------
was prepared in a batch mode with ratios of fly-ash-to-lime of 3:1 without the presence
of reaction products. These laboratory tests showed that, under certain preparation
conditions, the slurry solids surface area could exceed 60 nf/g and the slurry and
reaction product mixture could remain relatively free flowing with a total moisture level
exceeding 30%6.

One of the main objectives of the ADVACATE test program at the NCER site was to
produce the ADVACATE sorbent on a continuous basis. ADVACATE was prepared
for the first time in the presence of reaction products and grit from the quicklime. The
fly-ash-to-lime ratios in the ADVACATE slurry mix tank were 0.18:1 without fly ash
injection and 0.58:1 at the highest fly ash injection rate.

During the test program, ADVACATE slurry solids surface area ranged from 13 to 40
nf /g. These values are encouraging since they indicate that, even though there were
significant differences between the pilot plant operations and laboratory experiments, the
trend in producing a higher surface area sorbent remained.

The BFT surface area analyses of ADVACATE sorbent solids early in the test program
produced results that, based on material produced in batches at the laboratory scale,
seemed to be low. There was concern that the 3-hour residence time in the
ADVACATE sorbent preparation system was inadequate to produce high-surface-area
solids. Therefore, separate tests were conducted to evaluate surface area as a function
of slurry residence time. The following discussions present the experiments performed
during the ADVACATE test program that were specifically designed to investigate the
ADVACATE sorbent preparation capabilities.

"Semi-Batch" ADVACATE Slurry Residence Time Tests. Two approaches were used
during the ADVACATE slurry residence time testing. In the "semi-batch" test mode,
normal operation of the ADVACATE process was discontinued by stopping the slurry
feed to the mixer along with the quicklime and reaction product slurry feeds to the
ADVACATE mix tank. The ADVACATE slurry at the beginning of the test was
prepared in the normal manner using two tanks in series with an average combined
residence time of approximately 3 hours. Three "semi-batch" tests were ran. Tempera-
ture control problems during the first two tests may have confounded the results; there-
fore, the results from only one of the "semi-batch" tests are presented.

The "semi-batch" test was started while operating at the test conditions for ADVACATE
Test l-AS-6. In this test, slurry was allowed to recirculate through the attritor, feed tank,
and mix tank throughout the entire test. A slurry sample was taken from the mix tank
each hour for 12 hours. Water was added to the mix tank several times during the test
to keep the solids concentration below 30%. Although the addition of water resulted in
a temporary reduction in slurry temperature, the temperature ranged between 188° F and
213° F and averaged 202° F during this test

BET surface area measurements for the slurry samples are plotted in Figure 6. The first
data point plotted was at 1.5 hours, because the average slurry residence time in the mix

98-8

-------
tank when the ADVACATE slurry, lime, and reaction product slurry feed streams were
turned off was approximately 1.5 hours. The first ADVACATE slurry sample was taken
as soon as the feed streams were turned off. Measured surface area increased steadily
from 28 to 65 nf /g between 3 and 12 hours. No further increase in surface area was
observed at 15 hours of hydration time. Therefore, it was concluded that approximately
12 hours of residence time above 190° F was necessary to produce greater than 60 nf /g
surface-area ADVACATE slurry solids.

"Batch" ADVACATE Slurry Residence Time Tests. In the "batch" mode for
ADVACATE slurry residence time testing, the tanks were empty at the beginning of the
test. lime was slaked in water, then reaction product slurry was added to start the
ADVACATE reactions. Although three "batch" tests were ran, temperature control
problems during the first two tests similar to those in the "semi-batch" tests may have
confounded the results. Therefore, results from the third "batch" test are presented.

Quicklime was added to water in the ADVACATE mix tank to produce a fresh lime
sluny containing approximately 25% solids, then allowed to slake for 1 hour. Reaction
product slurry containing 35% solids from the reaction product mix tank was heated to
195° F in the ADVACATE feed tank, then mixed with the lime slurry at a reaction
product-to-lime mass ratio of 15:1 to begin the test. The resulting slurry was circulated
through the ADVACATE sorbent preparation system for four hours while hourly slurry
samples were taken. After stopping the recirculation, hourly slurry samples continued to
be taken for 12 hours. The slurry temperature averaged 204° F and ranged between
194° F and 213° F during the test.

BET surface area measurements for the "batch" test slurry samples are also plotted in
Figure 6. For this test, the beginning of the hydration time period corresponds to the
time when reaction product slurry was added to the lime slurry. The surface area
increased steadily from 26 nf /g at 2 hours to 62 nf/g at 10 hours, with a negligible
further increase to 63 nr /g after 13 hours.

Based on surface area results from the two ADVACATE slurry residence time tests
shown in Figure 6, a hydration or residence time of approximately 9 to 10 hours was
required to obtain ADVACATE slurry solids surface areas of greater than 60 nf/g.
During operation of the pilot plant, the BET surface areas were actually lower with the
ADVACATE slurry residence time at 10-13 hours than at 3 hours. This phenomenon is
not fully understood.

ESP Performance

At many existing power plants, the typical fly ash ESP inlet particulate concentration is
about 2.0 gr/acf. The requirements to retrofit these plants with a FGD process will be
site specific. The particulate emissions following an ADVACATE process will be
required to: 1) meet the emission standard for the existing facility (typically 0.03 to 0.1
lb/MMBtu); 2) meet the New Source Performance Standard (NSPS) for particulate

98-9

-------
emissions (0.03 lb/MMBtu); or 3) not increase particulate emissions above the existing
fly-ash-only emission rate.

A secondary objective of the test program was to measure the effect an ADVACATE
process has on the particulate collection performance of an ESP. Very high particulate
collection efficiencies are necessary in a retrofit application using an ADVACATE
process because the ESP inlet particulate loading is typically 50 to 175 gr/acf, depending
on the inlet temperature and the design moisture content of the ADVACATE sorbent.
The higher particulate loadings of this range were used during most test runs because the
reaction product flow rate was high to reduce the moisture content of the mixed
ADVACATE sorbent

The humidifieation of the flue gas stream that occurs with the ADVACATE process
improves the ESP electrical characteristics. Since the spark over voltage, which limits
the maximum operating voltage of the ESP, is proportional to the gas density, the
decreased temperature of the flue gas can result in an increase in operating voltages.
Voltage/current characteristics for electrical fields 1 and 4 during baseline fly-ash-only
testing and ADVACATE testing are plotted in Figure 7. A comparison of the curves for
field 1 shows the effects of both increased mass loading and lower temperature. A
comparison for field 4 shows the effects of lower temperature only, since most of the
particles that would induce space charge effects are removed prior to entering field 4.
The curves presented in Figure 7 demonstrate that electrical operation was improved in
both fields 1 and 4.

The ESP particulate collection performance results for the ADVACATE tests are
summarized in Table 3. The results in Table 3 must be qualified, however, because no
attempts were made during this short test program to investigate methods to optimize
the ESP performance. At the end of the test program, two cracked insulators were
found, the collection plates were found to be misaligned, and there was significant
buildup of solids on the high voltage insulators. Although these problems could affect
the spark rate, limit the maximum achievable voltages, and reduce collection efficiency,
no reduction in the secondary power levels in each field was observed.

The average particulate emission rate for each day of testing is plotted in Figure 8. The
particulate emission rate for the ADVACATE tests ranged from 0.019 to 0.242
lb/MMBtu at SCAs of 412 to 475 ft2 /kacfm. The SCA is calculated at cooled and
humidified gas conditions, which lowers the gas volume to be treated and increases the
SCA by approximately 10% above typical untreated flue gas inlet conditions.

Figure 8 indicates a trend of increasing particulate emission rates with time during the
test program. During the first 2 days of testing, the NSPS limit of 0.03 lb/MMBtu for
particulate emissions was easily met. However, later in the program, particulate
emissions were significantly higher than the NSPS limit One data point, on September
3, 1992, was exceptionally high at 0.242 lb/MMBtu. This data point was investigated in
detail, but no reason could be found to explain why the result was so high or to indicate
that the data were not valid. On September 17, field 4 tripped off line but was re-

98-10

-------
energized. Upon start up for testing the next day, field 4 could not be energized, so
testing continued with only three fields.

The ESP performance results indicated that very high particulate collection efficiencies
(>99.95%) were obtained. There were some difficulties in obtaining representative
measurements of the ESP Inlet particulate concentration. The inlet sample probe often
plugged very quickly or could not be inserted into the lower traverse point on the
horizontal, rectangular duct. It appeared that the particulate concentration in the gas
stream at the bottom of the duct was higher than in the upper sections of the duct. Due
to the difficulty in sampling the bottom port, where the particulate concentration
appeared to be very high, the measured inlet mass loadings could be biased low. Thus,
the actual ESP collection efficiency may be higher than reported. The outlet emission
rate, however, is not affected by the measured particulate concentration at the ESP inlet.

Equipment Operation

A limited number of tests were conducted during the ADVACATE test program,
partially because of equipment and operational problems. Some of these problems were
due to using existing equipment that was incompatible with the ADVACATE process
design. However, some of the problems were specific to the ADVACATE process. The
main problems specific to the ADVACATE process were handling and conveying several
solid and liquid process streams. Corrective actions were taken to resolve most
equipment and operational problems to complete the test program.

Table 4 highlights the main problems encountered during ADVACATE testing that were
associated with the process and describes the on-site and future proposed corrective
actions. Pilot-plant-specific problems are not included in this table. The problems which
were not resolved during the test program are currently under study.

Conclusions

From the data generated during ADVACATE process testing at the NCER site, several
conclusions can be drawn:

•	The ADVACATE process SQ, removal performance achieved 89% at a
stoichiometric ratio of 1.47. SQ performance was similar to that of the NCER
pilot spray dryer/ESP under comparable conditions.

The addition of fly ash to the pilot unit flue gas did not appear to impact SO*
removal performance.

•	The continuously prepared ADVACATE slurry solids surface areas ranged from
13 to 40 nr/g.

98-11

-------
•	The ESP particulate removal performance with an inlet grain loading between 100
and 200 gr/acf was better than 99.9 %. Emission rates varied from 0.015 to 0.242
lb/MMBtu, but most were generally higher than the NSPS limit of 0.03
lb/MMBtu with an SCA of approximately 440 ff/kacfm. These results were
obtained without tuning the ESP for the high inlet mass loading. Effects of using
the ADVACATE process with an ESP are not fully known.

•	Hie effects of slurry residence time and temperature, solids moisture content, and
slurry solids surface area on SO, removal performance are not fully understood
and need further study.

Conversion Factors

1.0 lb/MMBtu = 430.5 Ng/J

1.0 tf/kacfm = 0.1968 irf/irf/s

1.0 gr/acf = 2373 g/irf

1.0 inch = 2.54 cm

1.0 ft = 0.3048 m

1.0 tf/min = 472 cnf/s

1.0° F = 0.55° C

°F = 1.8° C + 32

1.0 lb/min = 27.216 kg/hr

1.0 ton = 907 kg

References

1.	U. S. Patent No. 4,804,521. "Processes For Removing Sulfur From Sulfur
Containing Gases." February 14, 1989.

2.	B. W. Hall, C Singer, W. Jozewicz, C. B. Sedman, and M. A. Maxwell. "Current
Status of the ADVACATE Process for Flue Gas Desulfurization." Journal of the
Air and Waste Management Association. Vol. 42, No. 1, p. 103-110 (January
1992).

98-12

-------
3.	U. S. Patent No. 5,047,221. "Processes For Removing Sulfur From Sulfur-
Containing Gases." September 10, 1991.

4.	U. S. Patent No. 5,100,643. "Processes For Removing Sulfur From Sulfur-
Containing Gases." March 31, 1992.

5.	10-MW Spray Dryer/ESP Pilot Plant Test Program High-Sulfur Coal Test Phase
(Phase III) Find Report Tennessee Valley Authority, Ontario Hydro, Electric
Power Research Institute, and Kentucky Energy Cabinet Laboratory, July 1988.
TVA/OP/ED&T-88/35.

6.	C. B. Sedmau, M. A. Maxwell, W. Jozewicz, and J. C. S. Chang, "Commercial
Development of the ADVACATE Process for Flue Gas Desulfurization,"
presented at the 25th IECEC, Reno, NV (August 1990).

Table 1

Typical Peabody/Martwick Coal & Ash Composition

Coal Component % (dry basis*)	Ash Component	%

Carbon	7298

Hydrogen	4.92

Oxygen	7.652

Nitrogen	1-65

Sulfur	3.05

Chlorine	0.03

Ash	9.72

1	Typical moisture of 10.24%.

2	By difference.

SiOj

50.96

A1A

18.84

FeA

19.67

CaO

3.33

MgO

0.83

Na20

0.38

KjO

2.39

SO,

2.19

Titania

1.04

Other

0.37 2

98-13

-------
Table 2

S02 Removal Performance Test Results

Parameter

I-AS-1

1-AS-l

l-AS-2

I-AS-3

I-AS-4

l-AS-5

l-AS-12

l-AS-13

l-AS-20

Date

6/5/92

6/18/92

7/4-5/92

07/10-13/92

7/17/92

7/29/92

9/2-3/92

9/6/92

10/1-2/92

Inlet Gas Flow, scfm

20,100

20,200

20,000

19,200

20,100

20,500

21,600

20,700

21,300

Inlet Temperature, * F

322

320

320

320

320

320

319

320

319

Inlet SC% Concentration, ppm

1,833

1,767

1,587

1,669

1,741

1,714

1,711

1,718

1,677

Approach Temperature, * P

20.0

20.3

20.0

20.0

20.0

19.9

20.2

20.0

20.1

Ply Ash Injection, lb/min

0

0

9.3

9.2

9,6

9.9

10.2

9.9

10.2

Mixer Outlet Moisture, % (dried at
14(7 F)

5.7

10.4

8.3

7.0

9.5

7.6

8.9

22.6

10.7

ADVACATB Slurry Residence
Time, hrs

2.8

3.1

3.5

6.1

14

2,9

13.0

10.1

3.2

ADVACATB Slurry Temperature,
*F

183

192

166

191

131

191

203

198

188

Reaction Product Ratio,
lb Reaction Product/lb Ca(OH^

1.6

1.8

1.7

0.8

1.5

1.7

1.9

1.7

1.2

Stoichiometric Ratio, Ca/S

1.36

1.49

1.46

1.25

1.46

1.47

0.99

1.25

1.58

Total Reaction Product Ratio,
lb Reaction Product/lb Ca(OH^

76

72

77

82

70

74

60

26

40

Test Results



















Test Average SQ Removal, %
Duct
ESP

Total System

73

7
80

77
10
87

53
19
72

76
6

82

84
3

87

75
14
89

49
17
66

64
15
79

77
11

88

Range (15-min avg data pts)

76-84

84-90

66-81

75-86

83-89

86-92

59-71

76-81

84-93

Lime Utilization, %

59

59

50

66

59

61

67

63

56

Reaction Product Solids Ga Util,, %

53

59

52

58

59

59

66

61

56

-------
Table 3

ESP Particulate Control Performance for ADVACATE Tests

Test

ESP
Fields

SCA1
(ff/kacfm)

Inlet
Loading
(gr/acf)

Outlet
Loading
(gr/acf)

Collection
Efficiency

(%)

Emission
Rate2
(Ib/MMBtu)

1-AS-l, 6/5/92

4

419

152.6

0.0048

99.997

0.015

1-AS-l, 6/18/92

4

433

125.7

0.0076

99.994

0.023

l-AS-2

4

459

133.1

0.0216

99.984

0.075

l-AS-3

4

475

168.2

0.0160

99.990

0.054

l-AS-4

4

454

173.5

0.0104

99.994

0.033

l-AS-5

4

440

1293

0.0122

99.990

0.040

l-AS-12

4

419

150.8

0.0715

99.951

0.242

l-AS-13

4

414

585

0.0266

99.953

0.087

l-AS-19

3

312

963

0.0568

99.936

0.178

l-AS-20

3

310

125.1

0.0575

99.952

0.182

lSCA at cooled and humidified conditions. Before cooling and humidification, a four-field SCA would be
nominally 390 ft2 /kacfm, and a three-field SCA would be nominally 290 ff /kacfm.

2 Baffle installed in field 4 during all tests.

98-15

-------
Table 4

ADVACATE Pilot Plant Operational Problems and Solutions

Problem

Solution(s)

Solids dropout in the duct,
increasing duct pressure drop

Increase duct velocity

Add fluidization sootblowers or other devices
to promote flow

Ensure good mixing and reliable slurry flow
Add dropout hopper
Relocate mixer

Plugged slurry nozzles in
mixer

Add on-line maintenance capability and spare
nozzles1

Add monitoring capability

Non-uniform blending of
slurry and reaction product

Change mixer design1
Reliable slurry and solids flow

Erosion of expansion joints

Add shields inside duct1
Provide setback1

ESP emissions

Tune ESP operation for high inlet dust
concentration

Moist dust buildup on ESP
support insulators

Provide heated purge air to insulators

ADVACATE material
process control

Develop on-line technique to measure
ADVACATE material

1 Solution implemented during test program.

98-16

-------
Preeooler

100

as
>
o

£
m
cc
CM

O

CO

Figure 1. National Center for Emissions Research
10-MW ADVACATE Process Flow Diagram

1	12.	1.4	1.6

Fresh Lime Stoich. Ratio (mols Ca(OH)2/moi S02)

Duct S02 Removal ESP S02 Removal Total System S02 Removal

Notes: All tests conducted at 320F In, with
3% S/0.03% d coal and Mississippi lime, 20F

approach, siunytempentture >ia>F.

figure 2. Effect of Lime Stoichiometric Ratio on SO, Removal
Performance for the ADVACATE Process

98-17

-------
1	1.2	1.4	1.6

Fresh Lime Stoich. Ratio (mols Ca(OH)2/mol S02)

No Fly Ash Injection

~

3-6 hr Mix Tank Res. Time
•

Notes: An tests conducted at320F In, 20F
approach, 3% S/0.03% Q coal. Mississippi lime,
2.1 -2.3s duct l

Slurry Temp < 180F
10-13 hr Mix Tank Res. Time

figure 3. Effect of Major Variables on Overall
SO, Removal Performance

160	170	180	190	200

Average ADVACATE Mix Tank Slurry Temperature (F)

With Fly Ash injection

Notes All teste conducted at 20,000 sdm, 320F,
20F approach, 1,46-1.49 stoich, 1.5-1.8 reaction
product/lime rafo, 2,9^5 hr mix tank res time

Without Fly Ash Injection

Figure 4. Effect of Slurry Temperature on Overall
SQ; Removal Performance

98-18

-------
90

— 100
(0
>
o

E

0)

a:

CM

O 80
CO

E
©

(O
>.

03

© 60

."5

70

-





A

A

o

¦ 6 A

o

•#A

•

¦

I

t I



!

SD-23F Approach

o

1	1.2	1.4	1-6

Fresh Lime Stoichiometry (mols Ca(OH)2/moi S02)

ADVACATE - 3-6 hr Res. Time

SD - 18F Approach

A

ADVACATE -10-13 hr Res. Time ADVACATE - No Fly Ash Injection
¦ ~

Notes: A0 tests conducted at 320F in, with
394 S/0.03% Ct coal and Mississippi (me.

ADVACATE tasJs conducted at a 20F approach.

Figure 5. Comparison of Spray Dryer and ADVACATE
Pilot Plant SO, Removal Results

E

CO

6	8	10	12

Hydration Time (hours)

Residence Time Test #5 Residence Time Test #6
(Batch)	(Semi^Batch)

Notes: Tests conducted at a reaction product/
Erne ratio of approx. 15, and an average siurry
temperature between 200and 205F

Figure 6. Effect of Hydration Time on Slurry Solids Surface Area

98-19

-------


220



200



180

<



E 160

""W"







c

140

£



Jk-»

3

120

Q





100

CO



T3

80

c

o



Q

m

©



CO

40



20



0

Fly Ash, Field 1

	-E	

Fly Ash, Fiel

		

Field 4

ADVACATE, Field 1

—e—

ADVACATE Field 4
—*—

0	10	20	30

Secondary Voltage (kV)

Figure 7. Voltage/Current Characteristics for First and Fourth Electrical
Fields During Fly-Ash-Only Conditions and ADVACATE Testing

3

1

0.3

II 025

m
c
o

0.2

0.15

¦S 0.1

0.05

£
LU

©

3
O

¦c
a

CL
£L
£0
LU

-

-







¦



~

~

-



¦

%



¦







*

¦



NSPS Limit





i i

	! 	 1	

1 1

I 	 1 		

5/31/92 6/14/92 6/28/92 7/12/92 7/26/92 8/9/92 B/23/92 9/6/92 9/20/92 10/4/92

Date

4 ESP Fields 3 ESP Fields (Field 4 failed)

Notas: All tosts conduced with 3% SC.03% O
cool. WriN^lm, andtfwbaflSah ESP
fieU4ki place.

Figure 8. ESP Particulate Emissions During ADVACATE Test Program

98-20

-------
NOXSO S02/NOx Blue Gas Treatment Process:
Proof-of-Cancept Test

John L. Haslbeck
Mark C. Woods
Warren T. Ma
Scott M. Harkins
James B. Black

NOXSO Corporation

P.O. Box 469
Library, PA 15129

Project Funded By:

U.S. Department of Energy, Pittsburgh Energy Technology Center
Ohio Coal Development Office
NOXSO Corporation
W.R. Grace & Co.-Conn.

Morrison Knudsen Corporation - Ferguson Group

99-1

-------
Introduction

The NOXSO process is a dry, regenerable flue gas treatment system that
simultaneously removes 90% of the S02 and 70-90% of the NOx from flue gas generated
from the combustion of coal. Hie process was successfully tested at small scale (0.017 MW)
on high sulfur coal (2.5%) at the TVA Shawnee Steam Plant. The test results are contained
in two U.S. Department of Energy (DOE) reports.1*2

Tests of a NOXSO process Development Unit (PDU, 0.75 MW) were conducted at
the Pittsburgh Energy Technology Center (PETC) under a cooperative research agreement
between NOXSO and the DOE. Testing in the adsorber was done by continuously feeding
sorbent into a fluidized bed adsorber and collecting the spent sorbent from the adsorber
overflow. Regeneration took place in a separate batch reactor. The test results were
reported by Yeh et al in 1987,3 and by Haslbeek et al. in 1988.4

A Life-Cycle Test Unit (LCTU, 0.06 MW) was built at PETC in 1988 to test the
NOXSO process in an integrated, continuous-operation mode. The LCTU test program was
designed to determine long-term effects of the process on the sorbent reactivity and attrition
properties. The sorbent was successfully tested for over 2000 hours on flue gas. The test
results were published by Ma et al. in 1990,5 and by Yeh et al. in 1990.6

The Proof-of-Concept (POC) test is the last test prior to full-scale demonstration. The
POC test collected all of the information needed to design the full-scale NOXSO plant: e.g.,
data pertaining to materials of construction, process performance and cost, process safety,
process control, sorbent activity, sorbent attrition, heat recovery, etc. The POC plant (5
MW) is located at Ohio Edison's Toronto plant in Toronto, Ohio. Flue gas was first
introduced to the plant on November 23, 1991. The current test results and process
performance, along with a summary of process economics, are presented in this paper.

Pilot Plant Description

The NOXSO POC plant at Toronto, Ohio is shown schematically in Figure 1. A slip
stream of flue gas (equivalent to 5 MW coal-fired power) is extracted from either Boiler #10
or #11 of Ohio Edison's Toronto Power Plant. The flue gas first flows through a 250 hp,
F.D. fan, then is cooled to 160°C (320°F) or lower by spraying water into the flue gas
ducts. The cooled flue gas then enters a 3.2 m (126") diameter fluid led adsorber. The
NOXSO sorbent, 1.23 mm diameter 7-alumina beads impregnated with 5.0 wt% Na,
removes the S02 and NOx simultaneously from the flue gas as it passes through the fluid bed
adsorber. The cleaned flue gas mixes with the hot offgas from the sorbent heater, then
passes through the baghouse to remove all particulate before the gas vents to the atmosphere
through the power plant's stack. A cyclone is installed after the adsorber to recycle fines to
the bed (50% efficient on 50 micron diameter particles).

The spent sorbent in the adsorber overflows into the dense-phase conveying system
where 377 kPa (40 psig) air lifts the sorbent 24.4 m (80 ft) high to the top of the sorbent
heater, which is a 2.34 m (92") diameter, 3-stage, fluid bed vessel. A natural gas-fired air

99-2

-------
heater supplies the hot air to heat the sorbent in the sorbent heater to 621 °C (1150° F).
During the sorbent heating process, all the adsorbed NOx and a small portion of adsorbed
S02 (.01%) desorb from the sorbent. The hot sorbent heater offgas can either be vented to
the atmosphere through the power plant's stack or mixed with the cleaned flue gas entering
the baghouse. The hot sorbent in the bottom bed of the sorbent heater underflows into a J-
valve. Either nitrogen or steam can be used to carry the sorbent through the J-valve into a
1.22-m (48") diameter moving-bed regenerator. Natural gas is the first regenerant to treat
the hot sorbent in the regenerator at a temperature of 1130°F. The sulfate on the sorbent is
reduced to SOz, H2S, and sulfide on the sorbent. Steam is then used to hydrolyze the sulfide
to H2S which occurs at a temperature of 1080°F. The offgas from the natural gas treater
mixes with that from the steam treater before the combined stream enters the incinerator, in
which all the sulfur species are converted to S02.

Regenerated sorbent flows into a second J-valve which conveys it to 1.73 m (68")
diameter, 3-stage fluid ted cooler. A fan supplies ambient air to cool the sorbent. The teat
of the regenerated sorbent is recovered by the cooling air which is then used as the
combustion air in the air heater. The cool sorbent in the bottom bed of the cooler overflows
into a 1.83m (72") diameter surge tank. A third J-valve is used to transport the sorbent from
the surge tank to the adsorber.

The first two J-valves isolate the reducing environment (regenerator) of the NOXSO
plant from its oxidizing environment (heater and cooler train). The steam (for normal
operation) or nitrogen (for start-up) enters the two J-valves to carry the sorbent upward.

Since the steam is introduced at the lowest point of the J-valve, which is also the highest
pressure point between the two vessels, a steam barrier is created to prevent the mixing of
the reducing gas with air or vice versa. The third J-valve is operated using air to lift the
sorbent from the surge tack to the adsorber.

The NOXSO pilot plant differs from a commercial unit in two respects. First, since
the POC plant uses only a slip stream of flue gas from the power plant, the amount of NOx
evolved from the sorbent heater is too small to affect the NOs thermal equilibrium inside the
boiler furnace. Therefore, NO* is not recycled to the boiler as it would be in a commercial
unit. However, the ability of a coal combustor to reduce excess NO, introduced into the
combustion chamber was proven in simulated NOx recycle tests. The tests were carried out
using the Department of Energy, Pittsburgh Energy Technology Center's tunnel furnace and
227 kg/hr (500 Ib/hr) pulverized coal combustor.1,2 Additional NO, recycle tests were
conducted on a 227 kg/hr (500 lb/hr) small scale cyclone burner at the Babcoek & Wilcox
Research Center in Alliance, Ohio to obtain additional design date for the NOXSO
Demonstration Plant.3

The second difference between the POC and a commercial NOXSO installation is the
fate of the regeneration offgases. At the POC, the offgases are simply incinerated so that all
sulfur species are converted to S02 and then vented to the plant stack. In a commercial unit,
the regeneration offgases are sent to a sulfur recovery plant where the sulfur bearing gases
are converted to elemental sulfur. Depending on market conditions and regional demand,
either sulfuric acid or liquid S02 could be chosen as the end product rather than elemental

99-3

-------
suite. As the processes for making each of these end products are well established and
commercially available, they were not tested at the POC facility.

Reliability

From November 1991 through December 1992, the NOXSO pilot plant processed flue
gas for a total of 5213 hours. The first four months of operation were hampered by the
usual startup/shakedown problems encountered in a first-of-a-kind installation. For example,
the plant was shutdown the entire month of January to replace all seven grid plates in the
fluidized beds. Nevertheless, once the initial problems were solved, pilot plant availability
averaged 76% over the period of April through December 1992 (see Figure 2). This is an
excellent result for a pilot plant since the plant was subjected to frequent changes in
operating setpoints that stretched the limits of plant operability. The pilot plant operating
experience demonstrates that the NOXSO process is simple, reliable, and adjusts easily to
changing operating conditions.

Operation/Safety

The pilot plant was staffed with two operators. A minimem of two was required for
safety so that in the event of an injury to one man the other could assist the injured or call
for emergency response. In actuality, our experience is that less than one fiiU-time operator
is needed since process controls and safety systems are folly automated. Auxiliary personnel
for maintenance, repairs, Instrument calibration, etc., would be required as needed.

The NOXSO pilot plant operated for 17 months without a single lost man-hour due to
accidental injury. This safety record is a direct result of the design team's commitment to
safety from the start. Safety engineers from MK-Ferguson (design/construction general
contractor) and W.R. Grace & Co.-Conn, (sorbent manufacturer) conducted safety hazard
reviews of the plant as designed and as constructed. In addition, the plant was inspected and
approved by MK-Ferguson's Corporate Safety Officer prior to startup.

S02/NOx Adsorption

Figure 3 is a plot of S02/N0X removal efficiencies versus cumulative plant operating
hours. The data are averages computed over a minimum of four hours and a maximum of
twelve hours. The data are selected from periods in which the plant sulfur and nitrogen
oxides mass balance closures were 100 +15%. The removal efficiencies in Figure 3 vary
with time due to the fact that NOXSO process operating conditions were intentionally varied
to quantify their effect on process performance. He process operating conditions varied
were flue gas flow rate, sorbent circulation rate, adsorber sorbent inventory, adsorber bed
temperature, and adsorber inlet S02 and NOx concentrations. Also tested were two different
adsorber configurations: 1) a single-stage fluidized bed with flue gas cooling via water spray
into the ductwork approximately 90 feet upstream of the adsorber, and 2) two fluidized beds
in series with cooling via direct water spray into the beds. The vertical line in Figure 3
marks the time at which the second adsorber grid was installed. Note that both SOj and NO*
removal efficiencies improved with the installation of the second grid.

99-4

-------
Figure 4 is a plot of S02 removal efficiency versus adsorber gas residence time.

When the data are segregated into groups with essentially the same sorbent residence time,
an equation of the form, y=ax1/N, N>1, satisfactorily represents the data. This is true for
the entire database of 117 data points, although for clarity only a portion of the database is
shown in Figure 4. The correlation coefficients (i2) for the two curves shown in the figure
are 0.85 (53-59 min) and 0.89 (32-39 min).

The strictly empirical correlation is best for intermediate values of S02 removal and
short sorbent residence times when the relationship between S02 removal and gas residence
time is nearly linear. The correlation is worst for high values of S02 removal and gas
residence time, since the correlation gives no limiting value of removal efficiency, although
the actual limit is 100%.

In addition to gas ami sorbent residence time, SOj removal efficiency varies with the
concentration of S02 in flue gas inlet to the adsorber. Figure 5 shows that S(\ removal
efficiency is inversely proportional to inlet S02 concentration. The proportionality constant
(the slope of the lines in Figure 5) varies depending upon the ratio of flue gas flow to sorbent
circulation rate.

Figure 4 also shows that the two-stage adsorber consistently out-performed the single-
stage adsorber. This is seen more clearly in Figure 6 which shows the results of an identical
series of tests cm the one and the two-stage adsorber. For the one-stage adsorber, SOz
removal efficiency is shown to be inversely proportional to the flue gas to sorbent mass ratio,
all other operating variables constant as noted at the bottom of the figure. When the tests
were repeated with the two-stage adsorber, S02 removal efficiencies were higher by 5 to 10
absolute percentage points. This improvement is due to 1) better gas distribution with the
addition of the second grid plate and 2) counter-current flow of gas and sorbent so that in the
bottom bed of the adsorber partially sulfated sorbent is in contact with the highest
concentration of pollutants providing the driving force to put more sulfur on the sorbent. All
the data in Figure 6 were obtained at equal adsorber sorbent inventories, therefore the
pressure drop across the two-stage adsorber is only greater than the one stage by the pressure
drop across the second grid plate. (2-3" H20).

Figure 7 shows NO, removal efficiency as a function of flue gas to sorbent mass
ratio. As is the case with S02, NOx removal efficiency decreases in proportion to the
increase in mass ratio, all other operating variables constant. The line drawn in the figure
through the one-stage data has a correlation coefficient (i2) of 0.98. The two-stage data show
the same trend but removal efficiencies are 6 to 12 absolute percentage points higher than the
one stage. The best line through the two-sage data extrapolates to 86% NOx removal
efficiency at a flue gas to sorbent mass ratio of 4.6. The two-stage/bed spray data point
shown in Figure 7 is 93.5% NOx removal at a mass ratio of 4.6. This shows the effect of
adsorber bed temperature on NO* removal. Data obtained over an adsorber bed temperature
range of 250-356 °F show a definite trend of increasing removal efficiency with decreasing
bed temperature. Further improvement is probable at bed temperatures lower than 250°F.

99-5

-------
This trend was best illustrated in tests where the flue gas was spiked with SO* and
NOx from pressurized gas cylinders. Figure 8 shows NO* removal efficiency as a function
of inlet adsorber NOx concentration from 300-1065 ppm. This is the range of NOx
concentration that exists in flue gas from coal-fired utility boilers. All tests were run at flue
gas to sorbent mass ratio of 4.2 to 5 and total bed pressure drop of 19" H20 in the two-stage
adsorber. The date in Figure 8 clearly show that adsorber NOx removal efficiencies of 86-
88% are achievable at 917 to 1000 ppm inlet NOx using the two-stage adsorber with in-bed
water spray.

Figure 9 shows that S02 removal efficiency increases as the concentration of NOx in
the incoming flue gas goes up. This is because the S02 and NOx adsorption mechanisms do
not proceed independent of one another. In one-step in the mechanism, NO and Oj combine
with SOz on the sorbent's surface to form Na^O*, a stable compound.

Sorbent Regeneration

The NO, adsorbed in the adsorber desorbs as the sorbent is heated to the sulfur
regeneration temperature in the sorbent heater. The heater bottom bed temperature at the
pilot plant was typically controlled at 1150°F. Laboratory studies show that NOx desorption
is complete around 700°F. In the pilot test, the NOx gas mass balance between the adsorber
and the sorbent heater was consistently within 100±15% closure. At the pilot plant, the
sorbent heater offgas (700-900 ppm NOx in air) was vented to the stack. In a commercial-
scale plant, the NOx desorbed in the soibent heater is recycled to the coal burners where a
significant fraction of the NOx is destroyed. Pilot-scale tests of NOx recycle were performed
in 1992 at Babcock & Wilcox, Alliance, Ohio. The results are summarized in the fwiai
report "An Experimental Study of NOx Recycle in the NOXSO Flue Gas Cleanup Process*.3

Sorbent sulfur regeneration is achieved by contacting sorbent first with a reducing gas
and then with steam. The gas/sorbent contacting was accomplished in two sequential
countercurrent moving bed reactors. Although tests have shown that H2S, C jHg, Hj, CO,
H2/CO (synthesis gas), and CH4 are suitable regenerant gases, pipeline natural gas was used
exclusively at the pilot plant because of cost and availability.

The sulfur regeneration mechanism is a net endothermic reaction. The heat of
regeneration is equal to 17.7 kcal per gmole sulfur at (1112°F). Since no additional heat is
input to the regenerator, the sorbent's sensible heat must sustain the reactions. Tests have
shown that the threshold temperature for the reaction to proceed is approximately equal to
the auto ignition temperature of natural gas which is 1116°F. In order to maintain a
reasonable rate of reaction, the regeneration temperature must be about 1136°F. Allowing
for heat loss when transporting soibent from the sorbent heater to the regenerator, the
temperature in the bottom bed of the heater must be at least 1150°F.

The pilot plant regenerator was equipped with sorbent sampling ports at various
elevations. Sorbent samples were periodically withdrawn from these ports and tested for
sulfur content. These data were used to gauge the extent of the reaction with time and to
estimate the rate of sulfur regeneration.

99-6

-------
The regenerator may be modelled as ail integral reactor with a constant gas
concentration profile during a steady operation period. The sulfur molar balance gives the
rate of sulfur regeneration as a function of sorbent flowrate, inventory, etc. as shown below:

FsSadX = rg dw

where Fs is the sorbent flowrate, lb/hr,

rs is the sulfur regeneration rate, (lb sulfur)/hr/(lb sorbent),

Sa is the sulfur fraction of the spent sorbent,

S is the sulfur fraction of the regenerated sorbent,

X is the conversion of sulfur regeneration, X=l-S/Sa, and
W is the sorbent inventory, lb.

The slope of the plot of X versus W/(FsSa) will give the rate of sulfur regeneration.
Figure 10 shows the plot. The legend on Figure 10 indicates where the sample was
extracted: upReg at the top of the downward moving bed, gasReg at the end of natural gas
treatment and gas+stmReg at the end of steam treatment. The data were obtained during
two steady operating periods when the sorbent heater bottom bed temperature (1150 °F) and
the natural gas and steam feedrates to the regenerator were held constant, but the sorbent
circulation rate, residence time, and spent sorbent sulfur content varied.

The data in Figure 10 can be approximated by two lines, indicating that the
regeneration consists of two main reactions. The slope of the lines is the reaction rate. Hie
rate of the first reaction, the one that occurs in the upper section of the regenerator, is eight
times higher than that of the second reaction. Since the sorbent cools as it passes through the
regenerator, the change in rate might be due to the temperature drop. However, at two
different sorbent heater bottom ted temperatures, the lines associated with the second
reaction have approximately the same slope, as shown in Figure 11. Only the rate of the
first reaction is increased at the higher heater temperature. Also at X=0.6 in Figure 11, the
sulfur regeneration shifts from the first to the second reaction regardless of heater bottom bed
temperature.

The first reaction is associated with natural gas regeneration, the second with steam
treatment. The contribution of steam treatment to the sulfur regeneration can be seen in
Figure 10. Hie figure shows that steam treatment begins to work at X=0.6; however, steam
has little effect once the gas treatment proceeds to X=0.8. Furthermore, the value of X
does not change once W/FsSa exceeds 100.

The adsorbed sulfur oxides are reduced to H2S and S02 in the regenerator/steam
treater. Figure 12 shows H2S/S02 ratio plotted against die spent sorbent sulfur content.
H2S/S02 ratio increases when the spent sorbent sulfur content is less than 0.8 wt%. Once it
exceeds 0.8 wt%, the H2S/S02 ratio appears to be independent of the sulfur content. The
mean H2S/S02 ratio is approximately 0.4. For a sulfur plant, the H2S/S02 ratio is important
as it must be adjusted to 2.0 so as to have the correct reaction stoichiometry. When the ratio
is less than 2.0, it can be adjusted by reducing S02 to H2S. Alternately, sulfuric acid or

99-7

-------
liquid S02 can be produced as a by-product of the sulfur regeneration In the NOXSO
process.

Sorbent Attrition

Sorbent attrition is caused by physical and thermal stresses that come to bear on the
sorbent as it is transported through the processing loop and as it resides in the fluid beds.
These stresses can fracture sorbent beads and/or erode the surface of the beads. If the
sorbent bead becomes small enough, it can be entrained by the gas and exit the fluid bed.
Sorbent makeup is then required to maintain a constant sorbent inventory.

The rate of sorbent attrition equals the rate of sorbent makeup provided the starting
and ending sorbent inventories are equal. The sorbent makeup rate at the NOXSO pilot plant
for a 7 month period of operation is summarized in Table 1. The sorbent makeup rate is 3
PPH or 3/27,000 = 0.011% of total sorbent inventoiy per hour. This equates to replacing
the entire sorbent inventory approximately once a year. This makeup rate is slightly lower
than the makeup rate (0.016%/hr) used in previously published estimates of NOXSO process
operating costs.

Table 1
SORBENT MAKEUP RATE

Operation

Start date
End date
Flue gas, hrs

SorbentInventory

Total makeup, ibs	20,307

Sorbent lost, Ibs	- 6,415

Increase in sorbent inventory, lbs	+ 4,245

Net Sorbent Makeup, Ibs	9,647

Sorbent Makeup Rate, Ib/hr	3.00

Materials of Construction

Corrosion coupons were placed at six locations in the pilot plant. They were exposed
to 5500 hours of process operation. Physical and chemical analyses of the corrosion coupons
and of the materials used to construct the pilot plant were performed on shutdown.

The adsorber vessel and flue gas ductwork are exposed to the corrosive effects of flue
gas. The adsorber can be constructed of either carbon steel (with a corrosion allowance) or
specialty metals designed for flue gas exposure. The carbon steel corrosion coupons at the
adsorber inlet showed a corrosion rate of 3.8 mils/yr (1 mil = 0.001 inch). With a 30 year
lifetime a carbon steel duct would loose 0.114 inches of thickness. Carbon steel coupons in
the adsorber corroded at 2.1 mils/yr. Specialty metals for flue gas performed very well on

7/17/92
2/11/93
3,232

99-8

-------
the Met and showed 110 corrosion in the adsorber. 20CB3, C276, C22 INCO 625, and C4
all showed less than 0.15 mils/yr of loss. The stainless steels (316 and 304) showed
corrosion rates less than 0.5 mil/yr.

The adsorber at the pilot plant was made of carbon steel with acid-resistant linings on
the bottom shell ami the flue gas ductwork after the in-duct water spray. The adsorber
internals consisted of a Hastalloy bottom grid, a 304 SS top grid, and a carbon steel overflow
and downcomer. A 1" diameter hole in the ductwork developed 1' downstream of the water
spray nozzles after 16 months of operation. This was due to a failure of the acid-resistant
lining and a misalignment of die spray nozzle. Subsequently, spray nozzles were placed
above the fluidized bed and water was sprayed directly into the bed to cool the flue gas and
bed simultaneously. This has two advantages: 1) the acid dew point in the bed is lower than
in the flue gas; therefore the adsorber can be made of cheaper materials and 2) the flue gas
upstream of the adsorber can be maintained above the acid dew point so that specialty
materials are not required in flue gas ductwork.

Materials used in the regenerator of the NOXSO process are exposed to SOj, H2S,
H20, CH4 and trace amounts of COS and CS2 at temperatures near 1150°F. Although this is
a very aggressive environment for materials, appropriate materials have been identified for
use. HR160 showed no corrosion, and alonized 304 stainless steel showed only minor loss
(less than 0.33 mil/yr) when exposed in the regenerator environment. Haynes 556 and 446
stainless showed losses of 3 to 5 mils/yr and could be used with appropriate corrosion
allowances. In comparison, unalonized 304 and 316 SS corroded at 22 to 52 mils/yr while
carbon steel was destroyed.

The pilot plant regenerator and steam treater were made entirely of alonized 304 H
stainless steel. A vessel inspection after 5200 hours of operation showed that the alonized
material performed extremely well. The only problem was with 446 thermal overspray used
to cover the welds that had flaked off after 300 hours due to the difference in thermal
expansion between 446 and the alonized 304 H vessel. However, the base weld consisting of
a 312 root pass and 308 filler performed very well without the overspray.

The sorbent cooler was constructed entirely of carbon steel. The vessel internals
were made of types 304 or 316 stainless steel. Based on visual inspections, no corrosion or
erosion problems were encountered in the sorbent cooler.

The sorbent heater was made entirely out of type 304 stainless steel. The shell from
the top of the bottom head to the middle grid was alonized to protect against hot sulfate
corrosion. Over the life of the pilot plant, no measurable corrosion or erosion was observed
in the sorbent heater.

Economics

Data from the pilot plant have been incorporated into the design of a commercial size
NOXSO plant. Using this commercial plant design, an economic analysis was performed for
the NOXSO process. The basis for the analysis and cost information are included in Table

99-9

-------
2. The analysis was conducted for a 500 MW power plant burning 3% sulfur coal and
emitting 0.6 lb NO/MMBtu.

Since the NOXSO process is a combined S02/N0X removal process, it is not possible
to separate the cost of removing S02 from the cost of removing NO,. Consequently, an
assumption is matte that the cost of removing NO, is 3.0 times higher than the cost of
removing S02. The value of 3.0 represents a reasonable average for the relationship between
the cost of NOx and SO2 removal based on separate high efficiency technologies. This value
does not affect the overall economies, however, it does affect the relative cost of S02 and
NOx removal.

Emissions data is also listed in Table 2. Hie "Phase I S02 Limit" is calculated based
on allowable emissions of 2.5 lb S02/MMBtu. It is appropriate to consider over compliance
since fee high removal efficiency of the NOXSO process will allow a utility to generate SO*
allowances which can be sold to partially offset fee operating cost. A value of $300 has been
assumed for SO2 allowances. Beginning in the year 2000, the number of allowances
generated will decrease, however it is also likely that the value of allowances will be
significantly higher offsetting to some degree the reduction in the number of allowances
generated.

The annual operating and maintenance cost is $24.7 million with the cost of sorbent at
$10.1 million representing 41% of fee total. The capital cost of $257/kw is based on a
recent EPRI study.7

Revenues for the process will be generated by the sale of the sulfur by-product and
fee S02 allowances. The sulfur by-product can be elemental sulfur, sulfuric acid, or liquid
S02. The choice of sulfur by-product will be influenced significantly by the local demand
for the specific product. Since the market for sulfur is larger than the other two, sulfur is
used in this analysis. If fee market exists for sulfuric acid or liquid S02, either would be a
better choice since the revenue from sulfuric acid would be approximately three times more
than suite and liquid S02 would be six to eight times more. Making sulfuric acid or liquid
S02 would also result in minor increases in capital and operating costs.

The net levelized cost for the process is presented from three points of view. The
cost of buying, operating, and maintaining the plant will be $26,2 million dollars per year.
This translates to 8.5 mills/kwh of electricity produced. On a pollutant removal basis, it cost
$276 to remove each ton of SO* and $828 to remove each ton of NOx.

Conclusion

The 5 MW pilot plant has successfully shown that the NOXSO process can be scaled
up by a factor of 10. Maximum removal efficiencies of 99+% S02 and 95% NOs were
demonstrated over 6500 hours of process operation. Date from the pilot test may now be
used to design a commercial-scale plant wife an acceptable level of technical risk.

99-10

-------
Table 2

NOXSO PROCESS ECONOMIC ANALYSIS (1)

POWER PLANT PARAMETERS

GROSS CAPACITY
CAPACITY FACTOR
HEAT RATE
COAL HEATING VALUE
COALSULFUR
NOx EMISSIONS

ECONOMIC PARAMETERS

electricity

NATURAL GAS
SORBENT

NET SULFUR VALUE

S02 ALLOWANCE VALUE

FIXED CHARGE RATE (2)

REMOVAL COST NOr/REMOVAL COST S02

NOXSO PROCESS REMOVAL EFFICIENCIES
S02
NQr

EMISSIONS DATA

UNCONTROLLED S02
CONTROLLED S02
PHASE IS02 LIMIT
S02 ALLOWANCES GENERATED

UNCONTROLLED NOx
CONTROLLED NOx

POLLUTANT REMOVAL EFFICIENCY

OPERATING AND MAINTENANCE COSTS

FIXED (3)

VARIABLE (4)

NATURAL GAS
SORBENT
ELECTRICITY
TOTAL

CAPITAL COST

500 MW

mo«

10,000 Bta/kWh
12,000 Btu/lb
3.0%

0.6 Ib/MMBiu

S0.Q3 /kWh
$2J0/Mscf
$3.40®
ISO Aon
$300
10.6*

3.0

95%
80%

76.653 tans/year
3,833 ions/year
38.325 taas/yest
34,493 tonsfyear

9,198 toufyear
1.840 tons/year

93.4%

$5,714,000
5129.000
S5.131.000
510 J 12.000
$3,642,000
$24,728,000

5128400.000

$257 flcW

510347.750
5l.S20.435
$12,168,188

REVENUES

S02 ALLOWANCES
SULFUR VALUE
TOTAL

NET LEVELIZED COST

$26480.813/year

8-5 mills/kWh
S275Aon-S02
SS28Aon-NOx

(1) 1993 dollars.

(2> Based on 30 year book life, 20 year tax life, 38% composite federal and state tax.
and 2.0% for property taxes aad insurance,

(3)	Includes operating labor, fringes, aad supervision; maintenance labor and equipment;
and general add administrative expenses.

(4)	Includes process water and Clans plant catalyst.

99-11

-------
Acknowledgments

NOXSO Corporation wishes to thank the following for their commitment of personnel

and funds to the successful completion of the NOXSO POC test: The U.S. Department of

Energy's Pittsburgh Energy Technology Center, the Ohio Coal Development Office, Ohio

Edison Company, W.R. Grace & Co.-Conn, and MK-Ferguson.

REFERENCES

1.	J.L. Haslbeck, C.J. Wang, and L.G. Neal, et al., "Evaluation of the NOXSO
Combined N0,,/S02 Flue Gas Treatment Process," NOXSO Corporation Contract
Report submitted to U.S. DOE Report No. DOE/FE/60148-T5. November 1984.

2.	J.L. Haslbeck, L.G. Neal, and C.J. Wang, et al., "Evaluation of the NOXSO
Combined NOj/SC^ Flue Gas Treatment Process," NOXSO Corporation Contract
Report submitted to U.S. DOE Report No. DOE/PC/73225-T2. April 1985.

3.	J.T. Yeh, C.J. Drammond and J.L. Haslbeck, et al., "The NOXSO Process:
Simultaneous Removal of S02 and NO, from Flue Gas," Presented at the 1987
Spring National Meeting of the AIChE, Houston, Texas. March 29 - April 2, 1987.

4.	J.L. Haslbeck, W.T. Ma and L.G. Neal, "A Pilot-Scale Test of the NOXSO Hue Gas
Treatment Process," NOXSO Corporation Contract Report submitted to U.S. DOE
Contract No. DE-FC22-85PC81503. June 1988.

5.	W.T. Ma, J.L. Haslbeck and L.G. Neal, "life-Cycle Test of the NOXSO Process,"
NOXSO Corporation Contract Report submitted to U.S. DOE Contract No. DE-
FC22-85PC81503. May 1990.

6.	J.T. Yeh, W.T. Ma and H.W. Pennline, et al., "Integrated Testing of the NOXSO
Process: Simultaneous Removal of SO^ and NOx from Rue Gas," Presented at the
AIChE Spring National Meeting, Orlando, Florida. March 1990.

7.	J.E. Cichanowicz, C.E. Dene and W. DePriest, et al., "Engineering Evaluation of
Combined NO^/SC^ Controls for Utility Application," Electric Power Research
Institute, Palo Alto, CA. Dec 1991.

99-12

-------
NOXSO

PROCESS
TOWER

Figure 1. NOXSO POC at Toronto, Ohio

99-13

-------
100



Month

Figure 2. Pilot Rant Availability by Month
(December 1991-December 1992)

3.000

4,(MO	5,000

Cumulative Operating Hours

Figure 3. NOXSO Pilot Plant
S02/NOx Removal Efficiencies

6.000

7,000

99-14

A

-------
Adsorber Gas Residence Time (SEC)

Figure 4. S02 Removal VS Gas/Sorbent Residence Time
77 Data Writs Used In Non-Linear Regression

\



" \ iV



- \



\



o\

! I \

i i

1.000	1,500	2.000	2.500	3.000

S02 inlet To Adsorber, ppm

3.500

10000 pph.U- H20
[N0x)=350-<21 pfttl
—

7000t>Ctl.125'K20
(NOl]=18e-Z31 ppm
—©_

4,000

Figure 5. S02 Removal as a Function of Inlet S02
All Tests at 7000 scfm

99-15

-------
o

03

3	as

342-3W F «*» 3334M F <*» (ftp)
2£OF.250F(W6»*»9«ib«l»pn(f)

45

Rue Gas/Sorbent {lb/lb)

Figure 6. NOXSO Pilot Test
One-Stage VS Two-Stage Adsorber

100

ONE SJW3E ADSORBS?
—%r-

TWO STAGE ADSORBER

O

TWO STAGBBED SPfWY
~

Flue Gas/Sorbenl (LB/LB)

WH20 (on® stag*;. T & 7 *H20 (two Mags)
342-349 F (one stage), 333-344 F (two stage)
250 F. 250 f (MO siagtfMd Spray)

Figure 7. NOXSO Pilot Test
One-Stage VS Two-Stage Adsorber

99-16

-------
FUJ£ GAS COOLING
31M33F
#

(oo	aoo

Adsorber Met NOX (ppmw)

1,000

Rua en/Soiwnt» 4-5 ims
Mtortxr Prnsura Drop » 1PH20

Figure 8. NOXSO Pilot Test
Flue Gas Spiking with S02/NOx
{S02}=2300-2600 ppm Two-Stage Adsorber

FUJE GAS COOLING
319-333 F
*

BOTTOM BED COOLING
288-311 F

O

BOTH BEDS COOLING
251-25? f

o

200	400

F*» Gaj/Sort>«rt -4-5 KWIb
Adax&er Bad Pmssura Drop =19*M20

boo	aoo	1.000

Adsorber Iniet NOx (ppmw)

1.200

1,400

Figure 9. NOXSO Riot Test
Flue Gas Sinking with S02/NOx
[S02]=2300-2600 ppm Two-Stage Adsorber

9S-17

-------
88	tSO	ISO	200

W/(Fs*Sa), kg.sort>0nt/kg.S hr

Figure 10. Sorbent Sulfur Regeneration
(Open Symbol: 10-05-92d->10-18-92d)

(Solid Symbol: 11-2Q-92n->12-10-92n)

W/(Fs*Sa}, kg.sorbent/kg.S hr

Figure 11. Effect of Heater Bottom-bed
Temperature on Sulfur Regeneration

99-18

-------
J	I	I	I	I

0.4	0.6	0.8	1	1.2	1X	1.4

Spent Sorbent Sulfur Content, wt%

Figure 12. Regenerator Offgas H2S/S02 Ratio

99-19

-------
SOXAL™ DEMONSTRATION PROJECT AT
NIAGARA MOHAWK'S DUNKIRK STEAM STATION

David Hurwitz
Clifford Denker
Patrick Birdsall
Joseph Soltys
AlliedSignal, Inc.

7 Powder Horn Drive
Warren, NJ 07059

Peter Strangway
Niagara Mohawk Power Corporation
300 Erie Boulevard West
Syracuse, New York 13202

Abstract

Beginning in January 1993. AlliedSignal and Niagara Mohawk Power Corporation conducted a
six-month demonstration of the SOXAL™ Flue Gas Control Process at Niagara Mohawk's
Dunkirk coal-fired steam station. The Department of Energy's PETC office sponsored the
project, with additional funding provided by NYSERDA and ESEERCO. The SOXAL process is
an advanced regenerative system employing a sodium-based scrubber and AQUATECH™ bipolar
membranes to regenerate the scrubbing solution and to recover the sulfur values in a form suitable
for commercial use. Combination of the basic SOXAL process with urea/methanol NOx control
methods results in NO* removal efficiencies in excess of 90%. This paper discusses the successful
operation of a 3 MW pilot-scale demonstration over a six-month period, with continuous
operation and removal of over 98% of the SO2 and 90+% of the NO*. It also describes several
mechanical problems, not directly related to the process.

Introduction

The overall objective of the nominal 3 MW pilot plant was to demonstrate the technical and
economic feasibility of the SOXAL process for simultaneous removal of SOX/NOX contaminants
from the flue gas of a coal-fired boiler. The key demonstration was the integration of an
AQUATECH bipolar membrane system with proven sodium scrubbing and sulfur dioxide steam
stripping technology. Previously, the AQUATECH membrane system had been commercially
proven in applications unrelated to flue gas desulfurization. In addition, SOXAL was tested on a

100-1

-------
laboratory pilot unit for one year using a synthesized scrubbing solution of sodium sulfite.

The majority of the current project's funding was provided through a grant from the Department
of Energy's PETC office under a cost-sharing contract with AlliedSignal as the prime contractor.
Additional funding and the host ate were provided by Niagara Mohawk Power Corporation with
co-funding from the New York State Energy Research and Development Authority (NYSERDA)
and the Empire State Electric Energy Research Corporation (ESEERCO).

The Clean Air Act of 1990 made it necessary to accelerate the incorporation of acid rain control
systems into the planning process of some of those utilities and industrial facilities burning sulfur-
containing fuels, primarily coal. While many control systems east based on lime and limestone
scrubbing, these systems have severe drawbacks for incorporation into long-term emissions
control plans. Although the costs associated with disposal of large amounts of gypsum sludge
may be manageable today, the trend is towards more costly and less onerous disposal options. In
addition, these approaches simply take pollution out of the air and put it in the ground. Even if
economical today, limestone scrubbing does not provide an environmentally sound solution for
the long term.

Many SO2 control technologies are being pursued in the hope of developing an economical
regenerable system that recovers the SO2 as a useful commercial product, thus minimizing the
formation of waste. Some include the use of exotic chemical absorbents alien to the utility
environment and waste treatment facilities. These systems present new environmental issues to
the utility while attempting to solve old ones.

Sodium alkali scrubbing is an accepted, proven system for removing SO2 from gas streams. It is
the system of choice in industrial applications due to its lower capital requirements, high SO2
removal efficiencies (98+%) and low maintenance costs. A large number of sodium scrubbers
have been operated successfully at industrial as well as utility sites. The main drawback has

always been the higher cost of the sodium scrubbing solution versus calcium systems. SOXAL
was developed to minimize the cost of sodium scrubbing by regenerating the scrubbing solution
and recovering the sulfur in the form of commercial products such as liquid sulfur dioxide, sulfuric
acid, or elemental sulfur.

Another benefit of sodium scrubbing is its' ability to enhance existing NOx control technologies
based on urea or ammonia injection. Methanol injected as an adjunct to these systems converts
NO to NO2. The NO2 is then absorbed % the sodium scrubber providing overall NO* removal
efficiencies in excess of 90%. The SOXAL pilot unit at Dunkirk was adapted to simulate
incorporation of a urea/methanol NOx control process. NO2 will be injected into the flue gas to
approximate the composition resulting from urea/methanol injection.

100-2

-------
Process Description

The SOXAL process flow diagram (Figure 1) shows four major unit operations:

1 • Pre-scrubber - A water scrubber designed to remove chlorides and fluorides from the flue
gas as well as particulates escaping from the ESP. It also serves to quench the hot flue gas.

2.	Sodium sulfite scrubber - removes SO2 and NO2 from the flue gas.

3.	AQUATECH bipolar membrane cell stack system - regenerates the spent scrubber solution
(sodium bisulfite) back to sodium sulfite and produces sulfiirous acid for the stripper.

4.	Steam stripper - strips off" pure SO2 (98+%) from the suUurous acid feed.

The primary reactions in the sodium scrubber for SOx/NOx removal are shown below. SO2
removal is accomplished by reaction with the Na2S03 scrubbing solution to form sodium bisulfite
(NaHSCb). Note that a portion of the Na2S03 is oxidized to Na2S04 by reaction with oxygen
from the flue gas stream.

Na2S03 + S02 + H20 --> 2NaHS03

Na2S03 + 1/2 Qa — > Na2S04

The Na2SC>4 can be recovered in a salable crystalline form. Alternatively, a secondary
AQUATECH unit can be employed to convert the Na2S04 to NaOH and H2SO4. This NaOH is
recycled to the scrubber to reduce the alkali makeup requirements.

Removal of NO* by urea/methanol injection into the flue gas is accomplished in a two-stage
process. First, urea addition reduces approximately 70%-80% of the NOx in the flue gas to N2 as
shown in the reactions below. This reaction occurs at 1600-1900 °F.

NH2-CO-NH2 —> 2NH2 + CO

NH2 + NO —> N2 + H20

The remaining NO in the flue gas is oxidized to N02 by reaction with methanol. Addition of
methanol also reduces ammonia slip, thereby reducing deposition of salts in downstream ducting
and plume visibility problems. This reaction proceeds at 1000-1500 °F.

CH3OH + 302 — > 2HO2 + H20 + C02

H02 + NO —> N02 + OH

The N02 product from methanol injection is subsequently reduced to N2 by reaction with Na2S03
in the scrubbing solution. The Na2S03 is oxidized to Na2S04 in this reaction. The Na2S04
generated is eventually recovered as a solid crystalline product, or processed in the optional

100-3

-------
secondary AQUATECH unit.

2N02 + 4Na2S03 --> N2 + 4Na2S04

Regeneration of the spent scrubbing solution is achieved in the primary AQUATECH
regeneration unit using proprietary bipolar electrodialysis membranes. The bipolar membranes
separate water molecules into hydrogen (H4") and hydroxyl (OH") ions (Figure 2). NaHS03 in the
spent scrubbing solution is converted to Na2S03 as shown below.

2NaHS03 —> Na2S03 + H2S03

H2S03 —> H20 + S02 (gas)

In the primary regeneration unit (Figure 3), sodium cations (Na) migrate across the cation
selective membrane into the base compartment and become associated with OH" ions originating
from the bipolar membrane to form NaOH. Most of this NaOH reacts with NaHS03 to form the
regenerated Na2S03 for recycle to the scrubber. The HSCb" anions that remain in the acid
compartment associate with H"1" ions from the bipolar membrane to form H2SO3, The partially
saturated H2SO3 stream is continuously withdrawn from the primary regeneration unit and
subsequently decomposed into S02 gas and water in the steam stripper.

The above reactions describe what takes place in the pilot-plant system. However, in commercial
operations, the concentrated S02 gas would be passed through a cooler to condense most of the
water. The resulting gas stream would then be preheated, dried, compressed, cooled and stored
as liquid S02 product. An optional secondary AQUATECH unit comprising three compartments
can be employed to recover sodium alkali from the sodium chloride and sodium sulfate solution
from the stripper bottoms as shown below:

Na2S03 + H20 —> NaOH + H2S04

The amount of H2 S O4 production is essentially dependent on the degree ofNa2S03 oxidation in
the system. Alternatively, the sodium sulfate stream from the stripper bottoms can be evaporated
and crystallized to a salable Na2S04 crystal form.

Pilot Plant Operation

The layout of the pilot plant is shown in Figure 4. The operators of the demonstration facility
found the SOXAL process to be easy to operate and very responsive but were surprised by how
many unexpected mechanical and instrumentation problems occurred. Few issues surfaced
concerning the membranes; however, most of our time was spent fixing pumps, flow meters, weld
leaks, and assorted non-process related problems. Most of these failures were eventually
corrected. The majority probably would not have occurred if we had more experience in
choosing commercial scale components and instrumentation for scrubbing systems and if our
suppliers and contractors had conformed more closely to our specifications for materials of
construction.

100-4

-------
Instrumentation

Our greatest disappointment was the failure of the SC>2/NOx gas analyzer to operate throughout
the entire test period despite numerous attempts to fix the problems. We were forced to contract
for an outside CEM service beginning in April in order to complete the test program. Our PVC
flow sensors proved to be too fragile. When the PVC turbines were replaced by stainless steel
turbines, the problem was solved. We found our Vortek-type gas flow meters to be prone to
plugging. Pitot type devices would be recommended in the future. Initially, we also experienced
difficulties with some of the sensor electronics and various controllers. These problems, too,
were mostly solved with time.

Much of the instrumentation installed proved, however, to be extremely reliable and performed up
to expectations. These instruments included the variable area rotometers, differential pressure
level sensors, pressure transducers, pH sensors, pneumatic control valves, PLC, Data Acquisition
System and miscellaneous PID controllers.

Prescrubber/Scrubber

The prescrubber and scrubber were designed and supplied by Advanced Air Technology of
Arlington Heights, Illinois. They were chosen for their expertise in supplying sodium alkali
scrubbers to the industrial market and their ability to incorporate low oxidation into their designs.

The prescrubber measured 4.5 feet in diameter by 25 feet in height and was water based. The
sodium- based scrubber included two six-foot, packed (polypropylene-random) stages and
measured 4.5 feet in diameter by 40 feet in height.

Operation was easy and reliable. We observed no fouling although a lot of particulates were
removed by the prescrubber. The mist eliminators worked well and there appeared to be no
carry-over from the prescrubber to the scrubber. The only difficulties encountered were the
failure of one pump, some minor control problems and a few initial gas leaks.

Cell Stack/Membranes

The AQUATECH cell stack employed at the Dunkirk demonstration project employed 44 two-
membrane cells. Each cell was composed of a single bipolar and cation membrane. It required
only 176 ft2of cell area to handle the equivalent of 3 MW of flue gas.

The cell stack was operated for over three months, although not all continuously, and no
problems were experienced with the cell stack hardware. The standard commercial AQUATECH
bipolar membranes employed in the demonstration performed very well. While some individual
membranes were replaced, from time to time, the overall performance was well within

100-5

-------
expectations, A batch of experimental bipolar membranes which did not perform well early in the
test were replaced. At no time was the testing ever delayed due to membrane failure.

Stripper

The steam stripper, although representing conventional technology, presented us with numerous
problems. Most were related to materials of construction and inconsistencies that resulted in
leakage of SO2 and concerns about the durability of the column. Instrumentation Mures also
resulted in temperature control problems that gave the erroneous impression of low efficiency.
Once these problems were identified and corrected the column operated as designed.

Operation of the pilot plant has been achieved with minimal staffing. Two engineers and four
operators covered all the shifts required, including 7 days per week, 24 hour operation during
continuous testing in January and February.

Test Description

The original 3 MW pilot-plant test program at Dunkirk covered six months (Figure 5)
incorporating continuous operation and numerous parametric tests. Failure of the SOx/NOx
analyzer to operate severely limited the quantitative data collected during the first two months of
operation. During January and February of 1993 we were, however, able to demonstrate
continuous operation of the cell stack and its integration into the scrubber/stripper system. The
process was kept in balance providing regenerated scrubbing solution and concentrated SO2
which was sent to the stack in the case of the pilot-plant facility. This period of operation also
demonstrated the viability of the bipolar membranes in a sodium-based system. Unfortunately we
were unable to measure how much SO2 was being absorbed in the scrubber.

Due to a scheduled month-long shutdown of the Dunkirk Station boiler during March, we were
unable to resume testing until late April 1993. This allowed time to revise the parametric testing
to take place during a four-month time frame after boiler startup and to arrange for a CEM
service to set up to obtain accurate SO2 and NOx readings that were unavailable from the
defective S02/NOx measurement system.

Beginning in late April, we began parametric testing according to the following plan:

100-6

-------
Tablet

Test Plan for Parametric Studies

Test

•	Initial Baseline Studies

•	Absorber Parametric Studies

-	Lower Stage pH

-	Recycle Rate

-	Number of Beds

-	SO2 Concentration

•	Cell Stack Parametric Studies

-	Base pH

-	Recycle Rate
-Amps

-	Conversion

-	Cell Stack Temperature
Stripper Temperature
Optimized Baseline
Optimized Baseline
NO2 Concentration
Optimized Baseline - SO2/NO2

EDTA Addition

Objective
Establish Baseline

Maximize Absorption
Maximize Absorption
Minimize Oxidation

Minimize Oxidation/Maximize Absorption

Reduce Flush Cycle
Reduce Cost and Flush Cycle
Reduce Power Consumption
Reduce Power Consumption
Reduce Power Consumption
Optimize Efficiency
Maximize Absorption
Minimize Oxidation
Determine N02 Absorption
Continuous Operation

Reduce Oxidation

During the four months of parametric and continuous studies from April through July, the
Dunkirk boiler was shut down on weekends and was operated in a curtailed mode overnight due
to a lack of power demand. Parametric testing was run, therefore, on a decoupled basis, five days
per week In other words, when we ran studies on the absorber, the cell stack was shut down and
vice versa. We ran from a full storage tank of either spent or regenerated absorbent, depending
on the test. The part of the process not undergoing testing was ran overnight to replenish the
inventory for the next day's testing. Although the parametric studies were not appreciably
affected by the boiler cycling and shutdowns, continuous operation was affected. At no time were
we able to run the process for greater than five days due to a lack of flue gas.

Test Results

As of the time when this paper had to be written and submitted, the demonstration program was
still in progress and we have not had an opportunity to thoroughly analyze all the data collected.
The data collected to date is summarized in Table 2. During the first two months of the test, we
operated continuously, seven days per week. Unfortunately our S02/N0X analyzer was not
functioning properly and we have no definitive gas stream scrubber data. Since the testing was
restarted in April using an outside (trailer-mounted) testing service, we have consistently
demonstrated 98+% SO2 absorption (Figure 6). During this time, SO2 concentration in the flue

100-7

-------
gas has averaged between 1000 and 1500 ppm (Figure 7). We are confident that same high level
of absorption was achieved during our earlier continuous run. Preliminary data indicates that
higher SO2 levels (obtained by recycling recovered SO2 into the scrubber) enhances SO2
absorption and reduces oxidation. We will have an opportunity towards the end of the test period
to run a number of five-day continuous tests. Unfortunately, the curtailed operation of Niagara
Mohawk's boiler at the present time is preventing us from running a longer continuous test.

Oxidation in the scrubber is an important factor in the economics of the SOXAL process, when
there is no commercial use for the sodium sulfate formed. The SOXAL process allows for
recovery of the oxidized sodium values at a small additional capital and operating cost using a
secondary AQUATECH unit. The sulfur values, however, would be lost. Recent preliminary
data indicates that total oxidation, during the SO2 only (no added N02) part of the demonstration,
Ms well within the design parameters even without the use of additives or any attempt to
minimize oxidation through adjustment of the operating conditions.

The major parameters associated with the operation of the AQUATECH cell stack are the
membrane durability and power consumption. Power consumption for the cell stack has
consistently been in the range of 1100 to 1300 kWh/ton S02 (Figure 8), depending on the
parametric test performed. This corresponds with anticipated values and is consistent with the
consumption used in EPRI's 1990 evaluation of SOXAL process economics.

During over 1600 hours of pilot plant operation to date, the commercial AQUATECH bipolar
membranes have proved to be extremely durable. Some experimental bipolar membranes had to
be replaced due to delamination. A wash process is used to minimize fouling of the membranes.
Parametric testing has resulted in the identification of operating conditions that minimize the need
to acid wash the membranes.

The majority of the testing remaining to be completed at this time involves 1) additional tests at
increased SO2 levels, 2) NO2 injection, 3) continuous operation at optimized conditions, and 4)

minimization of oxidation through addition of EDTA.

Conclusions

Although testing was incomplete at the time when this paper had to be written and submitted,
some very positive conclusions had already been achieved.

1.	Continuous integrated operation of the total absorption and regeneration system has been
demonstrated.

2.	The ease of operating the absorption and regeneration systems independently has been
demonstrated. This is important in dealing with variable loading, use of off-peak power and
general maintenance, where it may be desirable to operate parts of the system separately for
a period of time.

100-8

-------
3.	Stable bipolar membrane performance has been proven,

4.	The ability to operate the process with minimal staffing has been demonstrated. AlliedSignal
1ms operated the test with a staff of two engineers and four operators covering either a three
or four shift schedule. Except during parametric testing the demonstration was run by an
operator and engineer or an operator working alone.

5.	Consistent 98+% SO2 removal. SO2 removal has been demonstrated up to almost 100%
during parametric testing.

6.	Power consumption and oxidation consistent with design expectations for SO2 removal only
operation has been proven.

The demonstration has proven to be successful to date although instrumentation and mechanical
issues have limited the amount of quantitative data available from the first two months of
continuous testing.

100-9

-------
Dirty Am Gas
S02.HC1

Purge Stream

HCI

XI

Clean
FhieSas

NaOH
Makeup

Recovered
SO,

Flue Gas
SO.

i

J

Scrubber Faad
NajSO,

Stripper Feed
HzS0>lNaIS04

Base

Acid





vy

Spent Absorbent
NaHSOj, NajSO,

Steam



Met so2

Assoesfft AifWPfP

Cm

Smac

I - Purge Stream
NajSO,

Smw
Srmm

Figure 1
SOXAL Process Flow Sheet

o

h2o

Anode

OH-

pr

W, v.;. .vVggSI

C^ivAviwS®

OifSSsiSli

h2o

OH" H+

o

Cathode

Figure 2
Bipolar Membrane Operation

100-10

-------
Sodium Sulfite
Na2SOs

--V-

T

©T© ©

T

T

^	

Sulfurous Acid
h2so3

oh-

H*

OH"

©

— SO."

H*

SO,*

Anode

Na1



Na»-

OH-

Na*

Sodium Bisulfite 4 4 4 4
N3HSO3 	'	—Cell Unit	L

1

©

Cathode

Figure 3

Primary Regeneration Unit Operation

Flue
Gas

Figure 4
Equipment Layout

100-11

-------
Figure 5
Pilot Plant Operations Schedule

100

98

¦ ¦

i ¦

c
o

o
«
¦Q
<
-------
<<
-------
Table 2



Summary of Test Data - Daily Averages









S02

Flue Gas

SO?

Power

Test



Concentration

Flow Rate Absorption

Consumed

Code

Test Description

fppm)

©SCFM)

(%)

(kWh/ton SCW

A

Simultaneous* Baseline







A

Simultaneous Baseline









A

Simultaneous Baseline









A

Simultaneous Baseline

916

7,9922

98.8



A

Simultaneous Baseline

1202



99.3



A/H1

Simultaneous with 1.25x Feed to Base

1062

5,205

96.3



A/H1

Simultaneous with 1.25x Feed to Base

1115

4,798

97.5



A/H2

Simultaneous with 1.5x Feed to Base

1149

7,562*

98.9



A/H2

Simultaneous with 1.5x Feed to Base

1189

7,3662

97.2



A/H2

Simultaneous with 1.5x Feed to Base

988

4,758

96.6



B

Absorber Baseline

861

5,201

96.8



B

Absorber Baseline

1385

5,609

98.7



B

Absorber Baseline

1365

5,201

98.8



B

Absorber Baseline

1316

5,709

98.4



C

Cellstack Baseline







1122

C

Cellstack Baseline







1251

C

Cellstack Baseline







1082

D1

Absorber Bottoms at 5:1 Ratio

1361

5,309

99.6



D1

Absorber Bottoms at 5:1 Ratio

1278

5,175

98.9



D1

Absorber Bottoms at 5:1 Ratio

1085

5,711

91.9



D1

Absorber Bottoms at 5:1 Ratio

1182

5,402

97.7



D2

Absorber Bottoms at 1:1 Ratio

1086

5,744

100.0



E2

Absorber Recycle at 60 gpm

1146

5,880

97.9



E3

Absorber Recycle at 45 gpm

1123

5,771

95.9



F1

Absorber with One Stage

1211

5,865

82.7



G

Absorber with Recycled S02

2141

5,407

99.6



G

Absorber with Recycled S02

2167

5,394

100.0



G2

Absorber w/ Recycled S02, 5:1 Ratio

2228

5,377

99.4



H2

Cell stack with 1.25x Feed to Base







1131

11

Cell stack at 80 amps/sq. ft.







1049

12

Cell stack at 125 amps/sq. ft.







1331

J1

Cell stack at 80% Add Conversion







1334

J2

Cell stack at 120% Add Conversion







1292

Notes:

1.	"Simultaneous" indicates continuous operation of both absorption and regeneration
processes. All other tests conducted are decoupled.

2.	These flow rates in ACFM. All others in Dry SCFM.

100-14

-------
OPTIMIZATION OF ADVANCED COOLSIDE DESULFURIZATION PROCESS

M. R. Stouffer, W. A. Rosenhoover, J. A. Withum, J. T. Maskew
CONSOL Inc., Research & Development
40£X) Brownsville Road, Library, PA 15129

Abstract

This paper describes the development status of a high efficiency sorbent injection desul-
furization process (Advanced Coolside). The Advanced Coolside process involves the
use of a contacting device to humidify flue gas to near saturation while removing fly ash.
Hydrated lime, injected into the highly humid flue gas downstream of the contactor, cap-
tures S02 before being removed in the existing particulate collector. The high humidity
allows high S02 removal. High sorbent utilization is achieved by sorbent recycle, the
potential for which is increased with upstream fly ash removal. The objective of this
project is to develop a low-capital-cost process for Clean Air Act compliance, capable of
over 90% S02 removal and of sufficiently high sorbent utilization to be cost competitive
with wet limestone scrubbing over a wide range of compliance situations. The original
performance targets of 90% S02 removal and 60% sorbent utilization were exceeded in
1000 acfm pilot plant operations. The 90% S02 removal target was achieved with com-
mercial hydrated lime at sorbent utilizations of 70-75%. Up to 99% S02 removal was
attained at sorbent utilizations of at least 60% A simplified equipment design was
tested and operability was confirmed in pilot plant tests. An interim economic evaluation
was completed based on these results. This economic analysis compared Advanced
Coolside with wet limestone FGD on the basis of an "ntil" plant design, i.e., as if it were a
demonstrated technology. Projected capital costs are approximately 40% lower than wet
limestone scrubbing costs over the range of coal sulfur contents (1.5-3.5%) and plant
sizes (160-500 MWe) evaluated. Levelized S02 control costs are lower than wet scrub-
bing over the range examined. The economic study identified several design improve-
ments for cost reduction. Ongoing development work will focus on these design
improvements and on sorbent improvement. The goal is to establish at least a 20%
levelized cost advantage over wet FGD to provide a sufficient incentive for
commercialization of this lower capital cost process.

Background

In-duct dry sorbent injection technology has been actively developed in the U.S. since
the early 1980s. The performance of these processes has been well-established through
the development of the Coolside process (CONSOL)1"6 and the HALT process
(Dravo)7,8 and through the DOE duct injection technology development program.9"11
These development efforts have included pilot-scale tests, proof-of-concept tests, and a

101-1

-------
full-scale utility demonstration. Established performance is in the range of 40-50% S02
removal at 2.0/1 Ca/S molar ratio and 20~25°F approach to adiabatic saturation tempera-
ture using hydrated lime as the sorbent. Additionally, the 105 MWe demonstration of
the Coolside process at the Ohio Edison Edgewater Station4*6 showed that an S02
removal of 70% can be attained by improving calcium hydroxide sorbent activity with
sodium-based additive injection at a 0.2 Na/Ca molar ratio (—32% sorbent utilization).

Process performance data and economic analyses support the attractiveness of duct sor-
bent injection for site-specific applications.12 However, the applicability as a compliance
option for the Clean Air Act or other regulations can be expanded by increasing S02
removals and sorbent utilizations. The performance targets for developing an advanced
process (90% S02 removal and 60% sorbent utilization) represent a substantial improve-
ment over previous technology.

The Advanced Coolside process is being developed using a 1000 acfin pilot plant13 The
pilot plant was used in previous development and scale-up of the Coolside process;2 it
was modified to include all elements of the Advanced Coolside process. Process devel-
opment has focused on improving the design of the contactor and on improving sorbent
utilization by optimizing sorbent recycle. A test program to investigate sorbent
improvement was recently initiated. This report will discuss progress in these areas,
results of the interim economic study and approaches for future process improvement.

Description of Advanced Coolside Process

Figure 1 show a schematic of the Advanced Coolside process. The process achieves
greater S02 removal and sorbent utilization than previous duct sorbent injection pro-
cesses by operating at a higher flue gas humidity and by more fully exploiting the
potential of sorbent recycle. The key to the process is a gas/liquid contacting device
downstream of the air preheater. The contactor serves two purposes: to nearly saturate
the flue gas with water, and to remove most of the coal fly ash from the flue gas. The
sorbent is injected downstream of the contactor into the highly humid flue gas. Hydrated
lime is very active for S02 capture near the saturation point, even in the absence of
liquid water droplets. Because the flue gas is already humidified prior to sorbent injec-
tion, there is no strict residence time requirement for droplet evaporation. S02 is
removed in the duct and by the sorbent collected in the existing electrostatic precipitator
(ESP) or baghouse. The heat of reaction between S02 and hydrated lime raises the
temperature of the flue gas by roughly 8-10 °F for each 1000 ppm of S02 removed.
Therefore, the particulate collector can be operated at an elevated approach to
saturation without flue gas reheat. However, because hydrated lime activity is highly
sensitive to the approach to saturation, this reaction heat effect also acts as a limiting
mechanism for S02 capture.

The spent sorbent is captured by the existing particulate collector as a dry powder. It
can be disposed of with the fly ash or separately. Sorbent recycle is an integral
component of the Advanced Coolside process. Laboratory and pilot plant tests have

101-2

-------
shown that recycle sorbent is quite active for S02 capture at high humidity. The
potential for recycle is increased because fly ash is removed separately before sorbent
injection. Furthermore, recycle sorbent performance can be improved by a simple physi-
cal pre-treatment step prior to re-injection; the nature of this pre-treatment step is
currently a proprietary feature of the process.

Design optimization has focused on the flue gas/water contactor. For the initial pilot
plant testing the contacting device was a Waterloo scrubber.

14,15 This js

a commercially

available device, marketed by Turbotak Technologies Inc. (Turbotak), and used primarily
for removal of submicron particles. The Waterloo scrubber consists of a conditioning
zone, a centrifugal fan and a mist eliminator, and uses two-fluid nozzles to finely atomize
water sprays at a liquid/gas ratio of about 1 gal/1000 acf.

Experimental

Hie Advanced Coolside process is being developed using a nominal 1000 acfm (0.3 MWe
equivalent) pilot plant. In addition, exploratory work is being conducted in fixed-bed
laboratory reactors.

Figure 2 is a schematic of the pilot plant. It was used in the development of the
Coolside process and in the scale up to a 105 MWe demonstration. In this develop-
ment the plant was shown to adequately simulate large-scale process performance. The
pilot plant was modified to fully simulate integrated operation of the Advanced Coolside
process, including combined flue gas saturation and fly ash removal by a contactor,
sorbent injection into nearly saturated flue gas, and semi-continuous sorbent recycle. A
baghouse is used for particulate collection. As described below, the baghouse could be
operated to simulate S02 reduction expected in an ESP. The pilot plant is described in
more detail in Reference 13.

Discussion
Recycle Optimization

The improvement in desulfurization performance which allowed project performance
targets to be exceeded resulted primarily from recycle optimization. By more folly
exploiting recycle, sorbent utilization efficiencies of 70-75% were attained, while
maintaining S02 removal around 90%. Also, high S02 removals ranging from 90% to
over 99% were attained, while maintaining sorbent utilization of at least 60%.

Recycle optimization tests were conducted in the 1000 acfm pilot plant in a semi-
continuous manner. Spent sorbent was removed from the baghouse hopper on a regular
basis; a portion of the material was discarded and the remainder, after pretreatment, was
returned in a batch to the recycle feeder hopper. Test durations were sufficiently long to
assure that steady-state continuous recycle was simulated closely (typically 20-70 hr).

101-3

-------
Tables 1 and 2 list process conditions and results for pilot recycle optimization tests.

Teste 1 through 4 (Table 1) were conducted with reheat before the baghouse (to a 25 °F
approach) and with frequent pulse cleaning to minimize baghouse S02 removal. The
purpose was to simulate conditions in a retrofit application with an existing ESP. In this
case, S02 removal in the ESP would be limited by bulk gas mass transfer. Based on
literature information and on theoretical calculations, an ESP removal of 30% of the
S02 remaining in the ESP inlet gas is a reasonable assumption. As shown in Table 1,
S02 removal in the baghouse with reheat averaged 5% (absolute), with most removal
occurring in the duct. Tests 5-9 (Table 2) were conducted with no baghouse reheat. The
9 to 12 °F baghouse approach temperature was a result of the flue gas temperature rise
from the heat of reaction. In these tests S02 removal in the baghouse was greater than
with reheat, although the large majority of S02 was still removed in the duct.

The recycle test results indicate that for systems with an existing ESP, 90% S02 removal
can be achieved at sorbent utilizations of 70-75%, substantially higher than the original
target of 60% utilization. For example, with a fresh Ca/S mol ratio of 1.2, duct and
system S02 removals were 87% and 90%, respectively (Test 2, Table 1). The increase in
sorbent utilization over the initial target has a substantial impact on economics, because
sorbent cost is a major component of the total S02 control cost.

The recycle results also indicate that very high efficiency S02 removal can be attained in
systems with a baghouse operated at a close approach to adiabatic saturation. For
example, 99% S02 removal was attained at 61% sorbent utilization (Test 9, Table 2). In
this test most of the S02 removal (88%, absolute) occurred in the duct. This capability
to achieve very high S02 removal may be attractive to new units using a baghouse for fly
ash collection. As discussed below, high removal efficiencies have been achieved at
higher sorbent utilizations (ca. 75%) by using small amounts of additives for promotion.

In the recycle tests in Tables 1 and 2, recycle ratios ranged from 3.3 to 6.9 lb/lb fresh
lime. Relatively high recycle ratios are possible because fly ash is removed upstream of
sorbent injection. Total dust loading ranged from 9.5 to 14J gr/scf. Pilot testing
indicated that recycle sorbent particles tended to aggregate during handling, pretreatment
and reaction; this is expected to improve the ability of an existing ESP to handle higher
dust loadings. For example, the mean particle size of sorbent entrained in the flue gas
upstream of the baghouse was 45 jwn, compared with a typical value of around 15 fim for
fly ash from firing of pulverized bituminous coal.

As shown in Tables 1 and 2, the recycle optimization tests were relatively long-term tests.
With one exception, operating durations ranged from 21 to 115 hr. This allowed process
operability to be evaluated. It also provides added confidence in data reliability.

Data reliability was confirmed by comparing utilizations based on gas analysis with those
based on solids analysis. Figure 3 shows a parity plot of utilizations by the two
independent methods for the tests in Tables 1 and 2 and for more recent tests with addi-
tive promotion. The good agreement of sorbent utilization data confirms the accuracy of
process performance data. It also confirms that steady-state continuous recycle

101-4

-------
conditions were closely simulated. Establishing near steady-state operation is critical in
assessing the sorbent usage of a process with a large inventoiy of recycle material. In-
duct S02 removal data accuracy was confirmed by manual flue gas sampling using EPA
Method 6 for selected ran periods. Continuous gas analyzers agreed with the manual
analyses within the accuracy of the EPA method. Duct exit gas was sampled using a
probe to minimize sorbent entrainment; the probe was also heated to minimize sorbent
reactivity.

Design Optimization

A major portion of the process development is devoted to contactor simplification. The
contactor is a key capital cost component, and the contacting device initially tested (the
Waterloo Scrubber) was designed for more stringent applications (i.e., submicron
particulate control) than required for Advanced Coolside. Because the process is applied
upstream of an existing particulate collector, some fly ash slippage through the contactor
is acceptable. It is only necessary to remove a large portion of the particulate mass (ca.
90%) to avoid recycling much of the inert fly ash. Approaches to reduce the capital and
operating cost of the contactor included eliminating the fan, an integral component of the
original Waterloo scrubber system, and redesigning the contactor to reduce water and
atomization air requirements.

Preliminary pilot plant studies indicated that these approaches are feasible. Tests were
conducted using the original Waterloo scrubber system with and without its centrifugal
fan over a wide range of atomizing air pressures and water flow rates. The test results
(Figure 4) indicated that high flue gas relative humidities can be achieved with or without
the fan, as long as sufficient water droplet surface area is generated in the contactor.
The test results also showed that the droplet surface area and, thus, atomization energy
can be reduced to below typical operating conditions with a relatively minor effect on
flue gas humidity. For example, reducing atomization air pressure from 45 psig to
25-30 psig resulted in a slight decrease in humidity, from ca. 99% to ca. 97%. Particulate
removal tests indicated that removal efficiency was not sensitive to the nozzle operating
conditions over the ranges tested and that the scrubber fan was not needed to achieve fly
ash removal greater than 90% (by wt). These results indicated that there is flexibility for
design and operating modifications.

Based on the pilot plant tests using the original contactor, a mechanically simpler con-
tactor was designed by Turbotak, marketer of the Waterloo scrubber. The new contactor
(Figure 5) consists of a spray chamber and a downstream mist eliminator. The spray
chamber uses two-fluid nozzles and an internal baffle design to promote droplet/gas/
particle interactions and a reduced velocity zone to allow droplet settling. Most of the
particles and water droplets are removed in the spray chamber. The mist eliminator
removes remaining droplets from the flue gas. The Waterloo scrubber fan was elimi-
nated in the new contactor, significantly reducing the cost.

Tests were performed which verified the humidification performance, particulate collec-
tion efficiency, and operability of the simplified contactor. Over 100 optimization tests

101-5

-------
were conducted to reduce atomization air pressure and flow and water flow relative to
the design conditions of Turbotak. Figure 6 shorn the results of these tests, plotted as
percent relative humidity vs. the relative droplet surface area. Relative droplet surface
area is defined as the total droplet surface area per unit flue gas volume produced at the
test operating conditions divided by the total droplet surface area per flue gas volume
produced at the original design conditions. Efficient humidification requires sufficient
droplet surface area together with good dispersion of the droplets in the flue gas; both
are affected by variations in nozzle operating conditions. Based on the results of S02
removal tests performed earlier, high S02 removal can be attained at relative humidities
of 95% and higher. The results in Figure 6 show that the droplet surface area can be
significantly lower than the area produced by the design conditions without significantly
affecting the humidification efficiency, provided good dispersion is maintained. Many of
the operating configurations that were tested still achieved >95% relative humidity even
though they produced droplets with as much as 50% less surface area than the design
conditions. To produce sprays with less droplet surface area, lower atomizing air
pressure and/or water flows rate is required.

Particulate removal tests showed that >90 wt % fly ash removal (measured upstream of
the mist eliminators) can be maintained with the nozzles operating at lower air pressure
and water flow rate than the design conditions. Figure 7 shows that most all of the
conditions tested gave better than 90 wt % fly ash removal and better than 95% relative
humidity.

Based on these results, alternative operating conditions for the contactor were identified
that reduce the water flow rate and atomizing pressure (Table 3). The pilot recycle tests
discussed above used these alternative operating conditions. As shown in the table, water
and air flow requirements were reduced by about half, and the air pressure requirement
was reduced from 45-50 to —30 psig, while maintaining humidification (>95% relative
humidity) and fly ash removal efficiency (> ca. 90%). These alternative operating condi-
tions will result in lower operating and capital costs (e.g., for air compressors).

Operability Observations

Pilot plant operating experience in tests up to 115 hr in duration is a positive indication
of the operability and retrofit potential of the Advanced Coolside process. Although the
pilot plant is not of sufficient scale to make a complete assessment of process operability,
observations of pilot plant operation provide initial information on key operability issues.

The contactor operability was simplified by the elimination of the fen. With the modified
design, no significant operability problems have been observed in over 1500 hours of
operation. The mist eliminator was washed periodically to maintain contactor pressure
drop at about 1.5 inches of H20.

Accumulation of solids on the duct walls was not an operating problem, even at very

close approach to saturation and with different duct configurations having short straight-
run residence times (<0.5 sec) and numerous changes in flow direction. There was

101-6

-------
generally a light surface coating of dry solids. At bends, there was somewhat more
accumulation. The amount of solids on the duct surface tended to reach a steady value
after 10 to 30 hr of operation, after which the rate of accumulation approached zero.
The solids were loose and easily removed. Solids accumulation only became a significant
operating problem if the solids injection system was oriented in such a way that the
sorbent impacted a wall directly after injection.

No major problems were encountered in preparing, handling and feeding the recycle sor-
bent. Operability of the pneumatic transport system was similar to that with hydrated
lime. A minor feeder modification improved feeding operation by reducing bridging in
the hopper. Operability of the recycle handling system was observed to deteriorate at
very high sorbent utilization (>70%). Operability was improved by adding the fresh lime
to the recycle material during pre-treatment and co-injecting the sorbents.

Baghouse operability was good at the close approach to adiabatic saturation (down to
10 °F) investigated in this program. The material did have a tendency for compaction
under compression and for hopper bridging at the lowest baghouse approach tempera-
tures. This would be an important consideration for a large-scale design.

Sorbent Optimization

Sorbent improvement can increase the attractiveness of the Advanced Coolside process
in several ways. Increasing sorbent utilization reduces sorbent usage and waste disposal
requirements. Increasing sorbent activity can reduce the required level of sorbent recycle
and could increase the applicability of the process for very high S02 removal levels.
Finally, the results of sorbent studies could allow use of lower cost sorbents by reducing
process sensitivity to sorbent source.

Pilot plant tests reported previously in this paper were all conducted with a single com-
mercial hydrated lime. A sorbent optimization test program is currently under way. The
program includes work in three areas: a lime hydration study, evaluation of alternate
sorbents, and evaluation of additive enhancement.

The objectives of the lime hydration study are to determine the effect of hydration vari-
ables on the properties of hydrated lime and to determine the effect of lime properties
on desulfurization performance. The hydration study is being conducted in cooperation
with Dravo Lime Co. using their continuous pilot (120 lb/hr) hydrator. Hydration vari-
ables being investigated in a statistical experimental design include the following:
quicklime source, quicklime grind size, hydration water temperature, residual H20 in the
product, and hydrator residence time. Hydrated limes will be characterized for chemical
composition and physical properties such as particle size distribution, BET surface area,
and pore size distribution. Desulfurization performance will be measured in laboratory
reactors and in the pilot plant.

Evaluation of alternate sorbents includes testing of different commercial hydrated limes
and testing of other sorbents, for example, specially prepared hydrated limes. Recycle

101-7

-------
tests have been conducted for three commercial hydrated limes. These hydrated limes
were obtained from different geographic regions and from lime plants among the ten
largest in the U.S. In tests at 1.2 Ca/S mol ratio, system SOz removals (with baghouse
reheat to simulate ESP removal) ranged from 86 to 90% for the hydrated limes with
surface areas ranging from 14 to 22 m2/g. Also, once-through (no recycle) screening tests
of different commercial hydrated limes from different sources and with varying surface
areas showed only small differences in S02 removals. These results suggest that process
performance is relatively insensitive to surface area and to commercial lime source. This
is an economic advantage, allowing use of the lowest cost sorbent available. Preliminary
screening tests have also been conducted on two specially prepared high surface area
hydrated limes, each with a BET surface area of about 35 m2/g. For one of these
sorbents, once-through S02 removals were higher than for any of the commercial
hydrated limes tested (by about 10-15%, relative). For the other, S02 removals were
lower than any commercial hydrated lime tested. Process performance must be verified
in recycle tests; however, these results indicate that a systematic investigation of the
effect of lime properties on desulfurization is warranted.

Previous laboratory studies7 simulating Advanced Coolside process conditions indicated
that sodium-based additives can substantially increase the utilization of hydrated lime (by
over 20% absolute). In the current test program, different approaches for additive
promotion will be investigated, including addition to lime during hydration. Based on
previous lab studies and literature information, additives to be evaluated include
Na2C03, NaCl, and CaQ2. Chloride additives are of interest because they could be
generated by neutralization of contactor recycle water. One pilot plant test was
conducted with NaCl promotion. Results are encouraging, indicating that sorbent
utilization can be increased to 80-85% without increasing recycle by using very small
amounts of additive — 0.025 Na/S mol ratio, about 1/15 of that employed in the conven-
tional Coolside process.1"3 In another pilot plant test conducted with 0.03 mol CaCl2/mol
fresh Ca, system S02 removal without baghouse reheat was 97% at 1.2 Ca/S. This result
indicates the potential of the process for high efficiency S02 removal in a new plant with
a baghouse. Further testing is under way using different additives and additive dosages
and varying process conditions.

Process Economics

An interim process economic study was completed based on current process performance
data with a commercial hydrated lime and a conceptual process design. The objectives
for this study were to confirm the potential economic advantages of the Advanced Cool-
side process and to identify priorities for further process development. A final economic
study will be conducted at the conclusion of the pilot-scale development program.

The economic study compared costs of the Advanced Coolside process with limestone
wet scrubbing. Figure 8 shows a schematic of the Advanced Coolside process conceptual
design used in the study; many of the process systems, particularly ancillary equipment,
have not yet been optimized. The limestone wet FGD design used in the study is shown
in Figure 9; the design includes forced oxidation and use of a single absorber module

101-8

-------
with no spare. Economic assumptions (Table 4) were selected to assure comparison on
an equivalent basis. Both processes were evaluated for 90% S02 removal, an assumed
capital life of 30 years, 65% capacity factor, and using the same retrofit factors. The
analysis was based on an "nth" plant design, using an 18% contingency for each process.
Assumed delivered reagent costs were $60/ton for hydrated lime with 1% inerts and
$15/ton for limestone. All costs are reported in 1992 dollars.

The economic study confirmed a substantial capital cost advantage for the Advanced
Coolside process. Figure 10 shows that for a 2.5% sulfur coal the capital cost was about
40% less than forced oxidation limestone scrubbing, over the 150-500 MWe range of
plant sizes studied. The relative difference in capital cost was about the same for 1.5 and
3.5% sulfur coals. The lower capital cost can be important to utilities in making compli-
ance decisions because it reduces financial and regulatory risk.

The economic study quantified the potential S02 control cost advantages of the
Advanced Coolside process. Figure 11 shows that the process has a lower levelized cost
($/ton S02 removed) than limestone wet FGD over a wide range of coal sulfur contents
and plant sizes. The cost differential ranged from 21% for 1.5% sulfur coal and a
150 MWe plant, to 11% for 2.5% sulfur coal and a 250 MWe plant, to break even for
3.5% sulfur coal and a 500 MWe plant

The interim study also indicated that there is potential for further improvement of the
Advanced Coolside process and identified areas for improvement with the greatest
potential impact on economics. Figure 12 shows a breakdown of the levelized S02
control costs for a medium plant size and medium sulfur coal. Areas identified for cost
reduction include sorbent cost and equipment capital cost for certain process systems.
Approaches for equipment cost reduction include further contactor optimization and
improvement in other systems on which optimization studies have not yet focused (e.g.,
recycle handling, waste disposal and handling, and flue gas handling).

The goal of further development is to establish at least a 20% levelized cost advantage
over wet FGD over a wide range of compliance situations. This would make it more
attractive for utilities to employ a newer, less established technology.

Future Work

Based on the process economic study, the focus of future process development will be to
increase the cost advantage of Advanced Coolside over commercial technology through
equipment design optimization and sorbent improvement. For the economic study,
Turbotak developed preliminary full-scale designs for the simplified contactor based on
the test results with the original contactor. The results of pilot tests using the new
contactor will be used by Turbotak to develop a commercial design to further reduce
costs. In addition, other commercially available gas/liquid contacting devices will be
evaluated for use in the process. Equipment design optimization efforts will be expanded
to look at other systems with potential impact on process capital cost, as identified in the

101-9

-------
economic study. The sorbent improvement work under way will continue as described
above. The goals are to reduce sorbent usage or the required recycle ratio and to allow
use of lower cost sorbent sources. Another area for future investigation is air toxics
control, particularly that of mercury. A literature analysis under way suggests that the
Advanced Coolside process has potential for substantial reduction in Hg emissions. The
capability for air toxics control would provide an additional incentive to use this
technology for S02 compliance.

Acknowledgment

This work was conducted under partial sponsorship of the U.S. Department of Energy
Contract DE-AC22-91PC90360. The authors are grateful to C J. Drammond, T. D.
Brown, R. E. Tischer, and others at DOE Pittsburgh Energy Technology Center for their
support and encouragement of this project.

References

1.	Yoon, H., M. R, Stouffer, W. A. Rosenhoover, and R. M. Statnick. "Laboratory and
Field Development of Coolside S02 Abatement Technology." Proceedings, Second
Pittsburgh Coal Conference, Pittsburgh, PA (September 1985).

2.	Yoon, H., M. R. Stouffer, W. A. Rosenhoover, 3. A. Withum, and F. P. Burke.
"Pilot Process Variable Study of Coolside Desulfurization." Environ. Progress. Vol. 7,
No. 2, p. 104-11 (1988).

3.	Stouffer, M. R., H. Yoon, and F. P. Burke. "An Investigation of the Mechanisms of
Flue Gas Desulfurization by In-Duct Dry Sorbent Injection." I&EC Research. Vol.
28, No. 1, p. 20 (1989).

4.	Stouffer, M. R., W. A. Rosenhoover, and H. Yoon. "Pilot Plant Process and Sor-
bent Evaluation Studies for 105 MWe Coolside Desulfurization Process Demonstra-
tion." Proceedings, Third International Conference on Processing and Utilization of
High-Sulfur Coals, Ames, IA (November 1989).

5.	Withum, J. A., H. Yoon, F. P. Burke, and R. M. Statnick. "Coolside Desulfurization
Demonstration at Ohio Edison Edgewater Power Station." ACS Div. Fuel Chem.
Preprints. Vol. 35, No. 4, p. 1463-72 (1990).

6.	Yoon, H., R. M. Statnick, J. A. Withum, and D. C. McCoy. "Coolside Desulfuriza-
tion Process Demonstration at Ohio Edison Edgewater Power Station." Proc.

Annual Meet. - Air Waste Manage. Assoc. Vol. 9A, Paper No. 91/102.8 (1991).

101-10

-------
7.	Babu, M., R. C Forsythe, C. F. Runyon, D. A. Kanaiy, H. W. Pennline, T. Sarkus,
and J. L. Thompson. "Results of 1.0 MMBtu/Hour Testing and Plans for a 5 MW
Pilot HALT Program for S02 Control." Proceedings of the Third Annual Pittsburgh
Coal Conference, Pittsburgh, PA (September 1986).

8.	Babu, M., I. College, R. C. Forsythe, R. Herbert, D. A. Kanary, D. Kerivan, and
K. Lee. 5 MW Toronto HALT Pilot Plant." Proceedings, 4th DOE-PETC Contrac-
tors' Meeting, Pittsburgh, PA (August 1988).

9.	OT)owd, W. "Duct Injection Experiments at DOE-PETC," Technical Update

No. 18, Duct Injection Technology Development Program. (January 1991).

10.	Brown, C. A., M. Maibodi, and L. M. McGuire. "1.7 MW Pilot Results for the Duct
Injection FGD Process Using Hydrated lime Upstream of an ESP," Presented at
the 1991 S02 Control Symposium, Washington, DC (December 1991).

11.	Felix, L. G., J. P. Gooch, R. L. Merritt, M. G. Klett, J. E, Hunt, and A G. Demlan.
"Scale-Up Tests and Supporting Research for the Development of Duct Injection
Technology," Presented at the 1991 S02 Control Symposium, Washington, DC
(December 1991).

12.	Nolan, P. S., D. C. McCoy, R. M. Statnick, M. R. Stouffer, and H. Yoon.

"Economic Comparison of Coolside Sorbent Injection and Wet limestone FGD
Processes," Presented at the 1991 S02 Control Symposium, Washington, DC
(December 1991).

13.	Stouffer, M. R., W. A. Rosenhoover, and J. A Withum. "Advanced Coolside
Desulfurization Process," Presented at the 1992 Summer National Meeting of the
American Institute of Chemical Engineers, Minneapolis, MN (August 1992).

14.	Spink, D. R. "Gas Solid Separation and Mass Transfer Using a Unique Scrubbing
Concept," presented at the AIChE National Meeting, Houston, TX (March 29-
April 2,1987).

15.	Spink, D. R. "Handling Mists and Dusts," CHEMTECH, p. 364-8 (June 1988).

101-11

-------
Table 1. Advanced Coolside Pilot Plant Recycle Tests with Flue Gas Reheat Prior
to the Baghouse and 1500 ppm S02.

Test

1

2

3

4

Test Duration, hr

36

115

13

73

Process Conditions









Fresh Ca/S, mol

1.4

1.2

1.5

1.2

Recycle Ratio, lb/lb fresh lime

4.5

6.9

4.4

6.7

Recycle Pretreatment

Yes

Yes

Yes

Yes

Baghouse Approach Temp., °F

23

23

24

22

Hydrated Lime

A

A

A

B

Process Performance









S02 Removal, %, In-Duct

83

87

84

80

System

90

90

90

86

Baghouse

7

3

6

6

Sorbent Utilization, %, By Gas Analysis

63

75

60

70

By Solid Analysis

63

70

59

71

Table Z Advanced Coolside Pilot Plant Recycle Tests with No Flue Gas Reheat
Prior to the Baghouse and 1500 ppm S02.

Test

5

6

7

8

9

Test Duration, hr

40

28

21

25

23

Process Condfttons











Fresh Ca/S, mol

1.2

1.5

1.2

1.6

1.6

Recycle Ratio, lb/lb fresh lime

3.3

3.5

4.9

3.9

3.8

Recycle Pretreatment

Yes

Yes

Yes

Yes

Yes

Baghouse Approach Temp., °F

9

12

9

11

12

Hydrated Lime

A

A

A

A

A

Process Performance











S02 Removal, %, In-Duct

60

70

81

91

88

System

84

90

88

97

99

Baghouse

24

20

7

6

11

Sorbent Utilization, %, By Gas Analysis

67

61

71

60

61

By Solid Analysis

68

63

68

58

61

Table 3. Optimization of Contactor Operating Conditions.

Base Design

Conditions
Nozzle Water Row
Nozzle Air Pressure
Nozzle Air Flow
Performance
Exit Humidity
Fly Ash Collection

1.13 gpm/1000 acfm

45-50 psig
17 scfm/1000 acfm

>98%

>95%

Alternative

0.6 gpm/1000 acfm

30 psig
9 scfm/1000 acfm

>98%

-95%

101-12

-------
Tabie 4. Key Assumptions of interim Process Economic Study.

Advanced Coolside

Forced Oxidation Wet FGD

Delivered Sorbent Cost
Waste Disposal Cost
S02 Removal
Capacity Factor
Capital Life
Retrofit Factor
Location Factor
Design Philosophy
Sparing
Indirect Costs
Construction

$60/ton, 7% inerts (hydrated lime)
$6.50/wet ton
90%

65%

30 years
Medium (1.22-1.34)

1.06

¦nth" plant, 18% capital contingency
Auxiliary equip, only, no major equip.
37.2% of direct
2 years

$1S/ton (limestone)
$6.50/wet ton

90%

65%

30 years
Medium
1.06

"nth" plant 18% capital contingency
Auxiliary equip, only, no major equip.
37.2% of direct
3 years

HYDRATED
LIME

Figure 1. Conceptual Diagram of Advanced Coolside Desulfurization Process.

101-13

-------
Figure 2. Schematic of Advanced Coolside Pilot Plant.

SORBENT UTILIZATION, %

65	75

BY GAS ANALYSIS

Figure 3. Parity Plot of % Sorbent Utilizations by Gas Analysis
and by Solids Analysis for Recycle Tests.

101-14

A

-------
Figure 4. Initial Pilot Test Data For Contactor Simplification. Droplet Surface
Area is Relative to that at Design Conditions.

GAS
FLOW

ii

—a

'M*

/ < . v

MIST
ELIMINATOR

Figure 5. Schematic of Simplified Contactor Design.

101-15

-------
. -i?i¦£$£'

"	fc

¦

•v .... . .

¦ i!##l
¦ ¦ - .

J*"

»»

RELATIVE SPECIFIC DROPLET SURFACE AREA

Figure 6. Optimization Test Results Showing Relative Humidity vs.
Relative Droplet Surface Area.

96	99	190

PERCENT RELATIVE HUMIDITY

Figure 7. Optimization Test Results Showing Fly Ash Capture Results for
Tests with >95% Relative Humidity.

(Fly ash capture was measured upstream of the mist eliminator.)

101-16

-------
Figure 8. Schematic of Advanced Cooiside Conceptual Design
Used in Interim Economic Study.

Figure 9. Schematic of Limestone Forced Oxidation Wet FGD Process
Design Used in Economic Study.

101-17

-------
WET
FGD

ADVANCED
COOLSIDE

300

5

m

250

CO
O

o

g

CL
<
o

200

150

100

50











i

* • • m











* m „

• a «

* • • •

»









2.5 % S
COAL











100

200

300

400

500

600

PLANT GROSS MW

Figure 10. Comparison of Capital Costs for Advanced Coolside and Wet Limestone
Forced Oxidation FGD at 2.5% Coal Sulfur Content and Varying Plant Sizes.

WET
FGD

ADVANCED
COOLSiDE

250

100 150 200 250 300 350 400 450 500 550 600
PLANT SIZE, GROSS MW

Figure 11. Comparison of Levelized S02 Control Costs for Advanced Coolside and
Wet Limestone Forced Oxidation FGD as a Function of
Coal Sulfur Content and Plant Size.

101-18

-------
Capital Charges = 53.7%	O&M Charges = 46,3%

Figure 12. Breakdown of Levelled S02 Control Costs for
262 MW/2.5% S Coal Case.

(Capital costs are effective costs, including capital maintenance.)

101-19

-------
ECONOMIC ANALYSIS OF FGD BY-PRODUCT DISPOSAL ALTERNATIVES

D, Lynn Forster
Department of Agricultural Economics and Rural Sociology
Hie Ohio State University

2120 Fyffe Road
Columbus, Ohio 43210

Jonathan Rausch
Cornell Extension Service
0949 Larsen Road
Geneva, NY 14456

Abstract

A mathematical programming model of electric utility by-product disposal alternatives
is developed. Least cost disposal includes widespread use of the by-product on
cropland; however, the vast majority of the material is landfilled. Agricultural use
represents substantial cost savings to utilities.

Introduction

Acid rain has long been suspected as a cause of the deterioration of streams, lakes,
forests, soils, and various fabricated structures. These resources are adversely effected
by acidic precipitation linked to increased sulfur dioxide and nitrogen oxides emissions.
Acid rain is formed when sulfur dioxide and nitrogen oxide react with other chemicals
in the atmosphere (Helme and Neme, 1991). The primary source of sulfur dioxide and
nitrogen oxide, as identified by the United States Environmental Protection Agency, is
associated with the combustion of coal used in the production of electricity (Helme
and Neme, 1991). The amount of sulfur dioxide produced depends upon the sulfur
content of the coal being combusted. Coal higher in sulfur inevitably produces higher
concentrations of sulfur dioxide and precipitation that is more acidic.

102-1

-------
The actual or potential degradation of resources by add rain are vast. For example, in
the Adirondack Mountains up to 15 percent of the medium to large lakes, those lakes
greater than ten acres, are chronically acidic due primarily to add rain, and 25 percent
of small late are likewise effected (Helme and Neme, 1991). Hie National Add
Predpitation Assessment Program (NAPAP) has estimated that nearly 20 percent of
the nation's lakes and streams have little or no add-buffering capacity, thus are
susceptible to current and future addification,

Sulfiir dioxide (SQ,) emissions are concentrated primarify along the OMo River Valley.
Forty-four percent of the U.S. S02 emissions are produced in this region by Ohio,
Indiana, Pennsylvania, Illinois, West Virginia, with the indusion of Missouri and
Tennessee, in addition, four of the five highest S02 producers are also among the top
ten nitrogen oxide (NOJ producing states (EPA, 1986). Clearly, the OMo River
Valley is a major producer of emissions associated with acid rain, and sipiificantly
impacted by legislation mandating emission standards.

Emission Abatement Policy

Title IV of the 1990 Clean Air Act addresses sulfiir dioxide, nitrogen oxide, and
particulate matter emissions assodated with the burning of fossil fuels. This legislation
mandates a 10 million ton (40 percent) reduction in the nation's sulfiir dioxide
emissions (based upon 1980 emission levels) by the year 2000, and a two million ton
reduction in nitrogen oxide (Claussen, 1991).

The add rain program developed by the EPA under this title sets a ceiling on sulfur
dioxide emissions from electric power plants and allows individual utility companies to
determine the most cost effective means of achieving these new mandates.

Compliance is expected to be achieved through conservation efforts, using fuels lower
in sulfur, purchasing of emission allowances, retrofitting existing plants with pollution
control devices, and/or a combination of the above.

Currently the only pollution reduction technology which can be used on existing power
plants to reduce S02 emissions to mandated levels is flue gas desulfurization (FGD).
Through the use of a sorbent, such as limestone, exhaust gases are "scrubbed" of S02.
One such process is referred to as a dry scrubber process. These FGD technologies
are capable of reducing S02 emissions by as much as 95 percent from current power
plant emissions (EPA, 1986). However, this process of "scrubbing" creates another
environmental concern-disposal of the used sorbent.

Not all power plants are expected to convert to dry injection technology. Of the 23
power plants identified in Ohio, potential converters to dry injection FGD technology
by the 2000 were estimated. In this analysis, a least cost transportation model was
developed depicting the least cost distribution of dry FGD by-product material from

102-2

-------
these potential electric power generating sites to numerous end use alternatives. The
scenario assumes six power plants retrofit with this dry injection technology. This
analysis is only preliminary since options regarding S02 emissions reduction at the
plants are still being weighted.

Proposed Uses of Dry FGD By-Product

LandfTlling of this by-product material is the current means of disposal. However, it
has been suggested that dry FGD by-product material has chemical properties which
makes it valuable as an agricultural lime substitute, soil amendment in coal surface
mine reclamation, and highway embankment construction material This analysis
focuses on dry FGD by-product use in agriculture land application, current coal
surface mine reclamation, and landfilling.

Agricultural Land Application

FGD by-product has many characteristics similar to agricultural lime, and it is
expected that dry FGD by-product materials could be a close substitute for agricultural
lime. Annual usage of agricultural lime was estimated for each Ohio county. These
estimates were based on annual sales of agricultural lime in 1987-92. Since FGD by-
product has 60 percent the neutralizing potential of agricultural lime, agricultural lime
sales were adjusted to estimate FGD by-product demand. That is, FGD by-product
has a lower total neutralizing potential (TOP), thus higher application rates are
necessary to achieve the same results as agricultural lime. The estimated quantity of
FGD by-product demanded reflects these adjustments in TNP.

Transportation of the FGD by-product is expected to be similar to that of agricultural
lime. Only slight modification of existing equipment is necessary to transport the by-
product from the source (power plant) to the destination (farm, coal mine or landfill),
and to apply the by-product to agricultural land. Under these assumptions, FGD by-
product is expected to be hauled from the power plant to various locations throughout
the state by tracks. Once the by-product has reached the farm it is expected that
conventional lime spreaders will apply this by-product to agricultural lands.

Strip/Surface Coal Mine Reclamation

Coal surface mine operations are required to reclaim lands which have been mined.
During the reclamation phase lime is often used to return the mined spoils back to a
pH level conductive to plant growth. It is expected that FGD by-product materials can
also be used for this same purpose.

102-3

-------
The estimated quantity of FGD by-product material demanded to meet current surface
mine reclamation work was based on the estimated area of surface mining in each
county. These estimates are derived from data reporting tons of coal sold in each
OMo coal mining county (1989). From these estimates an average number of tons per
acre of coal extraction were used to estimate the number of acres displaced by surface
coal mining in a given year. Based upon the estimated number of surface acres mined
and an application rate of 20 tons per acre, an estimate for the quantity of FGD by-
produce demanded can be determined. Note, the application rate is that which is
expected to be used in experimental field work.

The final end use alternative identified is that of landfffling. It is expected that FGD
by-product will be landfilled in the event that the available quantity of the product is
larger than its economical use in agricultural and/or coal surface mine reclamation. It
is farther assumed that landfiUmg is not constrained in the quantity of the by-product
that ran be accepted, and that landfills are in close proximity to the power plants or
sources of FGD by-products.

Objectives of this Research

The objectives of this research are (1) to develop a model to identify the least cost
disposal methods of FGD by-product among the three stated alternative end uses from
the producer or power plants perspective, (2) estimate the quantity of by-product used
in each alternative, and (3) estimate the shadow price associated with each alternative
end use. This particular component of the research does not address the social
amenities/disamenities associated with FGD by-product disposal. Examples of these
omitted off-site impacts include: deterioration of roads and bridges from increased
traffic; property value pin/loss from landfill activities, surface mine reclamation, and
abandon mine land reclamation; and increased/decreased ground and/or surface water
quality from surface mine reclamation and abandon mine land reclamation, or
landfilling activities. However, work is in progress to quantify these
amenities/disamenities. Ones quantified the least cost disposal model can be re-
estimated with the appropriate increase/decrease in cost associated with the given end
use alternative. That is, the model can be modified to consider the social gain/loss
associated with each end use alternative, providing a least cost disposal model which
captures both a private and social accounting stance.

Model Development

Minimizing the total cost of transporting a product from some production point to
various destination or demand points can be done through a series of linear equations.
This type of least cost mathematical modeling, or transportation modeling, has been
applied to the distribution of FGD by-product material to various destinations

102-4

-------
throughout Ohio. This least cost transportation model is based upon six source nodes
(power plants) associated with the production of dry FGD by-product materials and its
use as a substitute for agricultural lime in 88 Ohio counties, as a soil amendment in 21
Ohio coal surface mining counties, and disposal in six landfill sites located in proximity
of the power plant

To formulate this model mathematically, the following terms are defined.

Zi = the number of tons of FGD by-product material available at the power
plant or source i, i=l,2,...,m;

bj = maximum number of tons of by-product required at each destination or
alternative use (e.g., county for agricultural lime, reclamation site, or landfill), j
= 1,2,...n;

Cjj = unit transportation and application cost from each source i to each
destination j, (i « l,2,...,m; j ~ l,2,...,n).

E aj = E b^ the quantity of FGD by-product available at source i must be equal
to the demand for FGD by-product material in alternative end uses (including
landfilling options).

aj > 0; supply of by-product at each source node is positive.

bj > 0; demand for by-product at each demand node is positive.

The problem then becomes determining the amount of FGD by-product material
shipped to each of these alternative end uses, given that the cost of distribution and
application of the by-product is known or can be estimated. Thus, the decision
variable, x^, equals the number of tons of by-product material shipped from each
source i to each destination j annually given some cost per unit shipped.

Hie transportation model estimated is:

Assuming:

N N

Minimum Cost = ^ J) c.x~

(1)

i /

subject to:

E Xjj = ai (i = 1,2,...,m);

£ Xg < bj (j = l,2,...,n);

Xjj > 0 (i = l,2,...,m; j = l,2,...,n).

(2)

(3)

(4)

102-5

-------
Equation (1) represents the minimization of total distribution costs, assuming a linear
cost structure for shipping, processing, and application of the dry FGD by-product
material. Equation (2) shows that the quantity of by-product shipped from each source
to all alternative end use destinations most be equal to the quantity of by-product
materials available. Equation (3) states that the quantity of by-product shipped from
all sources to each destination must be less than or equal to the maximum quantity of
by-product demanded at that destination. Finally, equation (4) indicates that the
quantity of by-product shipped from each source to each destination must be greater
than or equal to zero.

Data Used for Estimating the Transportation Model Using Six Power Plants

All estimates pertaining to the quantity of dry FGD by-product demanded at various
demand nodes for the two end use alternatives (agricultural and surface coal mine
reclamation uses) have been adjusted for a 25 percent rate of adoption. It is expected
that not all individuals using agricultural lime or reclaiming surface coal mines will
completely adopt this new technology. Thus, the model uses a conservative or lower
rate of adoption. However, it is important to note that the model can be re-run at
various levels of adoption.

linear distances from each power plant or source of diy FGD by-product to the center
of each county were estimated. There were a total of 528 possible agriculture routes
(88 counties supplied from six plants), and 126 possible coal reclamation routes (21
counties supplied from six plants).

Transportation costs associated with moving the specified distance were estimated. In
the case of agricultural land application, cost estimates were derived from the
agricultural lime industry. It is expected that the dry FGD by-product will be
transported in much the same manner as current agricultural lime. Thus, an estimate
of $0.10 per ton per mile was used. In addition to moving the product from the source
to the destination an application expense is incurred. Again, the application of the diy
FGD by-product is expected to be similar to agricultural lime, which 1ms an estimated
application charge of $3.50 per ton. This expense includes the spreader and
mechanism to load the product from a pile onto the spreader. Therefore,
transportation costs were calculated at $0.10 per ton per milk, then an additional $3 JO
per ton was added to each for ejected application costs.

Transportation of the diy FGD by-product to coal surface mine reclamation sites is
also expected to cost $0.10 per ton per mile. Application of the dry FGD by-product
is expected to be at significantly higher levels, potentially 20 tons per acre, than
application rates associated with agricultural use. Thus, different equipment is
expected to be used in distributing the by-product (e.g., a bulldozer or equivalent type

102-6

-------
of reclamation machinery), but application costs would not vary significantly and were
estimated at $3.50 per ton.

Cost estimates for landfilling were obtained from interviews with representatives of
electric utilities. All landfilling activities are regulated, and industry representtives
expected stringent landfilling requirements. Estimated costs were about $27 JO per ton
of material to landfill dry FGD waste and meet expected EPA guidelines.

Results

Quantities of dry FGD by-product shipped from the six power plants to each county
for agricultural uses, surface mine reclamation, and landfilling are shown in Table 1.

Figure 1 shows the percentage distribution of dry FGD by-product produced at six
Ohio power plants among the three end use alternatives. Notice that at five of the
power plants (Mountainer [Mount], Muskingum [Musk], Sammis, and Burger) at least
85 percent of the dry FGD by-product is landfilled. At two of the six power plants
(Edgewater [Edge] and Niles) nearly all by-product material is applied to agricultural
land. Surface coal mine reclamation accounts for negligible quantities of total FGD
by-product produced (1.03%) from all six power plants.

Figure 2 shows the quantity of dry FGD by-product distributed among the three end
use alternatives from each source. Of the total quantity of dry FGD by-product
produced, assuming a 25% adoption rate, little dry FGD by-product is used in
agriculture (8.87%) or surface coal mine reclamation (1.03%), and most is landfilled
(90.1%). This suggests that (a) a higher adoption rate among farmers and surface
mine operators would reduce the quantity of landfilled, (b) the amount of dry FGD
by-product produced greatly exceeds current quantities demanded or (c) additional end
uses for the by-product must be found. Based upon the estimated total output of dry
FGD by-product from these six power plants to be 4.1 million tons annually, current
agricultural land application and coal surface mine reclamation use at 100% adoption
is estimated to be about 1.7 million tons per year. This suggests that 2.4 million tons
of dry FGD by-product must be either landfilled or some alternative use must be
found assuming complete substitution of dry FGD by-product for agricultural and
surface coal mine reclamation.

Figure 3 shows the distribution of dry FGD by-product as a substitute for agricultural
lime from these six power plants. Quartiles were developed based upon the quantity
of dry FGD by-product estimated to be shipped to each Ohio county for agricultural
land application. The first quartile, or counties using the lowest 25% of dry FGD by-
product, are not shaded. Counties with shading indicate the second lowest 25% of
agricultural use through the largest 25%.

102-7

-------
Figure 4 shows a quartiled distribution, of dry FGD by-product for coal surface mine
reclamation. Since only 21 counties have coal surface mine activity, all quartiles are
shaded, and represent the lowest to highest quantities of dry FGD by-product used.
Again, the quantity of dry FGD by-product associated with a given quartile are in
hundreds of tons, and represent one-fourth of the total dry FGD by-product shipped
for coal surface mine reclamation from the six power plants selected.

The final geographic representation of dry FGD by-product distribution shows both
agricultural land application and coal surface mine reclamation, or a quartiled total
distribution of dry FGD by-product (Figure 5). Each quartile represents one-fourth of
the dry FGD by-product used as an amendment to land (both agricultural and coal
surface mine declamation lands). For example, four counties (Belmont, Coshocton,
Noble, and Tuscarawas) within the fourth quartile receive one-fourth of the total dry
FGD by-product applied to land.

The final objective of this research was to estimate the shadow price associated with
each demand or destination node. Shadow prices are imputed prices which reflect the
increase/decrease in total costs if one additional unit of product were available. In
this case, the additional unit is the use of one additional ton of dry FGD by-product on
agricultural land or in coal surface mine reclamation. Shadow prices for diy FGD by-
product are calculated as the difference between landfilling (non-binding constraint)
and agricultural land application or coal surface mine reclamation options (both
binding constraints). It would be expected that as the distance from the power plant
increases, the cost to move this by-product also increases, therefore the imputed value
of shadow price for binding end use options father from the source or power plant
would be lower. That is, the difference between landfilling and shipping dry FGD by-
product material greater distances would be smaller. Thus, counties located father
from dry FGD by-product sources would have lower shadow prices or lower cost
savings to the utility companies than would land application sites closer in proximity to
the power plant Figure 6 shows the shadow price associated with the distribution of
dry FGD by-product to each Ohio county.

Another interpretation of these shadow prices is the amount the power plant would be
willing to pay for the disposal of an additional ton of by-product in each end use
alternative. For example, Figure 6 shows Williams County in the first quartile or
having a shadow price between $3.29-15.01 per unit. Hie calculated shadow price for
Williams County is $3.29 per ton, suggesting the power plant would be willing to pay
the fanner up to $3.29 to use an additional ton of dry FGD by-product as opposed to
landfilling it at a cost of $27.50 per ton.

102-8

-------
Conclusion

Under Title IV of the 1990 Clean Air Act, electric power generating plants will be
required to reduce S02 and NOx emissions by about 40 percent no later than the year
2000. These emission standards are most significant for power plants which burn coal,
particularly high sulfur coal, as an energy source. In order to achieve compliance with
these mandates, power plants will have to reduce the quantity of coal burned, use fuels
lower in sulfur, purchase emission allowances, retrofit existing power plants with dean
air technology, or a combination of the above. Currently the only clean air technology
available to existing power plants is Hue Gas Desulferization (FGD) technology. EPA
has estimated that this technology can reduce S02 by as much as 95 percent.

However, dry FGD technology creates another environmental concern—disposal of the
used sorbent. Based upon current coal consumption estimates, Ohio could potentially
produce nearly 4 million tons of diy FGD by-product (used sorbent) annually from six
Ohio power plants. The objective of this research was to estimate a least cost disposal
model for the movement of this by-product to various geographic locations throughout
Ohio and for use as a soil amendment for agricultural land, coal surface mine
reclamation, and landfilling. In doing so, total disposal costs and quantities were
derived as well as shadow prices for each county (demand node) identified.

This analysis used one scenario for future dry FGD by-product distribution and uses.
Others are plausible and easily analyzed by this model. The linear transportation
model estimated suggests that the least cost disposal of diy FGD by-product would
cost $107 million to dispose of 4.1 million tons of diy FGD by-product material
annually. This is equivalent to average cost of $26.10 per ton of dry FGD produced.
In addition to landfilling, the two end use alternatives currently identified are dry FGD
by-product use as soil amendments on agricultural land and coal surface mine
reclamation. Of these three options, landfilling accounts for about 90 percent of the
total by-product produced, while agricultural and surface coal mine reclamation
account for about 8.87% and 1.03% respectively. Therefore, on average land
application of dry FGD by-product represents a savings of $1.40 ($2750 [landfilling]-
$26 .10 [average cost of disposal per ton]). However, at the margin (one additional
ton), land application represents a cost savings of about $14.81 per ton. That is,
electric power plants can reduce total cost of dry FGD by-product disposal by $14.81
per ton through land application rather than landfilling.

In this analysis, land spreading dry FGD by-product is dispersed widely over the state.
Agricultural use is concentrated in the western two-thirds of Ohio, while coal surface
mine reclamation use is important in the eastern one-third of the state. Yet, of the
alternative disposal options identified, the vast majority (3.8 million tons) of dry FGD
by-product is landfilled, followed by agricultural land application (374,276 tons) and
then coal surface mine reclamation (42,507 tons). Given these end use alternatives,
more FGD by-product will be buried in landfills than will be used in other alternatives
combined. However, electric utilities have an enormous economic incentive to supply

102-9

-------
dry FGD by-product to land. They could pay farmers substantial amounts and save
lanrlfflTing costs.

Acknowledgements

This research was conducted as part of the "Land Application Uses for Dry FGD By-
products" project which is a cooperative project of the Ohio Agricultural Research and
Development Center, Hie Ohio State University, The U.S. Geological Survey and the
Dravo lime Company. Funding support for this project was obtained from the Ohio
Coal Development Office (Columbus, OH) Grant No. CDO/D-89-35, The U.S.
Department of Energy (Morgantown Energy Technology Center, Morgantown, WV)
Award No. DE-FC21-91MC2806G, Dravo Lime Company (Pittsburgh, PA) Grant No.
RF768342, Electric Power Research Institute (Palo Alto, CA) Grant No. RP2796-02,
Ohio Edison Company (Akron, OH), American Electric Power (Columbus, OH), and
The Ohio State University (Columbus and Wooster, OH).

References

EPA Journal. "Acid Rain: An EPA Journal Special Supplement" Office of Public
Affairs, EPA Journal, Washington, DC, June/July 1986, Vol 12, No. 5.

Ned Helme and Chris Neme. "Acid Rain: The Problem." Office of Communication
and Public Affairs, EPA Journal, Washington, DC, Jan./Feb. 1991, Vol. 17, No. 1.

Eileen Claussexi. "Acid Rain: The Strategy." Office of Communication and Public
Affairs, EPA Journal, Washington, DC, Jan./Feb. 1991, Vol. 17, No. 1.

102-10

-------
Table 1

Estimated Tons of Dry FGD By-Product Shipped to Each End Use Alternative From
Each. Source or Power Plant



Mountain-

Edge-

Musk-







Total



eer

water

ingum

Niles

Sammis

Burger

Ship-

County

S25

Sll

S16

S9

S8

S12

ped

Agricultural Land Application













Adams

6,718

0

0

0

0

0

6,718

Allen

5,091

0

0

0

0

0

5,091

Ashland

0

0

0

1,821

978

0

2,799

Ashtabula

0

0

0

5,599

0

0

5,599

Athens

0

0

1,258

0

0

0

1,258

Auglaize

8,726

0

0

0

0

0

8,726

Belmont

0

0

0

0

0

4,921

4,921

Brown

6,523

0

0

0

0

0

6,523

Butler

2,021

0

0

0

0

0

2,021

Carrol

0

0

0

0

2,041

0

2,041

Champaign

4,048

0

0

0

0

0

4,048

Clark

3,194

0

0

0

0

0

3,194

Clermont

4,998

0

0

0

0

0

4,998

Clinton

7,002

0

0

0

0

0

7,002

Columbiana

0

0

0

0

6,853

0

6,853

Coshocton

0

0

0

0

0

11,259

11,259

Crawford

0

0

0

0

0

6,857

6,857

Cuyahoga

0

3,989

0

0

0

0

3,989

Darke

6,290

0

0

0

0

0

6,290

Defiance

0

11,016

0

0

0

0

11,016

Delaware

0

0

9,199

0

0

0

9,199

Erie

0

6,586

0

0

0

0

6,586

Fairfield

0

0

9,724

0

0

0

9,724

Fayette

5,612

0

0

0

0

0

5,612

Franklin

2378

0

0

0

0

0

2,378

Fulton

0

9,404

0

0

0

0

9,404

Gallia

3,917

0

0

0

0

0

3,917

Geauga

0

0

0

313

0

0

313

Greene

4,418

0

0

0

0

0

4,418

Guernsey

0

0

0

0

0

2,827

2,827

Hamilton

638

0

0

0

0

0

638

Hancock

0

0

3,577

0

0

0

3,577

Hardin

0

0

7,457

0

0

0

7,457

102-11

-------
Table 1 (eont.)



Mountain-

Edge-

Musk-







Total



eer

water

ingum

Mies

Sammis

Burger

Ship-

County

S25

Sll

S16

S9

SS

S12

ped

Harrison

0

0

0

0

456

0

456

Henry

0

5,745

0

0

0

0

5,745

Highland

14,993

0

0

0

0

0

14,993

Hocking

499

0

0

0

0

0

499

Holmes

0

0

0

0

0

3,237

3,237

Huron

0

0

0

16,983

0

0

16,983

Jacikson

109

0

0

0

0

0

109

Jefferson

0

0

0

0

1,373

0

1,373

Knox

0

0

8,513

0

0

0

8^13

Lake

0

0

0

1,539

0

0

1^539

Lawrence

148

0

0

0

0

0

148

licking

0

0

5,395

0

0

0

5395

Logan

1,009

0

0

0

0

0

1,009

Lorain

0

431

0

3,812

0

0

4,243

Lucas

0

4,770

0

0

0

0

4,770

Madison

712

0

0

0

0

0

712

Mahoning

0

0

0

2,438

0

0

2,438

Marion

0

0

1,903

0

0

0

1,903

Medina

0

0

0

4,879

0

0

4,879

Meip

116

0

0

0

0

0

116

Mercer

2,534

0

0

0

0

0

2£34

Miami

2,268

0

0

0

0

0

2,268

Monroe

0

0

0

0

0

370

370

Montgomery

611

0

0

0

0

0

611

Morgan

0

0

628

0

0

0

628

Morrow

0

0

1,068

0

0

0

1,068

Muskingum

0

0

5,665

0

0

0

5,665

Noble

0

0

398

0

0

0

398

Ottawa

0

6,070

0

0

0

0

6,070

Paulding

6,170

0

0

0

0

0

6,170

Perry

0

0

1,241

0

0

0

1^241

Pickaway

2,946

0

0

0

0

0

2,946

Pike

5,062

0

0

0

0

0

5,062

Portage

0

0

0

3,483

0

0

3,483

Preble

2,660

0

0

0

0

0

2,660

Putnam

0

0

2,501

0

0

0

2^01

Richland

0

0

0

0

0

4,119

4,119

Ross

1,630

0

0

0

0

0

1,630

102-12

-------
Table 1 (eont.)



Mountain-

Edge-

Musk-







Total



eer

water

ingum

Niles

Sammis

Burger

Ship-

County

S25

Sll

S16

S9

S8

S12

ped

Sandusky

0

3,019

0

0

0

0

3,019

Scioto

454

0

0

0

0

0

454

Seneca

0

5,542

0

0

0

0

5,542

Shelby

1,962

0

0

0

0

0

1,962

Stark

0

0

0

0

2,381

0

2^81

Summit

0

0

0

1,543

0

0

1,543

Trumble

0

0

0

3,846

0

0

3,846

Tuscarawas

0

0

0

0

2,695

0

2,695

Union

1,948

0

0

0

0

0

1,948

Van Wert

1,115

0

0

0

0

0

1,115

Vinton

153

0

0

0

0

0

153

Warren

1,720

0

0

0

0

0

1,720

Washington

0

0

1,963

0

0

0

1,963

Wayne

0

0

0

0

11,462

0

11,462

Williams

0

6,405

0

0

0

0

6,405

Wood

0

7,023

0

0

0

0

7,023

Wyandot

0

0

0

0

0

5,724

5,724

102-13

-------
Table 1 (cant.)



Mountain

Edge-

Musk-







Total



-eer

water

ingum

Niles

Sammk

Burger

Ship-

County

S25

Sll

S16 "

S9

S8

S12

ped

Coal Surface Mine Reclamation and Landfilld Alternative





Athens

0

0

312

0

0

0

312

Belmont

0

0

0

0

0

5,566

5,566

Carroll

0

0

0

0

1,637

0

1,637

Columbiana

0

0

0

0

1^68

0

1,268

Coshocton

0

0

0

0

0

3,904

3,904

Guernsey

0

0

0

0

0

287

287

Harrison

0

0

0

0

4,262

0

4,262

Hancock

262

0

0

0

0

0

262

Holmes

0

0

0

0

0

933

934

Jackson

2,049

0

0

0

0

0

2,049

Jefferson

0

0

0

0

3,954

0

3,954

Lawrence

55

0

0

0

0

0

55

Mahoning

0

0

0

409

0

0

409

Muskingum

0

0

1,858

0

0

0

1,858

Noble

0

0

6,917

0

0

0

6,917

Peny

0

0

841

0



0

841

Stark

0

0

0

0

404

0

404

Tuscarawas

0

0

0

0

4,343

0

4,343

Vinton

3,047

0

0

0

0

0

3,047

Washington

0

00

135

0

0

0

135

Wayne

0

0

0

0

66

o

0

66

landfill

757,524

0

869,447

0

1,689,128

386,656

3,702,755

Summary of Shipment Estimates











Total Ag.

120,393

70,000

60,490

46,256

37,823

39,314

374,276

Total Reclaim

5,413

0

10,063

409

15,933

10,690

42^07

Total Landfill

757,524

0

869,447

0

1,679,544

386,656

3,693,172

Total

883,330

70,000

940,000

46,665

1,733300

436,660

4,109,955

% Ag.

13.63

100.00

6.44

99.12

2.18

9.00

8.87

% Reclaim

0.61

0

1.07

0.88

0.92

2.45

1.03

% T-a-nrtfjH

85.76

0

92.49

0

96.90

88.55

89.86

% Total

100.00

100.00

100.00

100.00

100.00

100.00

100.00

102-14

-------
Percentage Distribution of FGD
By-Product By Source

1.2-T

Mount

S25

Edge

S11

Musk Nifes Sammis Burger
Si 6 SB	SB	S12

m % Ag. Use p//j % Reclainn Use ESS % Lancfi

Figure I. Percentage of Total FGD By-Product Produced and Distributed Among
the Three End Use Alternatives Selected by Plant

102-15

-------
Quantity Distribution Of FGD By-Product
By Source

-i.su	

Mount: Edge Musk NUes Sammis Burger
S25	S11	S16	S9	SB	S12

B51 to Use W7k Reclaim Use g^LsncfB

Figure 2, Quantity of FGD By-Product Produced and Distributed Among the Three
Stated End Use Alternatives % Selected Ohio Power Plants

102-16

-------
Figure 3, Quartiled Distribution of Agricultural Land Application of FGD By-
Product Assuming a 25% Rate of Adoption

102-17

-------
OinKlied distribution ©f
FCD hy product %m*4 m
curfKi eeal xnlna rKliiBition

Fini Qwilla
(S47-2JMB tons)

StcofMi Qt2»rill«

C3^W7-3^04 ton*)

Third Quuiflc
(3JiS3*4,34Z ions)

Fourth fiunifla
C5.S85-M1? tons)

Figure 4. Coal Surface Mine Reclamation Use of Dry FGD By-Product Assuming a
25% Adoption Level

102-18

-------
Figure 5. Agricultural Land Application and Coal Surface Mine Reclamation Use
of Dry FGD By-Product Assuming a 25% Level of Adoption

102-19

-------
Figure 6. Shadow Prices for Land Application or Electric Utilities Estimated Cost
Savings From Land application Versus Landfflling Dry FGD By-Product

102-20

A

-------
CONVERSION FROM DISPOSAL TO
COMMERCIAL-GRADE GYPSUM

An Alternative Approach for Disposal of Scrubber Wastes

Stanley K. Conn

Owensboro Municipal Utilities

4301 U.S. Highway 60 E
Owensboro, Kentucky 42303

Michael G. Vacek
Sargent & Lundy
55 East Monroe
Chicago, IL 60603

John T. Morris, Jr.
Wheelabrator Air Pollution Control
441 Smithfield Street
Pittsburgh, Pennsylvania 15222

Introduction

Owensboro Municipal Utilities' (OMU) Elmer Smith Units 1 and 2 were identified as
Phase 1 units in the 1990 Clean Air Act Amendments. To comply with the
amendment requirements, OMU opted to install a wet flue gas desulfurization (FGD)
system producing disposal-grade gypsum. The FGD system is common to both units.
Project engineering, procurement, and construction activities were proceeding on
schedule when difficulties were encountered with the development of a suitable
disposal site. These difficulties caused OMU to reevaluate its overall station waste
disposal plan. OMU's evaluation considered continuing with the development of an
OMU landfill, contracting for a third party to dispose of the waste in an engineered
landfill and modifying the process to produce a commercial-grade gypsum.
Ultimately, the FGD process equipment was modified such that gypsum could be
used for wallboard production.

103-1

-------
This paper summarizes events leading to this decision and OMU's economic
evaluation of the disposal alternatives- Also discussed are the technical
considerations for modifying the process and equipment design to produce a
commercial-grade gypsum. The revised process equipment design and arrangement
is also presented in the paper.

Background

The OMU Elmer Smith Station consists of two coal-fired generating unite with a total
capacity of 441 MW, Unit 1 is a cyclone-fired unit with a gross capacity of 151 MYV.
Unit 2 is a tangentially-fired pulverized coal unit with a gross capacity of 290 MW.
Both units fire coal from various sources in the Illinois basin. The sulfur content of
the coal ranges to a maximum of 6.0 lb/MBtu.

The work at the Elmer Smith Station to comply with the Clean Air Amendments is
identified as the Clean Air Project The Clean Air Project includes replacing the
existing electrostatic dust precipitators on each unit and installing a wet FGD system
that is common to both, unite. In addition, a low-NOx burner system will be installed
on Unit 2 to comply with the NOx provisions of the amendments.

Design activities were initiated on the project in March 1991 when Sargent & Lundy
(S&L) was selected as the architect-engineer (AE) for the project In November 1991,
Wheelabrator Air Pollution Control Corporation (WAPC) was awarded three
contracts to provide the FGD system, electrostatic precipitators and an acid brick-
lined chimney. WAPC's scope of work on these contracts is to design, furnish, and
install all "process-related" equipment for the project. Sargent & Lundy, as the AE, is
to provide balance-of-plant equipment facilities design such as foundations, structural
steel, building enclosures, all electrical facilities, control system, limestone handling
system, ash handling system modifications, booster fans, and station services (water,
air, etc.). A summary of the design features of major equipment associated with the
project is shown on Table 1.

In parallel with the S&L and WAPC efforts to design the on-site facilities, OMU also
contracted with Fuller, Mossberger, Scott and May (FMSM) of Lexington, Kentucky,
to locate a suitable disposal site for the waste gypsum and develop the site. After
several site selection screenings, OMU secured options on a 400-acre site that is
located 15 miles from the Elmer Smith Station. The design of the landfill was based
on a state-of-the-art lined landfill facility with a leachate collection system. Design
and permit preparation activities for the facility had been initiated when, in April
1992, pressure from a citizens group, local to the landfill site, caused OMU to
consider disposal alternatives other than an OMU-owned landfill.

To access the potential "market" for plant by-products, OMU and FMSM developed a

request for proposal (KFP) that was issued to landfill operators, bulk material
contractors and wallboard manufacture's. The RFP emphasized maximum utilization

103-2

-------
Table 1

Summary of Project

Electrostatic Precipitators

•	Unit 1

Type
SCA

Outlet emissions

•	Unit 2

Type
SCA

Outlet emissions
Reactant Preparation

•	Type

•	Number/capacity

•	% capacity each

•	Fineness
Absorbers

•	Number/capacity
Type/additives
S02 removal efficiency
L/G ratio

Material of construction
Primary Dewatering
Type

Solids in/out
Secondary Dewatering
Vacuum filter type
Number/capacity
Size

Filter area

Vacuum rapacity/hp
Solids at discharge

Design Features

Rigid electrode
238 tf/IOOO acfm
0.2 Ib/MBtu

Rigid electrode
266 ft2/"! 000 acfm
0.2 Ib/MBtu

Wet ball mills
2/16t/h
75%

95% through 325 mesh

2/67%

Limestone spray tower/none

96%

132 gpm/1000 acfm
317 LMN stainless

Thickener
12%/30%

Rotary drum
3/50%

6'0 x 8'L
151 ft2 each
530 acfm/30
75%

103-3

-------
of all plant by-products including bottom ash and fly ash. The RFP addressed both
disposal-grade gypsum that the process was currently designed to produce and
wallboard grades of gypsum that could be made in the system with modifications to
equipment. Two purity levels of gypsum were outlined in the RFP based on the
expected range of impurities in the potential limestone sources.

Bids were received in response to the RFP in July 1992. The bids encompassed total
by-product management (ash and gypsum marketing with disposal of balance) and
smaller scopes such as gypsum utilization with transportation by others. A total of
14 bids were received, and 7 were short-listed: 2 for total management, 3 for disposal
only, and 2 for gypsum utilization. An addendum was issued to the short-listed
bidders to obtain final pricing.

To support the OMU and FMSM effort, S&L and WAPC were requested to evaluate
technical requirements for wallboard quality gypsum based on technical information
provided by wallboard manufacturers in their bids. A summary of the wallboard
manufacturers' technical requirements are summarized on Tables 2 and 3. The
following sections discuss the technical evaluation of wallboard quality gypsum, the
design modification required to plant equipment to produce wallboard quality
gypsum, the results of the economic evaluation of disposal, and utilization
evaluation.

Table 2

Wallboard Manufacturer Technical Requirements

Mfq. A

Mfq. B

pH
Purity

Size criteria:

Free water

Min.

6

Max.

<10%
8

Desired

0% to 3%
6

97+%

Refected
>10%

8

<94%

Mean part
Surface area
Part area (X x Y)

20 Jim
3500 cm2/g

30 to 50 jim >70, <20p.m

>2000 jim >4000, <1000 Jim

Aspect ratio:

X to Y
Z to Y

Average particle
size (50 percentile)

<0.5, >0.1 <0.05, >0.75
30 to 50 jim >70 <20 Jim

2 to 5

>10

103-4

-------
Table 3

Specific Chemical Constituent Limits

Constituent

Fly ash

Mfg. A
Allowed
MaxJMin.
Limit

1.0%

Mfg. B

Allowed

MaxJMin.

Limit

Calculated
at

Maxiumum
Conditions

0.13% Max. 0.76%
(1250 ppm)

SiO,

1.1%

0.5%

0.41%

Fe.,0.

2W3

1.5%

Not given 0.45%

RA

3.5%

Not given 2.52%

Na

Not given 70 ppm

9 ppm

Mg

Not given 70 ppm

135 ppm

CaS03*%H20 1% max. 0.1%	0.72%

CI	120 ppm 100 ppm 100 ppm

Water soluble 600 ppm Not given Expected

<600 ppm

Worst Operating Conditions

0.2 Ib/MBtu ash loading and
60% fly ash collection in the
absorbers

60% FA collection, 51% Si02
in FA and 1.5% Si02 in LS.

60% FA collection, 43% Fe203
in FA and 0.2% Fe203 in LS.

60% FA collection, 25% Al203
and 43% Fe203 in FA, and
0.2% Al203 and 0.2% Fe203 in
LS.

25-250 ppm from Power
Magazine. Feb., 1988. 0.7%
Na20 in fly ash and 90%
solids and 90% wash
efficiency. 232 ppm Na in
46-gpm blowdown.

50-250 ppm from Power
Magazine. Feb., 1988. Based
on 90% solids, 90% wash
efficiency. 1.5% MgCOs in
limestone and 1 % Mg in ash.
Assumed that 50% of MgCOa
will dissolve.

At 99% oxidation, 97%
limestone utilization and 97%
CaC03

103-5

-------
Technical Feasibility Evaluation
Free Water

Free water is defined as the weight of the noncombined water divided by the total
weight of the wet sample. Excess moisture in the delivered product will require the
wallboard manufacturers to install special handling and drying equipment for the
synthetic gypsum that is not required for natural gypsum. As part of the wallboard
production process, the free water and a portion of the combined water must be
removed. Therefore, wallboard manufacturers prefer the lowest free water levels
possible with rejection levels ranging from 10%-15%. This point is the moisture level
where drying the gypsum is no longer economically feasible. Moisture will be
subject to penalties directly related to the costs of drying the gypsum.

Various vacuum filter options and centrifuges were considered to improve the solids
content of the final product with minimum changes to layout and equipment.
Utilizing rotary drum vacuum filters would provide 85% solids. Indexing, or carrier
belt-type horizontal filters and centrifuges would achieve in excess of 90% solids.
Horizontal belt vacuum filters have outperformed the rotary drum vacuum filters in
producing drier material and can achieve higher wash efficiencies due to counter-
current wash capabilities. A major consideration in the change from disposal-grade
to commercial-grade gypsum was minimizing the cost and schedule impact on the
project. The selection of vacuum belt filters for secondary dewatering equipment was
influenced by the fact that thickeners had been purchased and thickener vendor
engineering was complete. In addition, the water balance and distribution system
was designed without provision for a second high solids reclaim water system.
Technical evaluations indicated that thickeners in combination with vacuum belt
filters would produce an acceptable product. Financial and schedule evaluations
indicated that the thickener filter combination was the most cost-effective. The
indexing belt-type filter was selected over the carrier belt type after analyzing the
projected maintenance costs over a 5-year period. Material that is drier than the 90%
target will have a higher value. However, dusting of the gypsum during handling
and transportation may become a problem if the free water is less than 6%. A
centrifuge could also be used to produce drier than 90% solids. Centrifuges would
be expected to require a higher capital cost than both the rotary drum and belt
vacuum filters. The evaluation concluded that the horizontal indexing belt filter was
the most economic choice since guarantees of 90% solids were provided with
expected production between 91%-92%.

Sargent & Lundy recommended that a 4-day working stack be maintained to allow
for production and shipping flexibility. A storage building would be needed to
control moisture penalties, reduce potential contamination, arid minimize dusting and
product loss. Silo options were also investigated but proved to be more expensive
and less reliable than a more conventional pre-engineered building approach.

103-6

-------
Gypsum Purity

Gypsum purity is defined as the weight of CaS04'2H20 divided by the weight of the
dry sample. The minimum gypsum purity required by wallboard manufacturers is
between 94% and 95%. The purity at this level is primarily a market driven function.
Natural gypsum reserves can run in the 80%-90% range. The major concern for
wallboard production is not the level of gypsum, but is the makeup of the impurities.
The gypsum by-product purity can be optimized by maximizing oxidation, increasing
the available CaC03 in limestone, improving limestone utilization, and reducing fly
ash loading to and removal by the FGD system. Each of these items is discussed
separately.

Oxidation. Oxidation is defined as the total moles of CaS04*2H20 divided by the
sum of the total moles of CaS04*2H20 and CaS03*l/2H20. The higher the oxidation,
the higher the production of CaS04«2H20, which increases the gypsum contribution
to the total sample resulting in a higher purity.

A minimum of 99% oxidation is required to produce 95% pure gypsum. The original
sparger design was based on 98% oxidation of calcium sulfite to calcium sulfate.
WAPC's original design was to force oxidation to occur with one absorber operating
at 67% load. The design air stoichiometry was 2.4. Three 50% compressors were
provided in the original design. In changing the design to commercial-grade
gypsum, a 99.3% oxidation rate is required. In order to achieve the required
oxidation, the air stoichiometry was increased to 3.0. All three compressors must be
operated during MCR conditions to produce the required oxidation required for
commercial-grade gypsum.

The original sparger consisted of 4-inch pipes that were selected to minimize pressure
losses through the pipe to ensure adequate pressure and distribution of air to the
absorber. In changing to commercial-grade gypsum design, some of the sparger
pipes were increased to 6 inches in order to accommodate the higher air flow rate at
the original pipe pressure losses and acceptable distribution.

The current FGD design has a large solids residence time in the reaction tank in
addition to 32 feet of liquid level above the air sparger. OMU's anticipated 80% load
factor will further increase the average residence time. This design may achieve 99%
oxidation without requiring operation of the third compressor. However, provisions
were incorporated to allow three compressor operation.

Available CaC03 in Limestone. The available CaC03in the limestone is defined as
the percent of CaC03 dissolving in a laboratory experiment. A higher CaC03 content
in the limestone will reduce the quantity of inerts and other impurities such as SiOz,
Fe2Os, and R203, which will increase the gypsum purity. Table 4 presents the

103-7

-------
Table 4
Limestone Utilization

Gypsum
Purity. %

94

Available CaC03 in
Limestone, %

96

97

95

Required LS
Utilization, %

97.4

96.4

95.4

95

95

99.2

96

98.1

97

97.0

required limestone utilization at 95% through 97% available CaC03 content to
produce gypsum purities of 94% and 95%. The required limestone utilization
numbers are calculated based on 99% oxidation and 60% fly ash collection by the
absorbers at a loading of 0.2 lb/MBtu. The original limestone utilization guarantee
was based on a minimum 90% CaC03 content in limestone.

Table 4 indicates that the lower the limestone utilization value, the higher the CaC03
content required. For example, to produce a 95% pure gypsum, available CaC03
values of 95% and 97% would require very high utilizations of 99.2% and 97.1%
respectively. The table shows that a maximum gypsum purity value of 94% should
be contemplated to allow a realistic variation in limestone quality.

Limestone Utilization. Limestone utilization is a measure of the available CaC03
that reacts with the collected S02. The higher the limestone utilization, the lower the
unreacted CaC03 and MgC03 in the final gypsum product leading to higher gypsum
purity. Some amount of MgC03 in the limestone may react. The MgC03 credit is not

considered since the amount reacted depends on the reactivity of the specific

Table 5 presents the expected gypsum purity for available CaC03 between 94% and
97% at the design basis limestone utilization. A 96% limestone utilization can be
achieved for limestone with available CaC03 values between 95% and 97%. The
gypsum purity numbers are calculated based on 99% oxidation and 60% fly ash

collection by the absorbers at a loading of 0.2 lb/MBtu. The data indicate that the

limestone.

103-8

-------
Table 5

Expected Gypsum Purity

Available CaC03 in
Limestone Utilization, % Limestone, % Gypsum Purity, %

96

95

93.2

96

96

93.8

96

97

94.4

higher the CaC03 content, the higher the gypsum purity. For example, a limestone
with 97% CaC03 at 96% utilization will produce 94.4% gypsum purity.

A sulfuric acid addition system at the filter feed tank or upstream of the thickeners
could dissolve some of the unreacted carbonates (CaC03 and MgCOs) resulting in a
higher gypsum purity. For example, if 98% CaCOs utilization is achieved by acid
addition then 96% CaC03 limestone can be used to produce 95% pure gypsum. The
handling of acid adds to the complexity of the process. Also, if acid addition is not
monitored properly, corrosion problems can be created. Furthermore, acid addition
will add significant operating costs since 93% H2S04 costs approximately 4
-------
Fly Ashm In general, fly ash from properly operating precipitators will not impact the
wallboard quality of gypsum. Wallboard manufacturers' specified limits are between
0.13% and 1.0% If the fly ash loading to the FGD system is at the precipitator outlet
guarantee levels of 0.2 lb/MBtu and 60% of the ash is collected by "the scrubbers then,
the fly ash in gypsum would be 0.76%. In such a case, the manufacturers'
requirements would be impossible to meet. Fly ash loading is a function of the
precipitator design that could not be modified at that point in the project.

Excessive fly ash may cause whiteness problems; however, color consistency seems to
be a larger concern for U.S. wallboard manufacturers.

Chloride

High chlorides can cause paper adhesion problems in the manufacture of wallboard.
Wallboard manufacturers' requirements on chlorides in the gypsum are less than 100
to 120 ppm. This requirement can be met by washing the gypsum cake. The
recovered chloride will have to be purged to maintain the required chloride level
(<8000 ppm) appropriate for the current materials of construction. A continuous
blowdown from the filtrate pump (adjusted based on laboratory analysis) is required.
The indexing belt filter offers an 88% wash efficiency. A higher than 88% efficiency
can be achieved by adding an additional wash box to the belt filters at a later date if
needed.

The blowdown rate is inversely proportional to the wash efficiency. Therefore, the
higher the wash efficiency the lower the blowdown rate. Figure 1 is a graph of
blowdown versus chloride to sulfur ratio at 88% wash efficiency for various weight
percent solids. At 90% solids and 88% wash efficiency a blowdown flow rate of 50 to
280 gpm would be required (for Cl/S ratios of 0.02 to 0.1 respectively) to keep the
chloride level below 8000 ppm. A permit for blowdown up to 300 gpm is currently
being sought by OMU.

Chemical Composition of Gypsum

Combined Water, CaO and SOT The chemical formula of gypsum is CaS04-2H20.
Wallboard manufacturers have expressed the total gypsum purity of 95% into three
individual chemical species such as combined water, CaO and SC)3. The minimum
quantity of 19.9% combined water, 30.9% CaO and 44.2% S03 add up to total
minimum 95% CaS04«2H20 purity specified.

103-10

-------
Gypsum Cake @120 ppm d- with wash efficiency at 88%
Chloride Equilibrium Concentration - 8000 ppm

Slowdown, GPM

Chloride/Sulfur Ratio

Figure 1. Blowdown Versus Cl/S Ratio

Specific Constituent Limits. Table 3 is the list of constituents' maximum/minimum
allowable limits from wallboard manufacturers and the expected level at worst
operating condition at the Elmer Smith FGD installation. Specific items on the table
are addressed as follows:

•	As discussed earlier, expected fly ash level in the solids would be 0.76% and
0.7%, below a manufacturer's 1% requirement, at 0.2 and 0.185 lb/MBtu fly ash
loading respectively based on 60% collection by the scrubbers. The lower
requirement of 0.13% fly ash would be practically impossible.

•	Wallboard manufacturers' maximum Si02 limit of 0.5% and 1.1% will not be met
if the limestone contains greater than 2.5% Si02. High level at Si02 are a concern
to the wallboard manufacturers due to equipment wear. The wallboard
manufacturers have indicated that the Si02 limit can be exceeded if a cost
penalty is levied.

103-11

-------
•	R203 is a total amount of inerts including all metal and silica oxides but
excluding other inerts in fly ash and carbonates (unreacted CaC03 and MgC03).

•	One manufacturer's Na requirements of 70 ppm will be easily met but the 70-
ppm Mg level may be difficult to achieve. Other manufacturers have not given
any requirements for either Na or Mg. Instead of individual constituents they
look at total dissolved salts, which is discussed below.

•	One manufacturer's 1% maximum CaS034£H20 requirement will be met. The
expected level will be 0.72% at 99% oxidation. Other requirements of 0.1% are
very stringent and would require 99.86% oxidation.

•	The total water soluble salts should be substantially lower than the maximum
600-ppm requirement.

All the constituents listed in Table 3 are projected to be lower than the maximum
allowed limit by one manufacturer. The maximum limit by other manufacturers is
very stringent and will be difficult to achieve.

Other requirements (such as, aspect ratio, particle size, NH4, and ppm) are difficult to
monitor and, therefore, are not included in all manufacturer's specifications. A
minimum particle size requirement of 20 pm seems reasonable and should be
achievable. All other requirements, such as aspect ratio and x-y axis dimensions,
should be excluded from manufacturers' specifications.

Technical Feasibility Results

The technical feasibility evaluation concludes that wallboard quality gypsum is
technically feasible on the basis of the following:

•	Horizontal belt filters will be used for secondary dewatering in lieu of rotary
drum filters. Solids contents of 90% can be provided with a horizontal indexing-
type filter. The filter will be equipped with a two-stage wash system allowing
for a chloride level in the gypsum product that will not be greater than 120 ppm.

•	A continuous blowdown from the filtrate pump of the vacuum filter based on
chloride analyses is incorporated to control chloride level in solids and to
maintain the equilibrium level appropriate for the current materials of
construction. A blowdown stream of up to 300 gpm should be assumed. This
blowdown stream will also allow flexibility in coal procurement.

•	Limestone reagent with as high a CaC03 content as possible should be
purchased (a minimum of 96% CaC03 content may be required to meet a 94%
purity requirements). A higher CaCO, content in the limestone will increase
limestone utilization.

103-12

-------
• A minimum 99% oxidation can be achieved while operating the three 50% air
compressors that had been previously purchased. Sufficient conservatism exists
in the oxidation process design that may allow operation of only two
compressors.

Economic Results
Economics Analysis

Table 6 summarizes the evaluation of capital and operating costs associated with the
alternatives considered for disposing or utilizing the FGD system by-product. To
focus on the cost for alternate FGD disposal methods, the evaluation results
presented in Table 6 exclude any of the costs associated with reuse or disposal of fly
and bottom ash. The results shown in this table represent the optimum combination
of loading, hauling, and disposal and reuse bid costs based on the 14 bids received
byOMU.

The capital costs shown for the commercial-grade gypsum alternative represents the
incremental investment required for process and balance-of-plant supporting
equipment over the base cost for the rotary drum vacuum filter process equipment.
The costs considered in the process equipment category shown in Table 7 include the
incremental belt filter investment, air sparger modifications, and deletion of the radial
stacker and associated conveyors. In addition, building size and electrical power
supply system costs were impacted by the larger size of the belt filters and the larger
power requirement of the vacuum pump and auxiliaries. The total cost impact to the
process equipment is approximately $5.4 million. The $5.4 million incremental
investment compares to the $20 million investment required to develop an
environmentally acceptable landfill. No additional capital investment increment
would be required for a third-party landfill. The cost for a third-party landfill
operator to develop the landfill are included in the operating costs associated with
this alternative.

Operating costs included in the economic evaluation include costs for loading,
transporting, disposal, and/or utilization of final gypsum product for each
alternative. The commercial-grade gypsum evaluation assumes that 95% of the
gypsum will be sold for wallboard production. The remaining 5% is assumed to be
material that does not meet specification that would have to be disposed of in a
third-party landfill. For the commercial-grade gypsum alternative, an on-site barge
loading facility, owned and operated by a third-party contractor, was determined to
be the lowest risk economic choice. The overall cost of an OMU owned and operated
barge loading facility was comparable to the third-party operator costs, however,
OMU favored a lower initial capital investment approach to minimize the risk
associated with possible changes in the long term market for gypsum. Note, to avoid
compromising negotiations that are still in progress with potential gypsum
contractors, a breakdown of the operating cost components cannot be provided.

103-13

-------
Table 6

Economic Comparison of Solid Waste Disposal Alternatives

Commercial-
OMU-Owned Third-Party	Grade

Landfill	Landfill	Gypsum

Capital costs (1995$)

A. Landfill dev.

20,000,000

N/A

N/A

B. Sec. dewatering equip.

Base

Base

2,200,000

C. Gypsum stockout build.

N/A

N/A

2,108,000

D. Conveyors

Base

Base

434,000

E. Contingency

N/A

N/A

662.000

Subtotal

20,000,000

0

5,404,000

O&M Costs (1995 $)







A. Loading, trans., utilization







or disposal

1,906,000

2,763,000

1,245,000

B. Power

11.000

11.000

76.000

Subtotal

1,917,000

2,774,000

1,321,000

Limestone (1995 $)

1,016,000

1,016,000

1,632,000

IV.	20-Year PVRR	58,731,000 50,048,000 44,619,000

V.	20-Year PVRR Over Base	14,112,000	5,424,000	Base

The limestone cost evaluation presented in Table 6 represents the expected difference
between the higher quality stone required to meet gypsum purity requirements and
the lower quality, higher inert content percentage limestones.

OMU is currently in the process of procuring limestone for the project. Bids have
been obtained from suppliers in Kentucky and neighboring states. Preliminary
indications are that the premium for the higher quality stone can range to 100% of
the cost of the lower quality stone.

103-14

-------
Table 7

Development of Commercial-Grade Gypsum Process-Related Costs

Landfill Options Commercial-Grade Gypsum

A.	Belt filter

B.	Air sparger

C.	Delete rad. stacker

D.	Sec. dewatering building

E.	Electrical equipment

Total

Base

Base

Base

Base

Base

Base

50,000
(1,000,000)
400,000
250,000
2,200,000

2,500,000

The economic analysis shows a payback period of approximately 5 years on the
investment to convert to commercial-grade gypsum production. This payback was
sufficient for OMU to conclude that the commercial-grade gypsum conversion could
be justified. WAPC and S&L were authorized to proceed with the process and
balance-of-plant equipment changes in September 1992. OMU is currently in the
process of finalizing negotiations for gypsum purchase, barge loading services, and a
backup third-party disposal site.

Revised Equipment Arrangement

The arrangement of the revised secondary dewatering facility and gypsum stockout
building are shown on Figures 2 and 3. The three 50% capacity rotary drum vacuum
filter in the original design are replaced with two 100% capacity horizontal belt-type
vacuum filters. Each belt filter is designed for a capacity of 73,200 lb/hr and consists
of a filter assembly approximately 12-feet wide and 79-feet long weighing 12 tons, a
vacuum pump /washwater system skid approximately 12-feet wide x 16-feet long x
10-feet tall and a filtrate pump/receiver skid approximately 8-feet wide x 24-feet long
x 12.5-feet tall. The filters use a seamless cloth 2.5 meters wide with a filtration
length of 21 meters. The process flow for one belt filter is shown on Figure 4. The
filters will be continuously fed slurry by a recirculation loop from the filter feed
pumps to the filter feed tank. Modulating slurry control valves will be used to
control slurry feed to the filters. The length of the filter is broken down into a
number of process steps including a pool section, cake forming, first cake wash, first
cake dry, second cake wash, and final cake drying sections. "ITie filter cloth and cake
are moved from one processing step to the other as the filter indexes in increments of
approximately 30 inches. A countercurrent wash design in incorporated into the
filter system that uses wash boxes to distribute the washwater across the filter cake.

103-15

-------
Figure 2. Revised Secondary Dewatering Facility—Second Floor Plan

Spent vacuum pump seal water and spent cloth washwater are used for second-stage
cake washing. Filtrate water from the second stage wash is used for first-stage
washing.

Two 72-inch-diameter x 96-inch-high filtrate receivers are provided for each filter.
The primary receiver collects slurry filtrate and first-stage washwater for
transportation to the thickeners and for chloride blowdown. The secondary receiver
collects filtrate from the second wash stage for use in first-stage washing. Both
vacuum filtrate receivers are fabricated from rubber-lined carbon steel.

One filtrate pump is provided with each receiver. Filtrate pumps are constructed of
rubber lined carbon steel construction with gear box drives.

103-16

-------


| tew*. • * srocHour suiidinc

Figure 3.

Gypsum Conveyor System and Stockout Building

-------
w

00

Figure 4. Revised Secondary Dewaterlng Process Flow

-------
Each filter system includes one liquid ring vacuum pump designed to deliver 9040
acfm at 20 inHg and 450 bhp. The pumps come complete with a direct drive gear
box and 500-hp drive motor. Seal water requirements of the pump are minimized by
a 50% recycle of seal water. The seal water system also provides spent seal water
with an elevated temperature beneficial for use in second-stage cake washing.

Vacuum pump seal water and spent cloth washwater are collected in a washwater
tank of fire resistant FRP construction. A supplemental electric heater is provided in
the tank for use during system startup and when city water supply temperatures
drop below 40 °F. One washwater pump is provided to pump washwater to the
second-stage wash box.

Each vacuum filter includes a local control panel with a PLC for control of vacuum
filter indexing operations. Supervision of the PLC system will be by the owner's
distributed control and monitoring system (DCMS) via various alarm and status
signals transmitted from the PLC. All other belt filter system equipment will be
controlled through the DCMS.

The conveyor system will be designed to permit the simultaneous operation of both
vacuum filters. Each vacuum filter discharges gypsum product cake through a chute
and onto the reversing transfer conveyor for transportation to the forwarding
conveyor and finally the stacker conveyor system. The conveyor system is capable of
handling both vacuum filters simultaneous operation. The reversing transfer
conveyor can also discharge onto the emergency discharge conveyor with a final
emergency cake discharge elevation approximately 12 feet above grade. The gypsum
is conveyed from the secondary dewatering building to the gypsum stockout
building. A reversing shuttle conveyor in the stockout building allows material to be
stacked at any point in the building. The gypsum is removed from the pile and with
a front-end loader and moved by truck to an on-site barge loading facility.

Conclusions

There are several advantages to the production of a commercial-grade gypsum versus
a disposal-grade gypsum, but the advantages are not without some risks. In
addition, a decision to produce commercial-grade gypsum requires a site-specific
study. Some general conclusions from the OMU study are as follows:

•	The production of waliboard-grade gypsum is technically feasible. The

dewatering system is the main area requiring major capital investment.

•	The liability associated with a waste product are minimized with the wallboard
approach versus disposal in a third-party landfill.

•	The major economic factors are dewatering system and gypsum storage capital
investment, high-grade limestone costs, transportation costs to the wallboard

103-19

-------
plant, and gypsum utilization fees versus alternative costs (e.g., disposal,
alternate dewatering system, and limestone).

•	The economics are very sensitive to the long-term market for synthetic gypsum.
Length of contract and price adjustment clauses are very important factors to be
considered. The gypsum utilization fee in a pure economy should be a function
of competing gypsum mine costs and competing utilities. Only very close
power plants could compete economically with a gypsum mine located at the
wallboard plant Competition by utilities closer to wallboard plants could also
defeat the favorable economics of this modification.

•	Optimum economics occur when scrubber production matches wallboard plant
capacity to allow 100% synthetic gypsum wallboard production.

•	Maintenance of gypsum quality must be a prime consideration in addition to
removal efficiency when operating a FGD system.

•	A landfill should still be permitted to accept off-spec gypsum or in case of force
majure or contract termination.

Reference

1. W. Ellison and E. Hammer. "FGD-Gypsum Use Penetrates U.S. Wallboard
Industry." Power, pp 29-33. February 1988.

Bibliography

Ransom, J. M., Torstriek, R. L., and Tomlinson, S. V. "Feasibility of Producing and

Marketing Byproduct Gypsum from S02 Emission Control at Fossil-Fuel-Fired Power
Plants." EPA-600/7-78-192, October 1978.

Armour, D. W., Conn, S. K., and Frizzell, K. "By-Product Management Philosophy at
the Elmer Smith Station." Proceedings of the American Power Conference. 1993.

Collins, S. "Managing Powerplant Wastes." Power, pp 15-27. August 1992.

103-20

-------
PLANT GROWTH AND SOIL PROPERTIES RESPONSES
TO ADDITIONS OF DRY FLUE GAS DESULFURIZATION BY-PRODUCTS

W. A. Dick, R. C. Stehouwer, P. Sutton, J. M. Bigham, R. Lai,

S. J. Train a, E.L. McCoy and R.. Fowler
Department of Agronomy
The Ohio State University
Wooster, OH 44691

Abstract

Substitution of alkaline flue gas desulfurizatlon (FGD) by-products for conventional
liming materials in agriculture, mine reclamation and for soil stabilization are
potential uses for these by-products. Greenhouse and field studies demonstrated
improved plant growth when add soil or surface coal mine spoil were treated with
FGD by-products from three different sources. For agricultural soils, optimum
growth occurred when application rates were equal to the lime requirement of the
soil. For the acid spoil, optimum growth occurred when application rates were
between 6 and 12% (by weight). Leachate composition indicated that when FGD by-
products were applied at rates of 12% or lower, concentrations of elements of
environmental or regulatory concern remained vety low. Boron was an exception
when LIMB by-product was used to amend spoil and a period of time for the boron
to be leached from the amended spoil prior to planting of vegetation cm* crops may be
required when high application rates of LIMB by-product are used. There seems to
be little potential for adverse effects on water, soil and plant quality when FGD by-
products application rates are based on the amount required to neutralize the acidity
in soil or spoil. The improved plant growth observed in this study, when FGD by-
products ware applied to acid soil or spoil, represent a beneficial reuse of these by-
products.

Introduction

A number of diy FGD technologies have been developed in the past decade. The US
Department of Energy's Clean Coal Technology program and the 1990 amendments
to the Clean Air Act, mandating a 2-stage 10 million ton reduction in SO2 emissions
in the United States, have especially encouraged the development of these
technologies. Dry FGD technologies are generally smaller in scale and require a
lower capital investment than do the wet FGD processes. Dry FGD technologies are
also generally designed for retrofit on existing coal-fired power plants and, therefore,
represent an option for bringing older plants into compliance with clean air
legislation.

The classification of the diy FGD by-product as a solid, or even hazardous, waste by
most states has inhibited the development of beneficial uses of these materials.

Thus, in addition to installation and operating expenses, plants burning high sulfur
coal and using dry FGD technologies must also bear increasingly expensive landfill
disposal costs. We have recently completed a comprehensive study of the chemical,
physical, mineralogical and engineering properties of 58 dry FGD by-product
samples in Ohiol. With this information serving as a base, a greenhouse study was
conducted to demonstrate potential reuses of diy FGD by-products. These reuse

104-1

-------
applications, if environmentally benign or even beneficial, would significantly
reduce the costs of SO2 emission control.

Dry FGD by-products are composed of a mixture of conventional coal combustion
ash (either bed or fly ash), the SO2 reaction product (primarily anhydrite, CaSOa),
and unspent sorbent (generally lime, limestone, or dolomite). Due to the presence
of unspent sorbent, dry FGD by-products are usually highly alkaline with significant
neutralization potentiaP-5, Anhydrite does not neutralize acidity but may
ameliorate problems of A1 toxicity through the formation of Al(SC>4)x complexes. It
is also a source of both Ca and S for plant nutrition. Fly ash may supply other plant
nutrients (e.g. B, Mo, Zn, P, K)2. These FGD by-products may also be high in soluble
salts and contain some trace elements of environmental concerns.

Of the several dry FGD technologies tested in Ohio, four were selected for more
intensive by-product characterization: duct injection, lime injection multistage
burner (LIMB), pressurized fluidized bed combustion (PFBC), and spray dryer. Duct
injection involves injecting the sorbent (hydrated lime) into the flue gas as it enters
a humidification chamber in the ductwork downstream of the boiler and air heater.
Injection of fine water mist lowers the gas temperature and raises the humidity of
the flue gas to increase the reaction of the SO2 with the sorbent. Dry FGD by-product
is removed from the flue gas and contains the reaction products, fly ash, and
unspent sorbent. LIMB is a process whereby calcium based sorbent is injected
directly into the boiler where it calcines to CaO and reacts with SO2 and O2 in the
combustion gases to produce CaS04- The reaction product and unspent sorbent are
collected with the fly ash. In PFBC systems, a cataum-based sorbent (usually
limestone or dolomite) and crushed coal are introduced together into the boiler bed
where they are "fluidized" or suspended by jets of air. This mixes the coal and
sorbent and allows for reaction of SO2 and sorbent. Two by-product streams are
created - one being the heavier, granular bed ash material and the other the finer
materials which are removed by the particulate emission control equipment The
spray dryer scrubber includes a separate vessel located downstream of the boiler and
air heater where sorbent, usually a slurry of hydrated lime, is injected. Reaction
products and unspent sortsent are removed with the fly ash in the particulate
emission control system.

Based on the results of laboratory investigations of the properties of FGD by-
products, we proposed several greenhouse or tightly controlled field studies to
further develop beneficial reuses of these FGD by-product materials. The
applications tested and reported in this paper include using FGD by-products as an
alkaline amendment for acid agricultural soils and acid spoil from surface coal
mines. Growth responses and environmental parameters were measured.

Materials and Methods
Soil and Spoil

The soil used was a Wooster silt loam (fine-loamy, mixed, mesic, Typic Fragiudalf).
It had a pH of 4.6 and a lime requirement of 215 Mg/hectare (9.6 tons/acre) to reach
a pH of 7. The cation exchange capacity was 15.1 cmolc /kg (15.1 milliequivalents/100
g soil) and this soil was noted for Mn toxicity.

Spoil was obtained from the Fleming abandoned mined land (AML) site near New
Philadelphia in Eastern Ohio. This site contains approximately 10 hectares (25 acres)

104-2

-------
of exposed, highly erodible underclay bordered on two sides by approximately 18
hectares (45 acres) of unreclaimed spoil and 2 hectares (5 acres) of coal refuse. The

fH of the underclay ranges between 2.5 and 35 and that of the spoil between 32 and
0. Both the underclay and spoil were collected for use in the greenhouse study, but
only the results related to the spoil are reported.

Dry FGD By-Products Used

By-products from the LIMB, PFBC and spray diyer FGD technologies were used. The
LIMB by-product was collected from the full-scale commercial LIMB demonstration
at Ohio Edison's Edgewater plant in Lorain, Ohio. The PFBC by-product was
obtained from the demonstration facility operated by American Electric Power at the
Udd plant located near Brilliant, Ohio. The McCracken Power Plant at the Ohio
State University was the source of the spray dryer FGD by-product.

Detailed characterization data for the LIMB and PFBC by-products are summarized
in Table 1.

Experimental Procedures

After passing through a 10 mm screen, 48 leg erf air-dry Wooster soil was mixed with
FGD materials at rates of 0, 0.35, 0.7,1.4 and 2.8% by weight and placed into a FVC
pot (15 an diameter, 30 can deep). These rates approximate 0, 7.8,15.6, 31.2 and 624
Mg/hectare (3.5, 7.0,14, and 28 tons/acre) or 0, 0.25, 0.5,1 and 2 times the soil liming
requirement. Since the FGD by-products contained only 60% calcium carbonate
equivalency (Table 1), the rates required for neutralization of acid in the soil, as
compared to agricultural lime, were approximately double. Separate pots of each
mixture were planted with fescue and alfalfa. The treatments were replicated three
times.

The spoil from the Fleming AML site was air-dried, passed through a 10 mm screen,
and mixed with FGD by-product at rates of 0, 3,6,12 and 24% by weight
(approximately 0, 30, 60,120 and 240 tons/acre) using a complete factorial
experimental design. Sewage sludge (Table 1) was mixed with all treatments at a
constant rate of 6% by weight The sewage sludge was included because it has teen
our experience that addition of an organic material with alkaline material creates an
excellent amendment for reclamation of acid spoil and growth of vegetation. In
addition, four replications of spoil alone or mixed with 12% FGD by-product, but
without sewage sludge, were also included in the experiment. Mixtures were placed
in FVC pots (15 cm diameter and 30 cm deep) and the plots planted to Kentucky 31
fescue. The treatments were replicated four times and arranged in randomized
complete blocks.

The field experiment at Coshocton, Ohio involved inserting 21 PVC columns (30 an
diameter and 90 am deep) into the ground. The columns were filled (torn the
bottom to the top) with a layer of sand (5 cm), untreated Fleming spoil, and then
Fleming spoil treated to approximately 20 cm depth with either LIMB by-product (10,
20, 40% by weight) or spray diver FGD (10, 20, and 40%, by weight). An untreated
control was also included in the experimental design. An alfalfa and fescue
mixture was planted in the columns.

104-3

-------
Data Collection and Analyses

Similar data collection procedures and laboratory analyses were applied to materials
from pots containing treated Wooster soil or Fleming AML spoil,

Leachate was collected from pots containing treated soil and spoil at various times
(including before planting) by adding enough water to collect from 100 to 200 ml of
volume. Leachates were analyzed for pH, electrical conductivity, metals (AJ, As, B,
Ba, Ca, Cd, Co, Cr, Cu, Fe, Hg, K, Li, Mg, Mn, Mo, Na, P, Pb, S, Sb, Se, Si, Sr, V, and Zn
by taductivelv coupled plasma spectrophotometry (ICP)), and anions (F, CI"' NO3*,
SO3 and SO4 by ion chromatography (IC)).

Immediately following the first leaching the pots were weighed to determine water
content at field edacity. Soil water content in the pots was maintained at
approximately 25% below field capacity throughout the experiments to insure that
plants were not under moisture stress, that reducing conditions did not occur in the
pots, and that no leachate drained from the pots between intentional leaching^.
Artificial lighting was used to provide 14 hours of light, and the temperature was
maintained on a 25/20*C day/night cycle.

Hants were harvested eveiy 30 to 32 days with the first harvest occurring 94 days
after planting and continuing for six cuttings (approximately nine months after the
mixes were prepared. Cuttings were dried at 65°C, weighed, ground and digested
with perchloric add. The digests were analyzed for the same metals, using ICP, as
listed for the soils.

After nine months, the pots were split using a saw and the treated soil or spoil
removed. Various chemical, physical and mineralogical properties were measured
and the results compared across treatments. Only some of the data collected at the
conclusion of the experiment will be reported in this paper.

Rant growth at the Coshocton field site was determined 125 days after seeding. The
amount of dry matter production as alfalfa or fescue was determined.

Analyses of variance (ANOVA) was conducted on the soil, spoil and leachate data
collected.

Results and Discussion
pH, Alt and Mn Changes

The primary benefit of liming an add plant growth medium is increasing the pH to
near neutrality and redudng the potentially toxic concentrations of A1 and Km. The
FGD by-products used in our tests were effective in rapidly raising the pH of the
Wooster soil and the Fleming AML spoil (data not shown). For example, untreated
spoil maintained a pH of approximately 3.0 and spoil treated with 6% sewage sludge
obtained a pH of 40. Spoil treated with 6% sewage sludge and 12% FGD, however,
resulted in a pH of 8.3 (LIMB by-product) or a pH of 8.0 (PFBC by-product). The pH
of all treated soil and spoil materials decreased over time, with the largest decreases
occurring near the column surfaces. After nine months, pH in the surface 5 cm
layer of the pots containing spoil never exceeded 8.0, even when FGD by-product
was applied at the highest rate (24% by weight). This pH would not restrict plant
growth, as many agricultural soils in the western states of the United States have

104-4

-------
natural pH values of 8.0. Also a soil pH of 8.0 would be much more conducive to
increasing overall plant growth than the pH originally observed for the soils and
spoil materials used in this experiment. Thus, the addition of FGD by-product to
acid soil or spoil would represent a beneficial reuse.

The leachate chemistry is a reflection of the chemistry that occurs in the pots. pH in
the first leachate from treated soil and PFBC treated spoil was ^proximately 8.u at
the highest application rates (Figures 1 and 2). The leachate pH from the LIMB
treated spoil, however, varied from 7.0 (3% application rate) to 12.0 (24% application
rate) in the first leachate. With time, leachate pH remained either unchanged or
tended to decrease (especially for the LIMB by-product treated spoil). After nine
months, the final pH or leachate at the highest PFBC and LIMB by-product
application rates (24%) were 77 and 85, respectively. This decrease in pH is
apparently due to carbonation of lime, portlandite, and peridase since the pH was
tending toward that of free carbonates in equilibrium with atmospheric carbon
dioxide. Carbonation was somewhat inhibited in pots with high FGD application
rates, however, because high pH and soluble salt concentrations inhibited plant root
growth and microbial activity in these pots. Plant root and microbial activity act as a
source of carbon dioxide and leads to increased carbonation of the FGD by-products.

Conditions of high pH and adequate moisture also promoted the formation of
cementitious secondary minerals. Extensive enttringite (Ca6Al2(SQ|)3(OH)i2'
26H2O) formation occurred with LIMB by-product above the 12% application rate.
Enttringite is highly expansive and cementititious and thus severely restricted root
growth and water movement below the depth at which it formed. At the 24% LIMB
by-product application rate, extensive enttringite formation occurred below 15 cm
depth (Figure 3). The PFBC by-product did not promote enttringite formation to the
same extent as did the LIMB Dy-product. This very different behavior of the two
FGD by-products with respect to secondaiy mineral formation is due to the differing
mineralogies and solubilities of their constituent components.

As previously mentioned, one of the primaty benefits of adding an alkaline
amendment to acid soil or spoil is to decrease the potentially toxic concentrations of
A1 and Mn. Concentrations of A1 and Mn in the leachates collected from the pots
were determined. The first leachates were collected from previously air-dry
materials and are not representative of equilibrium conditions in the soil or spoil
plus FGD by-product mixes. The final leachates, however, provide a much better
taction of equilibrium conditions and long-term impacts of using FGD by-
products.

Leachate concentrations of A1 and Mn decreased with increasing application rates of
FGD by-products (Figures 1 and 2). The highest concentrations of Al and Mn were
found in the untreated Wooster soil and Fleming AML spoil. FGD by-product
additions thus created an improved plant growth medium. The solubility of Al and
Mn is greatest under acid conditions® and decreases rapidly with increasing pH. The
beneficial effect of FGD by-products is due to increasing the pH of the system.
Application of sewage sludge at the 6% rate resulted in large decreases in Al and Fe
(data not shown) even though the corresponding pH increase was slight (from 3.0 to
4.0). This is due to the strong complexes that organic matter forms with these
elements?. An inverse relationship between soil organic matter and AP+ solubility
and plant growth toxicity has been reported^. These results support our previous
recommendations where we maintain that best results occur when a mixture of
organic and inorganic amendments are usedlO.

104-5

-------
Sulfur Concentrations in the Leachate

Increasing application rates of FGD by-products increased soluble salt concentrations
(as determined by electrical conductivity; data not shown). Similar patterns were
seen for S, Ca, and Mg which were the major elements in the leachate. Only S
results are reported (Figure 4) but torn these results we can also gain information
about the fate of other elements, such as Ca and Mg, in FGD by-products applied to
soil and spoil.

Concentrations of S (and Ca) in the soil and spoils treated with LIMB by-product,
where enttringite did not form, appeared to be controlled by the solubility of gypsum
(CaSO^ ¦ 2H2O). Because enttringite is much less soluble than gypsum, S
concentrations decreased in leachates where this mineral formed, even though the
amount of LIMB by-product added was greater. With PFBC by-product, leachate
concentrations of S (and Mg) would be controlled by epsomite (MgS04 • H2O). The
concentrations of S would, therefore, be greater where PFBC is applied because
epsomite is approximately 300 times more soluble than gypsum. The
concentrations of Ca, also increased in the PFBC as compared to the LIMB by-product
treated soil and spoil, due to the solubility of caldta The results in Figure 4 indicate
that the presence of Mg in the PFBC by-product gives this by-product a greater
potential for excessive salt loading than the LIMB by-product.

Concentrations of Regulated Elements in the Leachate

Concentrations of elements regulated with respect to land application of sewage
sludge are shown in Table 2. These regulatory levels for sewage sludge were chosen
as a basis for evaluating our leachate results because we had included in our
experimental design the addition of 6% sewage sludge.

Not listed in Table 2 are Hg and Fb. Mercury was below the detection limit (<0.04
mg/kg) in all leachates ana also was not detected (<0.0002 mg/kg) in Toxicity
Characteristic Leaching Procedure (TCLP) extracts of LIMB or PFBC by-products.

Lead also was not detected (<0.04 mg/kg) and the TCLP levels for Fb in the LIMB and
PFBC by-products were 0.005 and <0.001mg/kg, respectively. Cadmium
concentrations in the leachate showed either no effect or a decrease in concentration
with increasing rates of FGD application (Table 2). Thus, with respect to these
metals, there appears to be little potential for environmental contamination as a
result of FGD by-product used as a soil or spoil amendment.

Concentrations of Cu, Ni and Zn generally decreased with increasing rate of FGD
application. These results are consistent with most studies of trace metal behavior
in soils, which show that metal solubility and mobility decreases as pH increases
above the acid range to neutralitylU2. Copper, Ni and Zn were much higher in the
sewage sludge than in FGD by-products (Table 1). Since the same amount of sewage
sludge was applied to all pots, regardless of FGD rate applied, the amount of these
elements added to the pots would be approximately the same. However, at very
high application rates of FGD by-product (24% by weight), concentrations of these
metals were found to be increased in the leachate. This is due to high pH
solubilizing the organic matter in the sewage sludge. Soluble organic matter can
complex with metals bringing them into solution' and these complexes are then
leached from the soil or spoil at elevated levels.

Leachate concentrations of the oxyanion, B, tended to increase with increasing FGD
by-product application rate (Table 2). At the highest FGD application rate applied to

104-6

-------
the Fleming AML spoil (24%), however, the concentration of B decreased (data not
shown) due to formation of enttringfte in the pots. Concentrations of the other
oxyanions (As, Mo, and Se) seemed to be unaffected by FGD application (Table 2) at
rates of 12% or lower.

Both the FGD by-products and the sewage sludge materials contained As, B, Mo, and
Se (Table 1). Increased concentrations observed at the highest application rate (24%)
(data not shown) is due to both the increasing amounts of oxyanions added to the
pots as well as increased competition of sulfate and organic Mgands for anion
absorption sites in the soil or spoil. This competition would force some of the
oxyarrions, that would normally be retained in the soil or spoil, to enter the solution
phase and be leached from the pots. In contrast, enttrmgjte formation would limit
contact between the leachate water and As, B, Mo and Se in the pots, thereby
decreasing their concentrations in the leachate The increases in B concentrations
were much greater in leachates from spoil treated with LIMB by-product than the
FFBC by-product. This is not surprising since total B in the LIMB by-product was
much higher than in the PFBC by-product Most species of B that are phytotoxic to
plants are water soluble and will rapidly leach from the root zone. However,
growth of plants may be initially inhibited by B when high rates of some FGD by-
products are applied

Plant Growth Responses to FGD By-Product Applications

Plant (alfalfa) growth was increased in the acid Wooster soil when treated with
LIMB by-product or the bed and cyclone PFBC by-product (Figure 5). In the first
harvest the optimum growth generally occurred at application rates of from 0.35 to
1.4%. At the highest rates, equivalent to two times the lime requirement for this
soil, growth during the first harvest period was inhibited. This is attributed to
excessively high pH values that initially developed in the soil when these high rates
were applied. However, by the third harvest the inhibitory effect was no longer
evident and, in fact, the 2.8% application rate yielded significantly more alfalfa than
did the untreated control. By this time, carbonation had occurred and decreased the
soil pH and the inherent buffering capacity of the soil had time to react with the
excess alkalinity. Our results (Figure 5) indicate that dry FGD by-products may be
used as agricultural lime substitutes, provided the rates applied are approximately
equivalent to the lime requirement of the soil If higher rates are applied, then fall
applications would be recommended to provide sufficient time for chemical
reactions to occur in the soil, during the winter months, prior to seeding of crops.

Similar results were observed when the acid spoil from the Fleming AML site was
treated with LIMB and PFBC (cyclone) by-products (Figure 6). This experiment used
much higher rates, however, and the high pH inhibited fescue growth until the
fourth or fifth harvest. The growth of fescue on spoil treated with 12% PFBC by-
product during the first harvest period seemed poor (Figure 6), but this was due to
poor growth in one of the four replications. However, subsequent growth periods
showed good growth response to applications of the FGD by-products. On the basis
of our greenhouse experiments, we would recommend application rates of FGD by-
products onto acid spoil of approximately 6 to 12% by weight. This rate supplies
sufficient alkalinity to neutralize the acidity in the spoil but still avoids inhibition of
plant growth because of excess soluble salts or high pH.

The field columns at Coshocton compared alfalfa and fescue growth, under natural
field conditions, in the add Fleming AML spoil treated with FGD by-products that
differed in the sulfur forms. The LIMB by-product primarily contains gypsum

104-7

-------
(CaS04) while the spray dryer FGD by-product contained over half of its total sulfur
as calcium sulfite (CaSC^jl, Calcium sulfite is considered inhibitory to plant growth
and this experiment was designed to evaluate plant growth response to the two FGD
by-products. Optimum growth (15.2 g/ column) after 125 days occurred in the spoil
treated with 10% spray dryer FGD by-products (Table 3). Less growth (8.4 g/column)
was observed when LIMB by-product was applied at the 10% rate Apparently the
pH from the LIMB by-product treatment was too high for growth, even for fescue,
which generally can withstand more extreme conditions than alfalfa. As the
application rates increased, growth decreased few both the LIMB and the spray FGD
diyer by-products. The results reported are only from the first harvest, but if the
growth pattern in the field fallows that observed in the greenhouse, subsequent
harvests, in the columns treated at the 20% FGD by-product application rate, would
be improved over that observed after the first 125 days. Also the rates applied here
exceeded that in the greenhouse. Our best growth responses on the treated Fleming
AML spoil in the greenhouse occurred at the 6 to 12% rate, which is consistent with
what was observed in the field where optimum growth occurred at the 10% rate.
The CaS03 in the spray dryer FGD by-product did not seem to result in any greater
growth inhibitions than that observed for the LIMB by-product (i.e. gypsum
containing by-product). The pH of the spray dryer FGD treated columns was not as
high as where LIMB by-product was applied and CaS03 is much less soluble than
gypsum also minimizing its effect on plant growth.

Conclusions

FGD by-products appear to be highly effective as alkaline amendments for acid soil
and spoil Addition of FGD by-product effectively increased pH to near neutrality or
to slightly alkaline pH and decreased concentrations of soluble A1 and Mn. The
result of these chemical changes is an improved growing medium for plants.
Improved plant growth was, indeed, observed when FGD by-product was mixed
with acid soil or spoil. The addition of 6% sewage sludge with the FGD by-product
created the best conditions for plant growth.

Potential adverse effects on plant growth of high concentrations of soluble salts,
excessively high pH, and trace metal toxicities were not observed, even at rates
slightly in excess required to neutralize the soil or spoil.

Leachate composition indicated that at application rates of 12% or lower,
concentrations of elements of environmental and regulatory concern remained
very low. Most, in fact, were below drinking water standard levels. Boron was an
exception when LIMB by-product was used to amend the spoil. The limiting factor
for application rates of Mg-containing FGD by-products, such as the FFBC by-
product, is more likely to be high soluble salt concentrations, which may inhibit
growth and impact water quality. For the LIMB by-product the factors limiting use
are the initially high pH values that were observed and the potential for enttrmgite
formation. Both would inhibit plant growth. However, it must be stressed that
these limitations occurred only at very high application rates.

There seems to be little potential for adverse effect on water, soil and plant quality
when FGD application rates, based on the amount required to neutralize soil or
spoil acidity, are not exceeded. This last statement is made in the context that the
neutralization potential of the FGD by-products is approximately equivalent to 50%
calcium carbonate FGD by-products with lower neutralization potentials would

104-8

-------
require higher application rates and, thus, their environmental impact would differ
from what is reported here.

Acknowledgments

This research was conducted as part of the "Land Application Uses for Dry FGD By-
products" project which is a cooperative project of the Ohio Agricultural Research
and Development Center, The Ohio State University, The U.S. Geological Survey,
and the Dravo Lime Company. Funding support for this project was obtained from
the Ohio Coal Development Office (Columbus, OH) Grant No. CDO/D-89-35, The
U.S. Department of Energy (Morgantown Energy Technology Center, Morgantown,
WV) Award No. DE-FC21-91MC28060, Dravo lime Company (Pittsburgh, PA) Grant
No. RF768342, Electric Power Research Institute (Palo Alto, CA) Grant No. RP2796-
02, American Electric Power Company (Columbus, OH) Grant No. C-8276, Ohio
Edison Company (Akron, OH), and the Ohio State University (Columbus and
Wooster, OH).

References

1.	Land Application Uses for Dry FGD By-Products, Phase 1 Columbus, OH:

Department of Agronomy, The Ohio State University, April 1993.

2.	C. L Carlson and D. C Adriano. "Environmental Impacts of Coal Combustion

Residues." Journal of Environmental Quality. Vol 22, No. 2, p. 227 (1993).

3.	R. K. Fowler, J. M. Bigham, S. Traina, U. I. Soto, R. G Stehouwer, and E. L.
McCoy. "Properties of Qean Coal Technology By-Products." Agronomy
Abstracts. Madison, WI: American Society of Agronomy, 1993, p. 361.

4.	R. F. Korcak "Fluidized Bed Material as a Lime Substitute for Liming Leached

Mineral Soils." Journal of Environmental Quality. Vol. 9, No. 1, p. 147 (1980).

5.	G. L Term an, V. J. Kilmer, C M. Hunt, and W. Buchanan. "Fluidized Bed Boiler
Waste as a Source of Nutrients and lime." Journal of Environmental Quality.
Vol 7, No. 1, p. 147 (1978).

6.	H. L. Bohn, B. L. McNeal, and B. A O'Connor. Soil Chemistry, 2nd Edition
New York, NY: John Wiley and Sons, 1985.

7.	F. J. Stevenson. Humus Chemistry; Genesis, Composition, Reactions New
York, NY: John Wiley and Sons, 1982,

8.	N. V. Hue, B. R. Craddock, and F. Adams. "Effect of Organic Acids on

Aluminum Toxicity in Subsoils." Soil Science Society of America Journal. Vol
50, No. 1, p. 28 (1986).

9.	E J. Kamprath. "Exchangeable Aluminum as a Criterion for Liming Leached
Mineral Soils." Soil Science Society of America Journal. Vol. 34, No ?, p. 252
(1970).

10.	P. Sutton and W. A Dick. "Reclamation of Acidic Mined Lands in Humid

Areas." Advances in Agronomy. Vol 41, p. 377 (1987).

104-9

-------
11.	P. B. Woodbury. "Trace Elements in Municipal Solid Waste Composts: A

review of Potential Detrimental Effects on Plants, Soil Biota, and Water Quality."
Biomass and Bioenergy. Vol. 3, p. 239 (1992).

12.	R L Chaney, R J. F. Bruins, D. E. Baker, R F Korcak, J. E. Smith, and D. W. Cole.

"Transfer of Sludge-Applied Trace Elements to the Food-Chain," Land
Application of Sludge: Food Chain Implications. Chelsea, Michigan: Lewis
Publishers. 1987, pp. 67-99.

Table 1

Characterization of LIMB, PFBC(cycI), and Sewage Sludge Amendments.

Parameter

LIMB

PFBC(cycl)

Sewage Sludge

Particle size (%)





na*

Sand (0.05-2mm)

0

25.5



Silt (2-50pm)

90

74.1



Clay (<2pm)

10

0.4



Mineralogy (%)





na

Anhydrite (CaS04)

25

22



Calcite (CaCOs)

15

11



Dolomite (CaMg(C03)j)

nd2

23



Lime (CaO)

21

nd



Portlandite (Ca(OH)2)

5

nd



Periclase (MgO)

nd

13



Fly ash

30

32



CaCO, equivalent (%)

59.4

60.3

0

pH (1:1, water)

12,5

10.5

6.5

Total Chemical Analysis







Organic C (%)

na

na

31.2

Ca (%)

35.96

17.53

2.76

Mg (%)

0.60

10.64

0.34

S (%)

5.77

5.21

1.41

A1 (%)

3.52

3.93

3.40

Si (%)

6.58

7.24

na

Fe (%)

5.56

5.17

1.24

Na (%)

0.33

1.03

0.07

N (%)

na

na

3.8

P (%)

0.02

0.02

1.78

K (%)

0.91

0.50

0.15

Ba (%)

0.03

0.02

0.01

As (mg kg"1)

55.1

1.9

<0.03

B (mg kg*1)

233.1

171.2

31.1

Cd (mg kg1)

1.0

1.9

6.3

Cr (mg kg"')

28.0

36.9

315.2

Cxi (mg kg"1)

21.0

52.5

1174

Pb (mg kg1)

16.0

16.0

16.1

y-s

I
1

5.9

6.6

11.2

Ni (mg kg"1)

31.1

52.1

166.4

Se (mg kg"1)

8.1

5.6

<0.3

Zn (mg kg"')

86.0

74.0

1494

"not analyzed

104-10

-------
Table 2

Leachate Concentrations of Elements Regulated With Respect to Land Application of Sewage Sludge,*

Treatment

As

B

Cd

Cr

Cu

Mo

Ni

Se

Zn































lllfti jLj









Soil alone

<0.08

0.13

0.021

<0.004

<0.007

<0.018

0.06

<0.27

0.645

Soil + IX lime requirement (PFBC)

<0.08

0.80

0.019

<0.004

0.071

<0.018

0.04

<0.27

0.012

Soil + 2X lime requirement (PFBC)

<0.08

1.54

0.012

<0.004

0.064

0.024

0.03

<0.27

0.032

Spoil + sewage sludge

<0.08

0.41

<0.003

0.009

0.041

<0.018

0.24

<0.27

3.34

Spoil + sewage sludge + 12% PFBC

<0.08

0.92

<0.003

0.021

0.055

0.089

0.01

<0.27

0.06

Spoil + sewage sludge + 12% LIMB

<0.08

3.73

<0.003

0.008

0.026

<0.018

0.01

<0.27

0.05

' Not listed are Hg and Pb which were below detection limits (0.04 mg/L) in all cases.

-------
Tables

Dry Matter Yields of Alfalfa And Fescue 125 Days After Seeding When Grown on Fleming AML
Spoil Treated With Different Rates of LIMB and Spray Dryer FGD By-Products.

Dry Matter Production (a/column)

Treatment

Alfalfa

Fescue

in j- t

Total

Control

0

0

0

10% UMB

6.8

1.6

8.4

20% LIMB

0.04

0.5

0.54

40% UMB

0

0.1

0.1

10% Spray Dryer

14.2

1.0

15.2

20% Spray Diyer

2.8

0.09

2.89

40% Spray Dryer

0

0.01

0.01

LSD (0.05)

3.9

3.2

3.6

0.0 0.7 1.4	2.8

RATE (weight %)

Figure t, pH, aluminum and manganese leachate concentrations torn Wooster soil
treated with PFBC by-product The first leaching was conducted after _
seeding of the pots to alfalfa or fescue and the last leaching was after nine
months of growth.

104-12

-------
RATE (weight %)

Figure 2. pH, aluminum and manganese leachate concentrations from Fleming

AML spoil treated with PFBC by-product The first leaching was conducted
after seeding of the pots to alfalfa or fescue and the last leaching was after
nine months of growth.

104-13

-------
260

Cft

c

CD

40

°2G CuKgt

Figure 3, Representative X-ray diffraction pattern for Fleming AML spoil treated with 24% PFBC by-product.

-------
0

1

tn

0 3 6	12	24

RATE (weight %)

o

V

PFBC first leaching
LIMB first leaching

PFBC last leaching
LIMB last leaching

12

O

Q.

12

cn

a

LlJ
>-

<
U_

<

12

PFBC(cycl).

PFBC(bed)

o First Harvest
• Third Harvest
v Sixth Harvest

Figure 4 Suite concentrations in leaehates from Wooster soil and Fleming AML Figure 5.
spoil treated with either LIMB or PFBC by-product. The first leaching was
conducted after seeding of the pots to alfalfa or fescue and the last leaching
was after nine months of growth.

0.0 0.7 1.4	2.8

RATE (weight %)

Alfalfa yields (first, third and sixth harvests) from pots containing
Wooster soil treated with LIMB and PFBC (source, bed and cyclone)
by-products.

-------
o

CL

cn

Q

L±J

LlI
3
O
C/5
LlJ

10

6
4
2
10

i	r

LIMB

4 -

0

o First Harvest
• Third Harvest
v Sixth Harvest

+

PFBC(cycI)

12

24

RATE (weight %)

Figure 6. Fescue yields (first, third and sixth harvests) from pots containing Fleming
AML spoil treated with LIMB and PFBC by-products.

104-16

-------
The Stabilization of Orimulsion
Spray Dryer Waste for Landfill Disposal

S. Kuchibotla
Law Engineering, Inc.
Jacksonville, Florida, 32207

EH. Kalajian
Honda Institute of Technology
Department of Civil Engineering
Melbourne, Florida 32901

C-S. Shi eh
Florida Institute of Technology
Department of Environmental Science
Melbourne, Florida 32901

K.R. Olen
Florida Power & Light Co,
Engineering & Advanced Technology
Juno Beach, Florida 33408

Abstract

In this paper the results of a study to develop a fixation process to stabilize
Orimulsion lime spray dryer (LSD) waste are presented. Orimulsion is a candidate
alternative fuel for utility boilers because of its low cost and outstanding handling
and combustion characteristics. An obstacle to commercializing Orimulsion is the
need to install backend environmental control systems to reduce particulate and SOz
emissions. The objective of the study was to determine physical and chemical
properties of the raw Orimulsion LSD waste and its stabilized products and to
determine if the stabilized product is suitable for disposal in landfill. Hie results of
engineering studies indicated that the Orimulsion LSD waste had physical properties
similar to clay soil with low compressibility; however, the Orimulsion LSD waste has
a very low bulk density. The compacted raw Orimulsion LSD waste was non-
pozzolanic and is not recommended for landfill disposal. The stabilized Orimulsion
LSD waste was found to be acceptable for landfill disposal based on its compacted
dry density and compressive strength. Surface hardness of the stabilized Orimulsion
LSD waste product was found to be acceptable for transportation and handling. The
results of chemical studies have shown that stabilization of Orimulsion LSD waste
using Type I Portland cement minimized the leaching of Ca, V, and SO/" to
surrounding aqueous solutions. It is concluded that, using the mix design developed
in the study, Orimulsion LSD waste can be stabilized and disposed of in landfill.

105-1

-------
Intro due tion

Florida Power & Light Co. (FPL) has been, evaluating the possibility of converting one
or more of its utility boilers from residual fuel oil-firing to Orimulsion Orimulsion is
a new fuel form consisting of small amorphous particles of a naturally occurring
bitumen emulsified into water. It exhibits non-Newtonian pseudoplastic flow
properties and physically resembles black latex paint. The bitumen used to formulate
Orimulsion is mined from large reserves in Venezuela's Orinoco River basin.
Currently, Bitumenes Orinoco (BITOR), a subsidiary to the Venezuelan state owned
oil company, Petroleos de Venezuela, S.A., is marketing Orimulsion as an alternative
utility boiler fuel at coal equivalent prices.

Orimulsion resembles a high-sulphur Eastern bituminous coal, in terms of heat and
fuel sulphur contents, 12,700 Btu/lb and 2.90%, respectively. However, unlike the
experience of firing a boiler on coal, Orimulsion produces very little fly ash, e.g.,
about 0.10 gr/DSCF. In addition, Orimulsion fly ash is dominated by submicron
particles of non-pozzolanic magnesium and vanadium oxysulphates. In this regard
Orimulsion resembles a high-vanadium, high-sulphur Venezuelan residual oil, that
had been treated with a magnesium-based fuel additive.

A conversion to Orimulsion would require the installation of environmental control
systems to reduce particulate and S02 emissions. Based on the experience of
reducing S02 flue gas emissions from burning high-sulphur coal, the flue gas
desulphurization (FGD) system of choice would be a wet limestone forced-oxidation
(WLFO) scrubber. Recent information, however, strongly suggested that the life-cycle
costs of a lime spray dryer (LSD) may be competitive against WLFO systems for 90+%
FGD of flue gases with high S02 concentrations (1,2). The major variables to be
compared when evaluating these two competing processes are capital requirements,
the relative costs of limestone and lime, sorbent utilizations, and the costs associated
with solid waste disposal.

With regard to the latter of these variables, WLFO scrubbers produce gypsum, which
can be used for wallboard manufacture, soil conditioning, or safely stacked in an
approved landfill without further processing. The solid waste from a LSD, in
comparison, has no commercial value, as it is primarily calcium sulphite, and must,
therefore, be stabilized and landfilled. Solid waste stabilization is effectively and
inexpensively accomplished in the case of a coal fired plant by mixing the pozzolanic
coal fly ash with the FGD product, and water. The mixture is typically placed in an
approved landfill, where it is compacted and cured thereby forming an
environmentally acceptable waste.

Given the chemical characteristics of Orimulsion fly ash, an economical solid waste
stabilization process had to be developed, before LSD technology could be fully
evaluated. Since Orimulsion and residual oil fly ashes are very similar, it was
believed that the oil ash stabilization process developed at Florida Institute of
Technology (Florida Tech) under FPL sponsorship (3) might provide a basis for a

105-2

-------
process to stabilize Orimulsion LSD solid waste. Accordingly, FPL contracted
investigators at Florida Tech to carry out a study to determine the process
requirements to stabilize Orimulsion LSD solid waste.

The unconfined compressive strengths recommended for landfill disposal of ash
wastes are 100-150 psi (4,5, 6). Thompson et al (7) stated that development of
compressive strengths by spray dryer wastes in excess of 25 psi is considered
adequate to support the weight of the material when used in a fill, as well as the
weight of the trucks and other equipment used for placement and compaction. Spray
dryer wastes with compressive strengths in excess of IS) psi can be expected to
perform acceptably as structural fill material (4).

Preparation of Simulated Lime Spray Dryer Solid Waste

Orimulsion LSD solid waste samples were prepared by Radian Corp., Austin, Texas
using a laboratory-scale spray dryer. To simulate commercial operations a synthetic
flue gas was generated by burning natural gas at high excess air levels such that the
adiabatic flame temperature was essentially the desired spray dryer inlet temperature
of about 325°F. Gaseous SOz was metered into the synthetic flue gas to maintain a
concentration between 2450 to 2500 ppm(v). Feed slurry consisted of slaked lime,
recycled solids from a previous run carried out under similar conditions, and a
quantity of Orimulsion fly ash. The rate of addition of fly ash was adjusted to
approximate a flue gas concentration of about 0.10 gr/DSCF. Inlet gas velocity was
approximately 275 acfm, and the spray dryer was operated at a 22°F approach
temperature.

Despite the unorthodox technique used to introduce the Orimulsion fly ash, it was
believed that the solid wastes produced from the laboratory-scale spray dryer were
reasonably representative of those that would be produced in a full-scale commercial
operation. Accordingly, about 25 kg of waste sample was collected for each of two
desulphurization runs. The aim desulphurization levels were 72.5 and 90%, however,
because the laboratory spray dryer is normally operated at a constant slurry feed rate
to achieve a constant calcium-to-sulphur molar ratio it was difficult to control the test
unit precisely to a predetermined S02 removal level. After adjusting for air in
leakage and moisture, the corrected FGD percentages for the two samples were 64
and 85%.

Calcium utilization for the low and high desulphurization runs were 100 and 94%,
respectively. These utilizations are somewhat higher than those that might be
expected in a full-scale commercial unit because of scale down factors that make the
laboratory-scale unit more efficient The higher efficiency was also demonstrated by
uncharacteristically low calcium-to-sulphur molar ratios of 0.64 for the 64% FGD run
and 0.90 for the 85% FGD case.

105-3

-------
Methodologies

Characterization of Orimulsion LSD Waste. Methods used to characterize the raw
and stabilized Orimulsion LSD wastes are listed in Table 1. The physical properties
determined were moisture content, grain size distribution, specific gravity, Atterberg
limits, bulk density and moisture-density relationships, while the engineering
properties determined were unconfined compressive strength and surface hardness.
Moisture content was determined using procedures outlined by American Society of
Testing Materials (ASTM) C 566-78 (8) 'Total Moisture content of Aggregates by
Drying". Grain size distribution was determined using procedures outlined by ASTM
D 422-63 (8) "Standard Method for Particle Size Analysis of Soils". Specific gravity
was determined using procedures outlined by ASTM D 854-83 (8) "Standard Method
of Test for Specific Gravity of Soils." The Atterberg limits were determined using
procedures outlined by ASTM D 4318-84 (8) 'Standard Test Methods for Liquid limit,
Plastic Limit and Plasticity Index of Soils." The bulk density was determined using
procedures outlined by ASTM D 698-78 (8) and ASTM 1557-78 (8). The moisture-
density relationships was determined using procedures outlined by ASTM D 558-82
(8) "Moisture-Density Relationships of Soil-Cement Mixes." The unconfined
compressive strength tests were conducted in accordance with ASTM D 2166-85 (8)
'Test Method for Unconfined Compressive Strength for Cohesive Soil." The
American Foundrymen's Association method for evaluating the surface hardness of
the baked and cured cores was used to determine the hardness of the samples (9).

The chemistry study was carried out to determine the concentrations of Ca, V, and
S042" in both raw and stabilized Orimulsion LSD wastes. Dried and ground samples
were digested using the hydrofluoric-boric acid (HF-H3B03) technique (10). The
digest was analyzed for Ca and V using the Perkin-Elmer Model 5100 PC Zeeman
atomic absorption spectrophotometer (AAS) equipped with HGA-600 graphite
atomizer and AS-60 autosampler. Sulfate was analyzed using the Shimadzu UV-VIS
Recording Spectrophotometer UV-160A according to the method described in
Standard Methods for the Examination of Water and Wastewater (11).

Stabilization and Fabrication. In the stabilization methodology, Type I cement (10%
or 15%) and sodium metasiHcate commonly known as Metsobeads (0.75%, 1%, 1.25%,
or 1.5%) were used to stabilize both 64% and 85% LSD wastes. It has been
considered that use of Metsobeads will improve the engineering characteristics of
cement products. The moisture contents selected were +3, +5, +10, +20, +35 and +50
percent above Optimum Moisture Content (OMC). The OMC is defined as the
moisture content at which the soil can be compacted to the maximum dry density by
using given compaction (12). To obtain a relationship between the moisture content
and the unconfined compressive strengths of the LSD wastes, samples were
fabricated at various moisture contents relative to the OMC. Based on the results
obtained, the moisture content at the maximum strength could be determined.

The test samples were fabricated using the Harvard Miniature Apparatus (2.81 inches
in height and 1.31 inches inner diameter). A forty pound spring was used in the

105-4

-------
tamper for compaction and twenty-five tamps per layer were applied to each of the
five layers, which approximates Standard Proctor compaction. After compaction, the
samples were wrapped in plastic wrap and placed in plastic boxes. These boxes
(with their top open) were placed in the curing chamber (100% humidity). The air
dry samples were left in the sealed plastic boxes on a laboratory shelf at room
temperature.

Determination of Leaching Characteristics. Leaching characteristics of the selected

stabilized Qrimulsion LSD waste was determined by conducting both column and
tank leaching studies. The column leaching test was conducted according to the
method by Harder et al. (13), which simulated the percolation of natural waters
through a stabilized product The tank leaching test was based on the work by
Duedall et al. (14), van der Sloot et al. (15), an van der Sloot and De Groot (16),
which simulated the immersion of a stabilized product in a aqueous solution.

Results and Discussion

Physical Properties. Both the 64% and 85% Grimulsion LSD wastes had low
moisture contents ranging from 0.1% to 0.5% (Table 2). The mean grain size,
uniformity coefficient and coefficient of gradation as reported in Table 2 showed little
variation between the two Orimulsion LSD wastes. Therefore, the material can be
characterized as uniform and in the silt-clay size range. The specific gravity for the
64% and 85% LSD wastes were found to be 2.50 and 252, respectively (Table 2). The
Atterberg limits for the 64% LSD waste were as follows: liquid limit 35%, plastic limit
10%, and plasticity index 25% (Table 2). For the 85% LSD waste the liquid limit was
36%, plastic limit was 10%, and plasticity index was 26%. The results indicated that
both 64% and 85% LSD wastes had similar plasticity characteristics. The aerated and
settled bulk density for the 64% LSD waste were 27 pcf and 35 pcf, respectively
(Table 2). The aerated and settled bulk density for the 85% LSD waste were 30 pcf
and 35 pcf, respectively. Hie void ratios determined for both the wastes were in the
range 3.5-4.8 and were high compared to the void ratio of typical clay soil which
ranges between 1 to 2. The uniform grain size distribution of the Orimulsion LSD
waste results in poor packing and compaction of the wastes, thus providing low bulk
density and high void ratio.

Based on the Atterberg limits and grain size distribution the 64% and 85%

Orimulsion LSD wastes were classified as CL soil (clay of low compressibility) in the
Unified Soil Classification System according to ASTM D 2487-85 (8) "The Method for
Classification of Soils for Engineering Purposes."

105-5

-------
Table 1

Methods Used to Characterize Orimulsion LSD Waste.

Parameter

Method

Reference

moisture content

ASTM 566-78

8

grain size distribution

ASTM D 422-63

8

specific gravity

ASTM D 854-83

8

Atterberg Ixmits

A^STM D 4=310 — 84

8

bulk density

ASTM D 698-78

8



ASTM 1557-78

8

moisture-density relationships

ASTM D 558-82

8

unconfmed compressivs strength.

ASTM D 2166-85

8

Classification

ASTM D 2487-85

8

surface hardness

The American





y t it 1Y1 ^ ' S





Association Method 9

acid digestion

hf-h3bo3

10

leach.2-Tig* ch3ac t e^u xstxc

c^?lumLk leachxng

13



tank leaching

14, 15, 16

Table 2

Physical Properties of Raw Orimulsion LSD Waste.

Parameter

64% LSD waste

85% LSD

waste





moisture content (%5

0.1-0.5

0.2-0.5

mean particle size (Dso , mm)

0.013

0.011

uniformity coefficient

1.263

1.250

coefficient of gradation

0.88

0.90

specific gravity

2.50

2.52

Atterberg Limits





liquid limit (%)

35

36

plastic limit {%)

10

10

plasticity index (%)

25

26

bulk density {aerated, lb/ft3 }

27

30

g/cm3

0.433

0.480

bulk density (settled, lb/ft3 5

35

35

g/cm3 5

0.561

0.561

void ratio (aerated)

4.8

4.2

void ratio (settled)

3.5

3.5

105-6

-------
Chemical Properties. The concentrations of Ca, V, and S042" in raw Orimulsion LSD
waste are shown in Table 3. The concentrations of Ca in 64% LSD waste and 85%
LSD waste were approximately 25%. The 85% LSD waste showed a slightly lower V
concentration than the 64% LSD waste (0.19% and 0.22%, respectively). Hie sulfate
concentration in 85% LSD waste and 64% LSD waste were similar (17.3% and 16.7%,
respectively).

Enrichment of Ca in Orunulsion LSD waste resulted mainly from the scrubbing
process. In the scrubbing process, a suspension of Ca(OH)2 or CaC03 was used to
remove S02 from the flue gas resulted in the formation of calcium sulfate
hemihydrate (CaS04 -44H20) which was subsequently oxidized to yield calcium
sulfate dihydrate (CaS04 -2H20), i.e., flue gas gypsum (17). Major Ca compounds in
scrubber residues identified by X-ray diffraction are caltite, gypsum, and
hannebachite (CaS04 ~W20) (18,19).

The V concentration measured in the Orimulsion LSD waste was over 10 times
higher than the approximate 71 to 180 jig g"1 concentration found in coal scrubber
by-products (20). The V in the ash may have had an effect on the V concentration in
the scrubber residue because the composition of scrubber residues is, in general,
dependent partly on ash generation (17). Vanadium as V205 is found in flue dust
after fossil fuel combustion (21), which indicates a possible source of V in the
Orimulsion LSD waste.

In the LSD waste, the S042" was mainly derived from the sulfur dioxide removed in
the flue gases. In the scrubbing process, S02 is absorbed to form ultimately
CaS04 -%H20, CaS04 -2HzO or a combination of the two (17). Therefore, the S042"
measured from the acid digest of the Orimulsion LSD waste was correlated to the
amount of SO, removed from the flue gases.

Compaction and Strength of Orimulsion LSD Wastes. Mix designs for stabilizing
Orimulsion LSD wastes, the moisture contents evaluated and the 28-day compressive
strength are shown in Table 4. Factors examined included the content of cement,
water, metsobeads (sodium metasilicate) additives, and curing time. The unconfined
compressive strengths were determined at different compacted moisture contents to
obtain a relationship between the unconfined compressive strength and the moisture
content of the Orimulsion LSD wastes arid to determine the ideal mixing moisture
content. The 28-day compressive strengths for the 10 percent cement stabilized 64%
and 85% Orimulsion LSD waste samples increased between OMC and +3 percent
above OMC. As mentioned above, the OMC is defined as the moisture content at
which the soil can be compacted to the maximum dry density by using given
compaction. This indicates that sufficient water was available for the hydration of
cement to complete between OMC arid +3 percent above OMC. The unconfined
compressive strengths for both the 15 percent cement stabilized 64% and 85%
Orimulsion LSD waste samples exhibited similar properties and gained maximum

105-7

-------
Table 3

Concentration of Ca, V, and S04*~ in Orimulsion LSD Waste.

Elements	64% LSD Waste	85% LSD Waste

Ca (%)	25.1±0.2	25.7±0.2

V	(pg g"1)	2182*7	192Q±6

SO/" (%}	16.7±1.5	17„3±0.6

Standard deviation, ti— 3 .

Table 4

Orimulsion LSD waste Mix Designs Investigated in the Study.

LSD waste cement metsobeads moisture dry	28-d strength.

(%)	{%)	(%}	content density	(psi)b

{%)	(Ibs/cf)a

64	0	0	20	50.7	72

64	10	0	24-74	43.4-55.3	94-172

64	10	0.75-1.5	27	55.4	160-168

64	15	0	26-31	55.0-56.4	200-225

85	0	0	21	50.3	88

85	10	0	25-75	43.1-55.4	71-134

85	10	0.75-1.5	28	54.8-55.0	121-133

85	15	0	29.5-34.5 55.2	155-274

a 62.4 Ibs/cf = 1 g/cm3
b 1 psi = 6.89 kPa

105-8

-------
strengths at 3 percent above their corresponding OMC's. The results indicated that
the moisture content at mixing plays an important role in determination of the
engineering properties. The moisture content at which all the cement stabilized
waste samples obtained their maximum unconfined compressive strength was +3
percent above their corresponding OMC.

The unconfined compressive strengths of the 10 percent cement stabilized 64%
Orimulsion LSD waste samples were higher than the unconfined compressive
strengths obtained for the 10 percent cement stabilized 85 percent Orimulsion LSD
waste samples. As mentioned by Rodriguez et al. (4) the presence of sulfates (in
solid state) in soils decreases the unconfined compressive strengths obtained for the
soil-cement mixes. The presence of higher sulfate content in the 85% Orimulsion LSD
wastes explains the lower compressive strengths for the 10 percent cement stabilized
85% Orimulsion LSD waste samples. Rodriguez et al. (4) also indicated in their
research, the cement content plays an important role in improving the engineering
properties of the stabilized soils. It was found in the study that the cement improved
the unconfined compressive strength of the 15 percent cement stabilized 85% waste
samples, thus providing them higher strengths than the 15 percent cement stabilized
64% waste samples. Increase the content of cement in the mix overrided the effect of
sulfate on stabilization. Metsobeads had been found no effect on improving the
unconfined compressive strengths of the 10 percent cement stabilized Orimulsion
LSD wastes.

The effects of curing condition on the various mix designs were found to have very
little effect on the compressive strength. Samples cured in air had slightly lower
compressive strength (< 10%) than samples cured in a chamber. The compressive
strength of stabilized Orimulsion LSD waste increased with time, however, the
stabilized waste has greater than 60% of its compressive strength developed after 7
days as compared to compressive strengths at 28 days.

The surface hardness of the cement stabilized 64% arid 85% Orimulsion LSD wastes
were found to be in the range of 75 and 85, which classifies them as very hard. The
stabilized mixes can be transported mechanically, with minimum breakage.

Characteristics of Stabilized Orimtils ion LSD Waste. The optimum moisture
content, the maximum dry density and void ratios for the stabilized wastes are listed
in Table 5. As seen in the table there is no significant difference between the 10
percent cement stabilized 64% and 85% Orimulsion LSD wastes and the 15 percent
cement stabilized 64% and 85% Orimulsion LSD wastes. The dry densities for the
stabilized wastes increased approximately 4-6 pcf over the uristabilized wastes and
the OMC for the stabilized wastes increased approximately 4-9 percent over the
uristabilized wastes.

The void ratios for the stabilized 64% and 85% Orimulsion LSD wastes were in the
range of 1.8-1.9. The void ratios for the cement stabilized Orimulsion LSD wastes

were found to have reduced by 85 percent from the void ratios obtained for the raw

105-9

-------
wastes and 10 percent lower than the void ratios obtained lor the unstabilized
Orimulsion LSD waste. This indicates that the void ratio decreased with the addition
of cement, improving the dry density. Dry densities of the Orimulsion LSD waste
increased 66% when compacted at OMC as compared to the settled bulk density
shown in Table 2.

Table 5 also shows the concentrations of Ca, V, and S042' in stabilized Orimulsion
LSD products. Both mixes, e.g., 85% LSD-10% cement (mix 85%-10) and 85% LSD-
15% cement (mix 85%-15), had similar Ca, ¥, and S042" concentrations. The Ca
concentration was approximately 7% higher in the stabilized LSD products than in
the raw LSD waste (85% LSD). The concentration of V in mix 85%-lQ and mix 85%-
15 were 1310 p.g g'1 and 1270 jig g~\ respectively, which were about 600 |ig g'1
decrease from the raw LSD waste. The sulfate concentration, 18.0% in both mixes,
did not change significantly in the stabilized LSD products compared to 17.3% in the
raw LSD waste.

The addition of Portland cement with Ca additives (lime, limestone, and calcium
silicates) contributed to the high concentration of Ca detected in the mixes. The
primary Ca compounds in the stabilized products may be the products of the
hydration of cement, such as calcium hydroxide, calcium sulfate,
and calcium silicates (22), or scrubbing by-products such as CaCO^ CaS04 -2H20, and
CaS04 J/iH,0 (17).

Tank Leaching Study. The results of tank leaching study are expressed as the rate
of metal ions released per unit geometric surface area of the stabilized Orimulsion
LSD product, i.e., metal flux. The flux of leached metal was calculated using the
following equation:

CmxV

J=	

MW x A x t

where

J = Metal flux (mmole/mm2/d)

Cm = Measured concentration (mg/1)

V = Volume of leaching medium (1)

MW = Molecular weight (mg/mmole)

A = Surface area of the leaching sample (mm2)
t = Time lapsed between sample collection (d)

105-10

-------
Tables

Physical Properties arid Concentrations of Ca, V, and S042* in Stabilized Orimulsion
LSD Waste (n=3).

% LSD
Waste

Cement
(%)

OMC

(%)

Max Dry
Density
(lbs/cf)a

Void
Ratio

(%)

Ca

(%)

¥
fag/g)

so42-
(%)

64

0

20

50.7

2.1

NA

NA

NA

64

10

24

55.5

1.9

NA

NA

NA

64

15

26

56.4

1-8

NA

NA

NA

85

0

21

50.3

2.1

NA

NA

NA

85

10

25

55.4

1.9

32.0±2.5

1310±260

18.0±3,0

85

15

29.5

55.2

1.9

32.3±2.6

1270±200

18.0±3.0

a 62.4 lb/cf = 1 g/cm3

Both mixes have nearly identical flux values for Ca (Figure 1). The highest flux
values of Ca occurred within the first 24-hour (5.8 x 10"4 mmole/mm2/d and 7.0 x 10"6
mmole/mm2/d for mix 85%-10 and mix 85%-15, respectively) and decreased
exponentially after 48-hour. The decreasing Ca flux continued throughout the period
of the experiment. At the end of the study, approximately 8% of the total Ca was
released from mix 85%-10 and 10% from mix 85%-15.

For the leaching of V (Figure 2), both mixes exhibited similar flux trends with the
highest flux values occurred within the initial 12-hour (1.3 x 10"4 mmole/mm2/d).
Thereafter, the flux values decreased exponentially and approached -9.8 x 10"*
mmole/mm2/d by the end of the study. Approximately 0.2% of the total V was
released from both mixes by the end of the study.

For the leaching of S042", both mixes also showed very similar flux trends (Figure 3).
The highest flux values were detected initially, which was 1.9 x 10"5 mmole/mm2/d
for mix 85%-15 and 1.6 x 10"5 mmole/mm2/d for mix 85%-10. After initial leaching,
the S042" flux decreased exponentially until the 100-hour period and leveled down
slowly. Approximately 4% of S042" from both mixes was released by the end of the
study.

The flux values detected during the initial 24-hour period for Ca, V, and S042" can be
attributed to the availability of soluble fractions on the surfaces of the stabilized

105-11

-------
products. Shieh et al. (23) reported that leaching of elements from stabilized ash
products was confined to the surface layer less than 1 cm. The leachable Ca
originated from the ca-minerals present in the scrubbing reagents and the added
Portland cement, such as calcium hydroxide (Ksp = 6.5 x 1CT6), gypsum (Ksp = 2.5 x
10"5) and lime (17,24). As the leachable Ca faction on the surface of the stabilized
product becomes depleted, leachable Ca from the inner regions become an important
source contributes to leaching. Leachable Ca within the stabilized product can
diffuse toward the surface and into the leaching medium (25).

Similar leaching characteristic to Ca was also found for V leaching. The rapid
decrease in flux values after the initial 24-hour indicated that a major portion of
leachable V was released initially from the surface of the sample when the stabilized
product was exposed to the test solution.

The soluble form of SO/", such as flue gas gypsum which was available on the
surface and within the stabilized product, was released to the leaching medium.
Similar to Ca and V, the observed rapid decrease in sulfate flux after 24-hour could
be attributed to the depletion of soluble gypsum on the surface of tested samples.
Sulfate release is sustained by the presence of leachable S042" within the stabilized
product The flux continued to decrease, however, demonstrates the effect of
tortuosity and S042" depletion in the outer layers.

Column Leaching Study. The results for column leaching are expressed in terms of
the rate in moles of Ca, V, and S042* per hour. The following equation was used to
calculate the leaching rate (mole hr'1).

Cm x V x MW

R =	

t

where

R = Rate of leaching (mole/hr)

Cm = Measured concentration (g/1)

V = Volume of leachate collected (1)

MW = Molecular weight (g/mole)

t = Time lapsed between sample collection (hr)

The pH in the column leachate was high initially, near 12, and gradually decreased to
pH=ll by the end of the study period (360 hr). The leaching rate of Ca for both
mixes (Figure 4) was highest initially then decreased to a rate of 2 x 1£T5 mole/hr by
the end of the study. Mix 85%-10 showed a higher initial Ca leaching rate (1.1 x W4
mole/hr) than Mix 85%-15. At the end of the study approximately 10% of total Ca
was released from mix 85%-10 and 6% from mix 85%-15.

105-12

-------
High leaching rate of sulfate occurred within the initial 12-hr (24 to 60 x 10"5
mole/hr). The leaching rate was then decreased to less than 5 x 10"* mole/hr near
150-hour and remained level to the end of the study (Figure 5). The trends of
leaching rate for SO/" was very similar to that for Ca. At the end of the study,
approximately 4-5% of total S042" was released from both mixes.

The leaching rate of V for stabilized LSD mixes was somewhat different. Large
fluctuation in leaching rate was observed (Figure 6). At the end of the study,
approximate 1% of .total V was released from mix 85%-10 and 0.5% from mix 85%-15.

For column leaching, as test solution flowed through the samples, more surface area
within the stabilized LSD waste were exposed to the leaching medium.

Consequently, leachable elements originated at various points within the stabilized
product would become mobile when came in contact with test solution. Soluble Ca
within the stabilized product was readily available wherein Ca compounds, such as
calcium hydroxide, gypsum, and unreacted lime, contributed to the dissolved Ca
detected in the leachate.

The leaching characteristic of sulfate was similar to calcium. Soluble forms of S042*,
such as gypsum, was released from the stabilized product. High initial leaching rate
indicated that the majority of soluble sulfate was released immediately into the
leaching medium.

Conclusions

Based on the results of the study, a cement (Type I Portland) content of 10 to 15
percent mixed at a water content of 3 percent above the optimum moisture content
(OMC) is recommended for the stabilization of the Orimulsion LSD wastes to
increase dry density and to meet the unconfined compressive strength and hardness
requirements for landfill disposal. Stabilization of Orimulsion LSD waste minimized
the leaching of Ca, V, and S042". In general, the highest leaching of elements from
stabilized products occurred during the initial exposure periods, thereafter, leaching
decreases exponentially.

105-13

-------
References

1.	R. Rudy, Private Communication.

2.	K. Felsvang, B. Brown, and R. Horn, "Proceeding of title NLA Retrofit FGD
Effective Use of Lime Seminar," Philadelphia, PA, January 9-10,1991.

3.	C. Shieh, I.W. Duedall, E.H. Kalajian, and J.R. WOcox, 'Stabilization of Oil Ash
for Artificial Reefs: An Alternative to the Disposal of Oil Ash Waste." The
Environmental Professional, Vol. 11, pp 64-70 (1989).

4.	A.R. Rodriguez, D.J. Castillo, and F.G1. Sowers. Soil Mechanics in Highway
Engineering. New York: Trantech Publications Company, 1988.

5.	P.H. Wright and R.J. Paquette. Highway Engineering. New York: John Wiley
and Sons, 1987.

6.	O.G. Ingles and B.V. Metcalf. Soil Stabilization: Principles and Practice.
Australia: Butterworths Pty. Limited, 1972.

7.	C.M. Thompson, R.D. Anchord, and G.M. Blythe. "Laboratory Characterization
of Advanced S02 Control By-Products: Spray Dryer Wastes." Prepared for
Electric Power Research Institute, Palo Alto, California, May 1988.

8.	ASTM Annual Book of ASTM Standards. American Society of Testing and
Materials. Philadelphia. Pa., 1987.

9.	E.J. Yoder. Principles of Pavement Design. New York: John Wiley and Sons,
1959.

10.	D. Silberman and G.L. Fisher. "Room-Temperature Dissolution of Coal Fly
Ash for Trace Metal Analysis by Atomic Absorption Spectrometry." Analytica
Chimica Acta 106, pp. 299-307 (1979).

11.	M.A. Franson. Standard Methods for the Examination of Water and
Wastewater. American Public Health Association, Washington DC,
pp. 4-207 - 4-208 (1989).

12.	ASTM Annual Book of ASTM Standards. American Society of Testing and
materials, Baltimore, Maryland, 1982.

13.	P.J. Harder, M.J. Marcinak, N.J. Schlotter, A.L. Labotka, and I.W. Duedall.
"The Fixation of Fly Ash: Physical and Leachate Properties." Consolidated
Edison Company of New York, Inc., New York, pp. 121-169 (1981).

14.	I.W. Duedall, J. S. Buyer, M. G. Heaton, S. A. Oakley, A. Okubo, R. Dayal, M.
Tatro, F. J. Roethel, R. J. Wilke, and J. P. Hershey. "Diffusion of Calcium and
sulfate Ions in Stabilized Coal Ash." In: Waste in the Ocean. Vol. 1. Industrial
and Sewage Wastes in the Ocean. I.W. Duedall, B. H. Ketchum, P. K. Park,
and D. R. Kester, eds., Wiley-Interscience, New York, PP. 375-395 (1983).

15.	H.A. van der Sloot, J. Wijkstra, C. A. Van Stigt, and J.Hoede. "Leaching of
Trace Elements from Coal Ash and Coal-Ash Products." Iru Waste in the
Ocean. Vol 4. Energy Wastes in the Ocean. I. W. Duedall, D. R. Kester, P. K.
Park, and B. H. Ketchum, eds. Wiley-Interscience, New York, pp. 468-497
(1985).

16.	H.A. van der Sloot and G.J. De Groot. "Characterization of Municipal Solid
Waste Incinerator Residues For Utilization: Leaching Properties." Proceedings
of the Ash Utilization and Stabilization Conference (Ash H), Washington (1989).

105-14

-------
17.	I Odler and K.-H. Zysk. "Characterization of Products of Primary Flue Gas
Desulfurization." Ed. Gregory J. McCarthy et al. Materials Research Society
Symposia Proceedings 113, pp. 179-185 (1987).

18.	D. Rai, S.V. Mattigod, C.C. Ainsworth, and J.M. Zachara. "Leaching Behavior
of Fossil Fuel Wastes: Mineralogy and Geochemistry of Calcium". Ed. Gregory
J. McCarthy. Materials Research Society Symposia Proceedings 86, pp. 3-15

(1986).

19.	S.V. Mattigod, D. RaiJ.M. Zachara, and J.E. Amonette. "Mineralogy of
Weathered Flue Gas Desulfurization Sludges." Ed. Gregory J. McCarthy et al.
Materials Research Society Symposia Proceedings 136, pp. 3-8 (1988).

20.	AG. Eklund and D.M. Golden. "Characterization of Solid Wastes from
Advanced SOa Control Technologies for Electric Utilities." Ed. Gregory J.
McCarthy. Materials Research Society Symposia Proceedings 113, pp. 173-177

(1987).

21.	F.A. Cotton and G. Wilkinson. Basic Inorganic Chemistry. New York: John
Wiley and Sons, p. 388 (1976).

22.	A. Lea. The Chemistry of Cement and Concrete, 3rd ed. Arnold, London, pp.
177-249 (1970).

23.	C. Shieh, I.W. Duedall, E.H. Kalajian, and F.J. Roethel. "The Technology of
Energy Waste Stabilization for Utilization in Artificial Reef Construction." Ed.
D.W. Tedder and F.G. Pohland. Emerging Technologies for Hazardous Waste
Treatment American Chemical Society Symposium Series, Washington, DC
(1990).

24.	D.M. Roy, K. Luke, and S. Diamond. ''Characterization of Fly Ash and Its
Reaction in Concrete." Ed. Gregory J. McCarthy et al. Materials Research
Society Symposia Proceedings 43, pp. 3-20 (1985).

25.	T. Edwards and I.W. Duedall. "Dissolution of Calcium From Coal-Waste
Blocks in Freshwater and Seawater." Ed. Iver W. Duedall. Waste In The
Ocean Volume 4: Energy Wastes In The Ocean. New York: Wiley, p. 22 (1985).

Acknowledgements

Hie authors thank Blair A. Kennedy, The New Brunswick Power
Commission, Dalhousie Generating Station, for providing the Orimulsion fly ash that
made this study possible.

105-15

-------
10

~o

o

CD

I

o

X

ll

Mix 85%-10
	 Mix 85%-15

o

0

100 200 300 400
Time (hr)

500 600

Figure 1. Flux of calcium from stabilized Orimulsion LSD waste in the tank leaching
study,

200

I

TJ

CM

0)

o

E 100

<0

I

o

X
LL.

0

0

100

200

300

Time (hr)

400

500

600

Figure 2. Flux of vanadium from stabilized Orimulsion LSD waste in the tank
leaching study.

105-16

-------
CM

0)

o

CD

i

o

x
J=?

LL

100 200 300 400 500

600

Time (hr)

Figure 3. Flux of sulfate from stabilized Orimulsion LSD waste in the tank leaching
study.

120

0)

o

(£>

I

o

0

4->

m
DC

500

Time (hrs)

Figure 4. Leaching rate of calcium from stabilized Orimulsion LSD waste in the
column leaching study.

105-17

-------
30

n

CI)

— 20
O CX}

E

0>

o

-------
WHEN THE REGULATIONS GET TOUGH;
ADVANCED TREATMENT OF FGD BLOWDOWNS

Marek K. Mierzejewski
Infilco Degremont Inc.
2924 Emerywood Parkway
Richmond, Virginia 23294

David C. Olszewski
Keith R. Minnich
Aqua-Chem, Inc.

7800 North 113th Street
Milwaukee, Wisconsin 53201

Abstract

In cases where blowdowns from FGD systems are discharged to the environment,
they must first be treated, principally to remove heavy metals and suspended solids.
The established treatment techniques of metal precipitation and complexation followed
by clarification, while achieving these objectives, do not alter the wastewater's salinity.
If Federal or State regulations limit salt input into the environment, then evaporation of
the wastewater is required. Furthermore, the increasingly strict discharge limits being
set for heavy metals may mandate this zero discharge technology even if no salinity
controls are imposed.

Similarly, if organic additives are used in the absorber, the resulting high Biochemical
Oxygen Demand (BOD) of the blowdown must be removed biologically prior to
discharge.

To date, experience of these more advanced techniques is limited, with only one
evaporator working on an unblended FGD blowdown (in Germany), and with only a
few biological trickling fitters tackling the BOD problem.

This paper reviews key issues and precautions to be considered in designing FGD
wastewater treatment plants to these toughest new standards.

Introduction

The sources and characteristics of FGD wastewaters and techniques of their treatment
have been discussed in an earlier paper(1). Since presentation of that paper in
December 1991, three more contracts for FGD wastewater treatment plants (WWTP)
have been awarded, making a modest total of four such plants in the U.S. dedicated
to the treatment of unblended FGD blowdowns, in response to Phase I of the 1990

106-1

-------
Clean Air Act Amendments. Although only one of these plants has begun operation
(NIPSCO Bailly, in June 1992®), all the others are at varying, but early, stages of
design/construction. Since these plants are normally associated with saleable
gypsum FGD systems, more may be expected during Phase II, where the increasing
cost of waste FGD sludge disposal seems likely to lead to more utilities adopting the
saleable gypsum approach.

Of course, there are several more FGD systems installed, or being installed, than the
number of WWTP, and this reflects the easy first option of blowdown disposal that
utilities consider, namely discharge into ashponds. With tightening regulations, this
option may become less available.

The four U.S. installations have little in common: they are each in a different state, they
are designed to treat blowdowns from different FGD suppliers, and are required to
meet varying standards of treatment (Table 1).

Discharge standards vary from state to state and are being applied increasingly based
on water quality, rather than technology, standards, in effect requiring treatment to
produce an effluent having similar if not lower concentrations of species as the water
body into which it is discharged. Limits have been proposed for metals where an
alternative to the established treatment of these blowdowns has had to be considered.
For example, the New York State Electric and Gas Corporation (NYSEG) had little
option in the case of its Mil liken Station, located on the edge of Cayuga Lake, NY, but
to adopt a zero discharge option.

As regards BOD removal, the FGD WWTP at Pennsylvania Electric Company's
Conemaugh Station will include a biological stage using Sequencing Batch Reactors
(SBRs) to reduce the influent BODs from 290 mg/l (if dibasic acids (DBA) are dosed)
to 25 mg/l, a removal efficiency of over 90%.

Evaporation

Metal limits are one thing: total dissolved solids (TDS) another. Evaporation, yielding a
highly concentrated liquor or a crystalline product, is the technology of choice where
TDS limits are proposed. In Europe, it is debatable whether evaporation would be
proposed for any other purpose. This, however, reflects the generally "easier*
discharge standards that have to be met there as compared to the U.S. On some
projects in the U.S. very low allowable concentrations of heavy metals have been
demanded, considerably lower than in Germany, but with no limit on discharge salinity.
In this case, evaporation has nevertheless had to be considered, since physico-
chemical precipitation and coagulation alone, even using sulfides, may not attain the
set limits.

In Germany, discharges from FGD systems are regulated by the 47 Verwaltungs-
vorschrift (47 VwV) which, although it represents a minimum Federal limit that can be

106-2

-------
locally tightened, seems generous by U.S. standards. By contrast, the U.S. EPA,
through State Departments of Environmental Protection, is imposing standards which,
in many cases, exceed the German limits by one, and sometimes two, orders of
magnitude. The 47 VwV is currently under review, but even the proposed new
discharge limits make an interesting comparison with those for a current U.S. FGD
project (see Table 2).

The implications of these tighter standards are that the existing technology must be
further optimized and fine-tuned, with the additional requirement that the WWTP is
carefully operated. The alternative is to introduce supplementary technology, such as
evaporation, to achieve the limits.

The purpose of an evaporation/crystallization system is principally to minimize the
volume of the FGD blowdown, by producing a concentrated brine: an 80% volume
reduction is typical for an evaporator, and over 90% is possible using salt conversion.

Pretreatomnt Necessary for Evaporation

This pretreatment takes two forms: firstly the treatment that is required to permit
disposal of the salt product (e.g. heavy metal removal, suspended solids removal) and
secondly the pretreatment for the evaporator proper.

If a relatively high purity product is required, then heavy metals need to be removed
by the existing technology described earlier. This necessitates, in effect, a full
physicochemical precipitation/coagulation/clarification plant upstream of the
evaporator itself. Clarification is also required to reduce influent suspended solids to
less than 100 mg/l.

FGD blowdowns are saturated with calcium sulfate and other calcium sails. To
prevent these salts from scaling the evaporator heat transfer surface the evaporator is
"seeded" with CaS04 crystals so that instead of precipitating on the heat transfer
surface the salts precipitate on these crystals. The successful operation of seeded
slurry systems requires that the feed chemistry conforms to several parameters, and
pretreatment by addition of one or more compounds may be required.

A feed/distillate heat exchanger is used to preheat the feed with outgoing distillate to
improve the system energy economy. Anti-scaling agents are metered into the feed to
temporarily increase calcium salt solubility, thereby reducing scaling in the heat
exchanger only and facilitating proper deaerator performance.

Other pretreatments include:

• Acidification to a pH of 5.5 to 6.5 which converts carbonates to dissolved C02
which is removed in the deaerator.

106-3

-------
• Anti-foam dosing, if necessary, to control foaming and reduce solids carryover
into the distillate.

The feed is sprayed into a deaerator to reduce the amount of dissolved gases, it being
particularly important to remove dissolved oxygen to minimize corrosion.

Salt Conversion. Salt conversion is another pretreatment technique that has
advantages in the treatment of FGD blowdowns. This technique replaces the calcium
by sodium ions, allowing the treated FGD blowdown to be evaporated and crystallized
to yield a high purity sodium chloride product. This has the benefit of further reducing
the amount of waste that must be disposed and may yield a salt product that is
saleable. Salt conversion may be required when disposal of the wastes is very
expensive or not an option. Its major drawbacks are the additional processing
equipment required and its capital and operating costs. This option almost doubles
the capital cost when compared to evaporation of the calcium chloride based
wastewater. Although the operating costs of the evaporation systems are slightly
lower, the chemical costs are significantly higher, due to the amount of sodium
carbonate required.

Before salt conversion, tie Mg2+ must be removed by raising the pH, thereby
precipitating Mg(OH)2. Salt conversion itself is performed in a discrete reaction tank
after physicochemical pretreatment. Sodium carbonate is dosed to a mixed reaction
tank and the following reaction takes place:

Ca2+ + NaaCOg - CaC03 + 2Na+

The stoichiometry is such that a considerable amount of soda ash is required. For
example, a 30 gpm (6.8 m3/h) flow, containing 6 g Ca2+/L will consume 2.6 tonnes
Na^COj/d (2.9 US ton/d). The only positive point is that there is no sludge disposal
problem, as the calcium carbonate produced (2.45 tonnes/d in this example) can be
fed into the scrubber with the limestone slurry.

Most Suitable Evaporator Types

Evaporators can be grouped into two broad categories: thin film and forced
circulation. In the first type the brine is distributed over the heat transfer surface in a
thin film, with boiling taking place at the surface of the film. The distribution system
must be designed to completely wet the heat transfer surface. In forced circulation
evaporators, brine at its boiling point is pumped from a flash tank through a heat
exchanger where it is sensibly heated under pressure to prevent boiling in the heat
exchanger. From there, the heated brine is returned to the flash tank where it then
flashes at a lower pressure.

106-4

-------
In both types of evaporator the brine flow required in the heat exchanger exceeds the
feed flow and a recirculation pump provides the necessary flow to the heat exchanger
with the incoming feed being combined with this flow. After boiling takes place a small
portion of the flow is withdrawn as concentrate.

Mechanical vapor compression or multiple effect configurations can be used with
either type of evaporator. In mechanical vapor compression, the evaporated vapor
passes to the suction of a steam compressor: the compressed vapor is then used as
heating steam in the evaporator heat exchanger. In multiple effect evaporation, boiler
steam is used to evaporate vapor in the first effect. This vapor is used as the heating
steam in the second effect, which operates at a lower pressure and temperature than
the first effect. This process can be repeated several more times, until vapor from the
last effect is condensed in a heat exchanger which normally uses cooling water.
Mechanical vapor compression thus recycles the latent heat of vaporization, while
multiple effect evaporation uses the latent heat of vaporization of boiler steam several
times.

Thin film evaporators require much smaller recirculation pumps than forced circulation
evaporators. The smaller pump reduces both the capital and operating cost. This is
particularly significant when high alloy materials are required in hot brine service.

These evaporators are generally limited to non-crystallizing applications. However,
salts that are inversely soluble with temperature, such as CaS04, can be crystallized in
thin film evaporators by using the seeding technique. Forced circulation crystallizers
are used for concentration of brines that are saturated in NaCI, and would therefore be
used in any salt conversion system.

A calcium based FGD purge stream can be concentrated in a thin film evaporator to a
total solids concentration of approximately 33%, which is below the crystallization point
of CaCI2 in a solution at ambient temperature. A filter press system can be used to
remove CaS04 crystals from the 33% CaCI2 solution to minimize the amount of
impurities in the brine. A dry CaCI2 salt can be produced by transferring the 33%
solution to a spray dryer, although the usefulness of obtaining a dry CaCI2 product is
questionable, given this salt's deliquescent nature.A sodium based FGD blowdown
stream can be concentrated in a thin film evaporator to a total solids concentration of
approximately 25%. Further concentration of the sodium based brine can be
accomplished in a forced circulation evaporator. The NaCI crystals, which must be
removed by a dewatering device such as a centrifuge to limit the concentration of the
slurry, can be dried in a drum or fluid bed dryer. As with all aspects of FGD
blowdown handling and treatment, correct selection of the materials of construction is
very important. In evaporation, the higher temperatures and concentrations mandate
materials such as those given in Table 3. Rubber-lined carbon steel can be used as a
substitute material where temperature limitations allow.

106-5

-------
Utility Requirements

Table 4 shows the utility requirements per gpm (0.23 m3/h) of feed, as determined by
product type: 33% total solids calcium chloride solution, calcium based dry salt and
sodium based dry salt (i.e., after salt conversion). Comparative figures are given for
mechanical vapor compression (MVC) and multiple-effect (ME) evaporators. As
mentioned above, the utility costs for the sodium based dry salt option are lowest, but
chemical costs (not shown here) are the highest.

Evaporator Product Disposal

Liquid disposal requirements can be completely eliminated by the addition of
crystallization or spray drying after evaporation. Crystallization is less expensive and
less energy intensive than spray drying but cannot be used with calcium chloride
based wastewater.

Spray drying yields a dry product that is easier to dispose of than that from
crystallization, which yields a wet product (3-10% H20). An additional drying step can
be used to produce a dry product after crystallization. This is normally required if the
product is to be sold such as in the case of salt conversion. Some U.S. utilities are
investigating the possibility of using the liquid calcium chloride solution for road de-
icing. This requires pretreatment to remove heavy metals and other toxic
contaminants.

Experience: Full Scale and Demonstration

Evaporation of FGD blowdowns worldwide has been very limited to date. A power
station in Berlin, Germany, has the only known FGD wastewater evaporators currently
operating. Indeed Germany was expected to be the leader in this new environmental
treatment, but economic problems there have apparently delayed its implementation.
The U.S. has no FGD blowdown evaporators currently operating, although the first
system was purchased earlier this year (by NYSEG) and will be in operation in 1994.
With the current activity in the U.S. FGD market, there are numerous wastewater
evaporators proposed and rt appears the U.S. may become the leader in this field.

Regarding demonstration scale testing, Philipp Muller, the German affiliate of the
Degremont Group, has joint-ventured with Aqua-Chem Inc. to run tests with an
evaporator and crystallizer system at a German power station later this year.

The system has been designed with the flexibility to handle sodium chloride or calcium
chloride based solutions, either unseeded or seeded with gypsum. With the high
grades of materials used (i.e., titanium, Hastelloy, Alloy 20), it can handle a wide range
of chemistries, concentrations and temperatures. It also has many built-in automation

106-6

-------
features to reduce operator attention. The flow diagram for this plant is represented in
Figure 1.

The FGD blowdown, after heavy metals removal and gypsum desaturation, is
pretreated to remove the magnesium and convert the calcium to sodium. The
wastewater is then concentrated in two steps. In the first step, the falling-film
evaporator concentrates the solution to about 25% TDS (just under the solubility limit
for NaCI). Gypsum seed crystals are added if necessary to maintain a seeded slurry
operation. In the second step, the forced circulation evaporator/crystallizer
concentrates the solution further to produce a slurry with suspended salt crystals. A
portion of the slurry is discharged to a filter press to remove the salt crystals and the
other liquid is returned for further concentration. If required, the salt crystals can be
washed in a batch mode for higher purity.

COD and BOD Removal

As well as treating FGD blowdowns to new strict metal limits, the WWTP may need to
be adapted to reduce the concentrations of BOD5 and Chemical Oxygen Demand
(COD). To date, these stages have been required in some WWTP's in Germany and
the Far East, where they are introduced after the clarification stage. The COD of FGD
blowdowns is due principally to unoxidized sulfite ions and trace organics;
environmental authorities in the Far East also require removal of dithionate (S2062")
which can contribute to the COD.

Unreacted sulfite is removed by oxidation; hence, if the influent sulfite level is known to
be low, the first (oxidation) stage can be dispensed with. Trace organics are removed
by adsorption on granular activated carbon in pressure filters. With these two stages
the COD may be reduced from an influent level of 150-200 mg/l to less than 100 mg/l.
However, if dithionate is a regulated species, then ion exchange on a special synthetic
adsorbent is also required.

The BODs of FGD blowdowns is usually very low, and the usual discharge standard of
20-30 mg/l seen in Europe and the U.S. imposes no requirement for special treatment.
Only in the case where organic additives are used in the scrubber is a BOD removal
stage required. Organic additives, used to enhance S02 removal in the scrubber, are
normally mono- or dibasic carboxylic acids (such as formic or adipic acids), which
serve to buffer the scrubber pH to an optimum level. Being highly soluble, these
compounds pass through the scrubber and into the blowdown, from which they must
be removed to meet BODs discharge guarantees. One German FGD supplier's
process is based on a continuous 800 mg/l formic acid feed: in three WWTP treating
the blowdown from Saarberg Holter Umwelttechnik (SHU) scrubbers at Heyden,
Bexbach and Fenne Power Stations, the BOD is removed after physicochemical
treatment by a biological trickling filter.

106-7

-------
The BOD of carboxylic acids, expressed in grammes per gramme of pure substance,
increases as the carbon chain grows. Thus, formic acid has a BOD of 0.24 g/g,
acetic add 0.65 g/g and propionic acid 1.1 g/g. Tests done on DuPont DBA (a
mixture of succinic, glutaric and adipic acids - 4, 5 and 6 carbon chain respectively)
show a BOD similar to propionic acid. Thus, a DBA dose of 300 mg/l may be
expected to give a BODs of about 330 mg/l.

The acids are also readily biodegradable: in deionized water, a 300 mg/l DBA solution
showed an oxygen uptake rate (OUR) of 60-75 mg Og/Lh, evidence of very rapid
biological oxidation. When the DBA was spiked to the same concentration in an FGD
blowdown having a high total dissolved solids content (TDS) the OUR dropped to
20-30 mg 02/l.h, still showing a high biodegradability, similar to that seen with
domestic sewage®.

In common with evaporation of FGD blowdowns, there is very Irttie experience
worldwide with their biological treatment. Except with SHU installations in Germany,
the biological treatment of blowdowns containing organic additives has simply not
been encountered. The simplest carboxylic acid, formic acid, has been the subject of
most investigations, not least because of its extensive use by SHU. In SHU systems,
this acid is dosed at a recommended level of 800 mg/l, but this has been reduced in
several installations to 500 mg/l. Its function is to reduce the pH in the scrubber, to
about pH 4 in the spray zones and pH 5 in the sump, thereby enhancing the reaction
between sulfur dioxide {SOJ and limestone. In this way, a high efficiency removal of
S02 can be maintained at an elevated chloride concentration of say, 40,000 mg/l: in
effect, the organic additive helps to override the decrease in efficiency normally seen
with an increase in chloride concentration.

Pilot tests performed by Bettenworth etal.(4) investigated biological formic acid
oxidation using a two-stage process: a cinder packed trickling filter followed by
filtration through a bed of 1-3 mm grain size Biolite, a high surface area granular
medium.

The pilot was designed so as to maintain both the trickling filter and the biofilter in an
oxic state; with air being diffused into the bottom of the first unit, the effluent entering
the second filter was almost saturated with oxygen, and still contained 3-4 mg 02/l at
the outlet. Over a several week period (to allow for acclimation) the formic add and
chloride concentrations were brought up to 100 mg/l and 9,000 mg/l respectively.
The removal of COD, which is approximately 1.25 times the BOD concentration, never
dropped below 90%, with almost all (>97%) of the removal taking place in the trickling
filter, the second filter acting principally to remove biomass particles.

As with any biological system, the dassic BOD5: N : P nutrient ratio of 100:5:1 was
respected during these tests by appropriate dosing, and should always be followed in
full-scale plants.

106-8

-------
On the basis of these tests, a full-scale WWTP, combing both physicochemical and
biological treatment was installed at Saarbergwerke's Heyden station in Western
Germany in 1985. This was followed by another two WWTP at Bexbach and Fenne
stations.

The trickling filter at Bexbach is a 6 m (20 ft) diameter x 6 m high forced draft unit,
filled to 4.5 m depth with plastic media: physicochemically-treated water is fed over the
media by a rotary distributor. Two pressure filters follow the single trickling filter: the
design filtration velocity is 7 m/h (2.9 U.S. gpm/sq.ft.) through a 1.5 m (5 ft) deep
Biolite bed. (See Figure 2).

A 4-5 times dilution of the FGD wastewater with cooling tower water is practiced at
Bexbach to reduce the chloride and formic acid concentrations entering the biological
treatment and, more importantly, to drop the calcium sulfate concentration below the
equilibrium point. FGD wastewaters are frequently supersaturated in calcium sulfate,
and equilibrium is re-established during physicochemical treatment. However effective
such "desaturatiorf may be, there remains a risk that some calcium sulfate will
precipitate on the packing in the trickling filter. Furthermore, any unreacted lime
remaining after physicochemical treatment may also react with the carbon dioxide
(both atmospheric and respiratory) in the air in the trickling filter. This double risk of a
heavy mineral scale forming on a lightweight, high surface media might be expected to
lead to problems and this indeed has been the case. Firstly, the scale can prevent
good biomass attachment, such that BOD removal efficiency drops. More seriously,
collapse of the media due to the weight of the scale has been encountered in one
plant.

Alternative Biological Treatment

The choice in biological treatment lies between "attached growth", as described above,
or "suspended growth", such as activated sludge, or some combination of the two.

None of these options presents itself as an obvious choice for ail plants. If the
problems of scaling can be overcome (for CaS04 by better desaturation or by dilution
and for CaCOs by add dosing to a pH closer to neutrality) then attached growth is the
preferred method for low BOD wastewaters. In such cases, the trickling filter/biofilter
combination, as at Bexbach can be considered or alternatively, one of the new
generation of oxygenated upflow biofilters available today.

If, however, higher BOD levels, of 200 mg/l and above, are likely to be encountered,
then a suspended growth system is recommended. The argument against using this
option with low BOD wastewaters is that it is difficult to maintain the Mixed Liquor
Volatile Suspended Solids (MLVSS) of 2000-3000 mg/l, i.e., the biomass concentration
necessary to attack the BOD.

106-9

-------
All suspended growth processes depend on a wastewater entering an aeration basin,
where intimate mixing with the biomass and oxidation takes place. The treated
wastewater leaves the basin with the biomass, which must then be recovered, usually
in a clarifier, and returned to the basin to maintain the level of biological activity
necessary there (Figure 3). Solids recovery thus becomes critical: its efficiency
determines the viability of the system. Typically, suspended solids in the overflow from
a secondary clarifier will be in the order of 30 mg/l. An influent BOD5 of 50 mg/1, say,
may yield approximately 30 mg/l of new synthesized biomass: in such a case the
plant will be run on a knife-edge, always hovering between depleting the aeration
basin (where overflow solids > synthesized solids) and just maintaining the MLVSS
(overflow solids = synthesized solids). In actual practice such mathematically
straightforward conditions do not exist, and fluctuations in flow, disruption or
malfunction of operation can all cause excessive solids loss from the system.

The other problem facing suspended growth systems is similar to that which faces its
attached growth counterpart, viz. precipitation of CaS04 and CaC03. Whereas in a
trickling filter this phenomenon is most likely to manifest itself as scale, in a suspended
growth system these compounds may form discrete crystalline precipitates which will
settle readily: the problem then is how to purge them from the system. This requires
that these precipitates be both suspended by the aeration/mixing system and then,
after settling, be removable from the secondary clarifier. The first condition demands
vigorous mixing using static draft tube aerators or jet aeration, with an energy input at
least double the 20-30 W/m3 (75-115 W/1000 gal) energy input used in conventional
activated sludge design. Porous discs, used for fine bubble aeration, are to be strictly
avoided in view of their potential for mineral scaling and blockage in this application.

The second condition requires a scraped clarifier that will remove the solids as soon
as they are settled and not allow them to thicken to an immovable mass. Sequencing
Batch Reactors, where aeration and settlement is alternated sequentially between two
or more tanks (i.e., each tank serving alternately as aeration basin and clarifier), have
been recommended at one U.S. FGD installation, but are as yet without reference.
Their benefit certainly is in that the duration of all sequenced stages can be modified,
after installation, to obtain optimum treatment. The likely disadvantage with SBR's is
that since they are fiat-bottomed, unscraped tanks, mineral sludge will tend to settle
and thicken, gradually reducing the effective volume.

Summary

The tightening restrictions on all wastewater discharges, not just FGD blowdowns,
mandate that existing technologies of wastewater treatment do not stand still, but are
modified and supplemented accordingly. In the case of FGD blowdowns specifically,
the limits on heavy metals may soon be such that evaporation (i.e., the station
becoming zero discharge) will be the rational solution. The problem will, however, be
replaced by another one: hew to dispose of the salt liquor or cake produced (a
question not unlike that of saleable gypsum disposal). The option of discharge into

105-10

-------
ashponds will not be ever present. Likewise, in throwaway systems, the wastewater
treatment problem will not always be "buried" with the sludge; waste FGD sludge
leachate treatment is already under investigation for a Southeastern U.S. utility.

The restrictions put on BODs in the discharge are quite different to those put on heavy
metals and salinity, in that they are avoidable. The facile solution to achieving
compliance may be simply to discontinue the use of the offending organic additive,
since a "typical" FGD blowdown, without added organics, will meet the usual BOD
discharge limits seen in U.S. and Europe of 20-30 mg/l. If, however, enhanced S02
removal is required, and organic additives are used, then treatment to remove the
BOD will be required.

The fact that such treatment is biologically based requires that this stage follow the
physicochemicai removal of heavy metals, without which inhibition of the microbial
processes may occur. Furthermore, the biological removal system requires a constant
feed of substrate (in this case the organic acids) if the bacteria are not to die. This
precludes operation of the FGD system with intermittent or variable application of the
organic additive.

Both these more advanced removal techniques have something in common: there is
very little experience worldwide in either. The first U.S. evaporator installation working
on undiluted FGD blowdown, and one of the first in the world, will be starting up in
1994. Demonstration scale evaporation tests by a U.S. company are also due to start
in Europe this year. As to biological treatment, experience is currently limited to
removal of formic acid at a handful of stations in Germany. (It should be noted that
the zero discharge option of evaporation would also remove concerns over BODs at a
stroke.) According to current forecasts, at least one FGD WWTP in the near future
may be equipped with biological treatment equipment.

The next few years, particularly once Phase II of the Clean Air Act Amendments comes
into operation, should see interesting new developments in this advanced FGD
wastewater treatment technology.

106-11

-------
References

1.	M.K. Mierzejewski, "The Elimination of Pollutants from FGD Wastewaters",
presented at the EPRI/EPA/DOE 1991 S02 Control Symposium,

Washington DC (December 1991).

2.	Michael Sicinski, David LaValle and Marek K. Mierzejewski, "Cleaning the Water
at a Clean Air Plant; Early Operating Data from the NIPSCO Bailly FGD
Wastewater Treatment Plant", Paper lWC-92-50, 53rd International Water
Conference, Pittsburgh, PA (October 1992).

3.	Mark J. Briggs, Radian Corporation, Milwaukee, Wl. [personal communication]

4.	H. Bettenworth, H. Pflug, E. Dieterle and M. Wolkl, "Untersuchungen zum
biologischen Abbau von Ameisensaure im Abwasser von Rauchgasent-
schwefelungsanlagen", VGB Kraftswerktechnik (Germany), Vol. 64, No. 11,
p. 993-999 (1984).

106-12

-------
*

TBS
M +

Table 1

U.S. FGD Wastewater Treatment Plant

Utility

State

Station

FGD System

Oneration

NIPSCO

IN

Bailly

Pure Air

1992

NYSEG

NY

Mi Hi ken

SHU

1994

Atlantic
Electric

NJ

B.L England

GEESI

1994

Penelec

PA

Conemaugh

ABBES

1994

Controlled species are given in this column
Total Suspended Solids
Various metals, heavy and other

Treatment*

pH.TSS
Zero Discharge

pH, TSS, M +

pH, TSS, M + ,
BOD

Table 2.

Comparison of Germany's 47 Verwaltungsvorschrjft (47 VwV)
Discharge Limits with those from a current U.S. Utility FGD Project

Species

47 VwV

u§

Cd

50

100

Hg

50

2

Cr

500

50

Ni

500

30

Cu

500

100

Pb

100

50

Zn

1000

100

V

5000

—

s2-

200

—

Mn

—

500

Se

—

200

As

—

36

All units are parts per billion (jig/l)

106-13

-------
Table 3.

Materials of Construction for Evaporators



cr < 20.0001

cr < 50.000

Cr > 50.000

Tubes

Tl Gr 2

n Gr 2

Ti Gr 2, 12 or 7
Hastelloy C

Upper Tubesheet

Ti-clad

Ti-clad

Ti-clad or
Hastelloy clad

Lower Tubesheet

316 with
2.5 Mo

Ti-clad
6 Mo SST2

Ti-clad

Hastelloy clad

Lower Shell

316 min Mo

6 Mo

Hastelloy clad

Demisters

317

6 Mo

Hastelloy

Piping

317

6 Mo

Hastelloy

Pumps

Cd4MCu3

Cd4MCu

Hastelloy > 180°F
Cd4MCu < 180°F

Upper Channel and
Vapor Separator

316 min Mo

6 Mo or
Hastelloy clad

Hastelloy clad

Distributor

Inconel 625

Inconel 625

Hastelloy

Notes:

1	: All chloride [CP] concentrations in mg/l.

2	; *6 Mo stainless steel = AL6XN, 254SMO, or equivalent

3	; Cd4MCu should contain nitrogen (0.15% min.)

106-14

-------
Steam
kW

(MM BTU/hr)

Electricity
kW

Fuel Gas
kW

(MMBTU/hr)

Cooling Water
m3/h

(gpm)

Table 4.

Utility Requirements By Product Type
per gpm of Feed (7% Total Solids (TS) in Feed)

33% TS Calcium
Based Solution

MCV
0

(0)
9,3

0.11
(0.5)

ME2

0.8

0
(0)

2.95
(13)

Calcium Based
Dry Salt
MVC/ME3 ME

Sodium Based
Dry Salt
MVC ME

44	1^8	cr-f	fi	oq c

(0.150) (0.054) (0.174) (0) (0.135)

6.1

11.7

(0.04)

1.1
(5)

0.8

11.7
(0.04)

3.4

(15)

5.5

1.2

(0.004)

0.07
(0.3)

0.7

1.2

(0.004)

2.52

(11.1)

MVC1 : Mechanical Vapor Compression

ME2 : Multiple-Effect (3-effect for Calcium Based and 4-effect for Sodium
Based)

MVC/ME3: This option requires a combination system.

106-15

-------
o

2

05

Figure 1.

Demonstration Scale Evaporator/Crystallizer for FGD Slowdown Treatment

-------
trickling filter

BLOWER

DILUTION TANK

FROM

PHYSiCOCHEMICAL
TREATMENT

TREATED
WASTEWATER

Figure 2.

Biological Treatment of FGD Blowdown at Bexbach Station

INFLUENT







TREATED

AERATION BASIN

— air

CLARIFIER

WATER

RETURN SLUDGE

WASTE SLUDGE

Figure 3.

Activated Sludge Plant Schematic

106-17

-------
A

-------
DEVELOPMENT OF HASTE WATER CONCENTRATION & SOLIDIFICATION

CVCiPFRf XfAD -utpm fFTTTT? r*lCI flffCTIT T?1TO T 71HTniI "DT liMTn
SIOX&E rUK if IS J. r±jUi!» VsAd IJ£oUJUI: UKi Aill lUu JkrJLj/iNX

S. Tsubouchi
T. Miwada
Chubu Electric Power Co., Inc.
1 Toshin-cho, Higashi-ku
Nagoya, 461-91 Japan

S. Kotake
N. Ukawa
Mitsubishi Heavy Industries Ltd.
2-5-1 Marunouchi
Chiyoda-ku, Tokyo, 100 Japan

Abstract

Wet limestone gypsum FGD system has been widely installed in
coal-fired thermal power stations for environmental protection
by removing the acid-rain causing sulfur oxide compound from
their flue gas. This system inevitably disposes waste water
containing chloride, heavy metals, nitrates and COD
components, which requires processing with a fairly expensive
and complex water treatment system before discharging. In
order to solve this problem, Chubu Electric Power Corp. and
Mitsubishi Heavy Industries have jointly developed a new FGD
waste water treatment system where waste water is collectively
concentrated with the combination of electrodialysis and
evaporation and then solidified with flyash and cement. A
demonstration test with a 20MW equivalent pilot plant
conducted for almost 9 months at the Shin-Nagoya Power Station
adequately proved the efficacy of the system.

Introduction

Recently in coal-burning thermal power stations, large number
of wet type flue gas desulfurization (FGD) plants are being
installed for environmental protection. In desulfurization
plants, sulfurous acid gas in the flue gas is made to be
absorbed by limestone slurry absorbent and recovered as by-
product gypsum while the gypsum filtrate is recycled back into
the system. In this recycled filtrate, a part of the flyash
and halogen etc., contained in the flue gas also get taken in.
Among these, in particular, if cl" ion concentration in the
absorbent fluid becomes high, the materials of the constituent

107-1

-------
equipment and apparatus of the FGD plant get corroded while
its desulfurization efficiency also get affected. As a
countermeasure, a fixed quantity of waste water is constantly
bled out of the absorber to avoid excessive build up of these
harmful components and a waste water treatment facility is
provided to get rid of them.

Presently, in the waste water treatment facility,
corresponding to the characteristics of the waste water
constituent, such high grade treatment systems as; heavy
metals treatment system, fluorine treatment system, biological
nitrogen treatment system, COD treatment system etc., are
provided.

Although these waste water systems have been widely adopted
not only by the electric utilities but by other companies as
well, the fact remains that they are complicated, need large
installation space and are expensive to equip and operate.

To resolve these problems, a desulfurization system without
any waste water discharge has been developed in which the
waste water from the process after being concentrated is
solidified into a stable form by mixing with cement and flyash
to be disposed for land reclamation purposes. The system was
demonstrated with a pilot plant, the summarized results of
which are reported hereunder.

Research, and Development Particulars

The research and development particulars of this facility is
shown in Figure 1. During the elementary research, to carry
out comparative system study in respect of process, material
selection, and economical trial balance, in particular for the
waste water concentration method, laboratory scale tests were
conducted for the three methods; (1) evaporation tower method
using boiler flue gas excess heat (2) method using steam
heating and multi effect evaporator, and (3) combined method
using electrodialyzer and evaporator.

After comparison of the tests made from the technical and
economical aspects, it was confirmed the method (3) to be the
best.

On the basis of these findings, a pilot test plant equivalent
in capacity to a 20 1? electricity generating facility was
installed within the premises of Chubu Electric Company's
Shin-Nagoya thermal power station. The installation and trial
operations of the pilot plant were completed by 1991 year end,
and during the 1992 January-September period it was subjected
to demonstration tests. A general view of the pilot facility
is shown in Figure 2.

As waste water for the test, actual waste water drawn from the
FGD plant of Chubu Electric Power Company's Hekinan Power
Station No.1 coal-burning boiler unit and blended with
chemical additives for adjustment of chemical composition was

107-2

-------
used. Further, the flyash used was also obtained form the
same boiler while the cement used was the market available
common portland cement.

Outline of Waste—Water-Free FGD System Pilot Plant

Pilot plant outline of the waste-water-free FGD system is

shown in Figure 3. The FGD system waste water, typical
composition of which is shown below, is a complex mixture of
flyash, chloride/fluorine ions, heavy metals, nitrogen
compoundsr COD components etc.

Item

pH

SS

Ca2*

Mgz+

Alz+

Fe3+

Properties

5-6
mg/1	500

mg/1 7,000
mg/1 2,500
mg/1	50

mg/1	50

Item	Properties

Na+	mg/1	60

CI"	mg/1	20,000

F"	mg/1	50

T-N	mg/1	440

S042"	mg/1	1,000

COD	mg/1	70

As a pre-treatment, the waste water is filtered to remove the
suspended solids, and then, after being primarily concentrated
with electrodialysis and further secondarily concentrated with
evaporator, is mixed and kneaded with cement and flyash to be
formed into a stable solid substance. Specification of main
equipment for each facility is described hereunder while
outline of each facility is described in the following.

TJnxt

Equipment

Specification

Filtration

Primary
Filter

Model

Filtration area
Treated flow volume

Precoated
Type
1.32 mz
1.4 m3/h

Secondary
Filter

Model

Filtration Area

Treated flow volume

P Or3+*

Type
0.377 mz
1.4 m3/h

Electro-
dialyzer

Electro-
dialyzer

Model	!

Membrane area	:

No.of membrane pairs;

Filter
Press Type
0.368 nr
180 pairs

Secondary
Concentration

Evaporator

Model

ucau llallolcl /ilea

Wiped-Film

Type
0.4 m2

Solidification Mixer	Model

Kneader	Capacity

Vibro-Mixer
188 kg/h

107-3

-------
Filtration Facility

The thickener supernatant, of FGD system waste water used for
this test contains about 500 ppm of suspended solids (SS) . In
order to prevent system performance deterioration due to
clogging and adherence on the electrodialyzer membranes by
these SS compounds, a pre-coated type 2-stage filtration
equipment was provided as a pre-treatment means.

Both the 1st stage and 2nd stage filters used in this
equipment were of ceramic made with their external surface
coated with pearlite as a filtration aid. These filters were
washed at regular intervals.

Electrodialyzer Facility

The filtrate from the filtration equipment is sent to the
electrodialyzer shown in Figure 4. The electrodialyzer is
consisting stacks of narrow compartments through which the
feed solution is pumped. These compartments are separated by
alternating cation-exchange and anion-exchange membranes which
are selectively permeable to positive and negative ions
respectively. The terminal compartments are bounded by
electrodes, for passing direct current through the whole
stack. Plastic-meshes are inserted in the solution
compartments to keep the cation- and anion-exchange membranes
apart and to promote uniform liquid flow.

The "Selemion" membranes of Asahi Glass Co., Ltd, are used for
the pilot test.

When the electrodes are connected to a DC source, ion
migration begins, as shown schematically in Figure 5 for a
group of chambers well within the stack. In each chamber
cations travel from right to left; anions in the opposite
direction. For the chamber D, the anion permeable membrane on
the right does not admit cations from the right, and the
cation-permeable membrane on the left similarly act as a
barrier for the negative ions from the left. As a result, the
filtrate cl" concentration decreases in the chamber D and
increases in the neighboring chambers C.

With this univalent selectivity, S042" ions are kept at low
concentration to enable the gypsum saturated waste liquid
circuit theretofore to be maintained as a unsaturated circuit
in spite of having it subjected to concentration. In
addition, it permits the liquid to be secondarily concentrated
with the evaporator to the final value of 250,000 ppm in terms
of cl" concentration without precipitation of gypsum. The
concentration of the liquid is decided by the density of
current flowing in between the membranes. Increase in current
density increases the concentration degree of the liquid while
a decrease in its value decreases the concentration degree.
As the current density in an electrodialyzer increases, the
migration amount of the ions also increases to bring up the
concentration degree of the liquid and thus the treating
capacity of the equipment. On the other hand, when the

107-4

-------
surplus current density is impressed, electrolysis of water

starts to occur. Further, due to the so-called Donan
equilibrium phenomenon, the concentration of S042" and Ca2+ ions
within the membranes starts to increase to cause gypsum scale
formation. Accordingly, the current density is decided from
the view point of preventing these phenomena.

In the said pilot plant, operation of the electrodialyzer with
high current density has been made possible while to prevent
gypsum scale generation a periodical membrane washing process
has been adopted.

Operation conditions for these were decided beforehand with
laboratory scale test facility and applied to the
demonstration test plant. Summary of these conditions are
stated below.

a.	Membrane current density	s 0.8 A/dm2

b.	Cl concentration of FGD effluent	: 20,000 ppm

c.	Cl concentration of diluted liquid	: 16,500 ppm

d.	Cl concentration of concentrated liquid	: 70,000 ppm

Secondary Concentration Facility

The liquid concentrated in the electrodialyzer is sent to a
secondary concentrating facility for further concentration.
For secondary concentration evaporative concentration is
carried out with a wiped film evaporator using steam as the
heating source. In the evaporator, the supplied liquid drops
from the evaporator top along the heat transfer wall surface
and gradually gets evaporated and concentrated while reaching
the evaporator bottom. During evaporative concentration, the
problem likely to occur is scaling of the heat transfer wall
surface. A means to avoid this is to prevent generation of
gypsum in the concentrated liquid by reducing the afore-stated
solubility product of gypsum in liquid. In combination with
this, low pressure evaporation is carried out employing low
temperature and pressure steam to prevent scaling due to
localized overconcentration of liquid on the heat transfer
wall surface.

Cl concentration for the secondarily concentrated liquid was
set at 250,000 ppm. To control this concentration, the
phenomena of boiling point elevation in conjunction with Cl
concentration in liquid was adopted.

The falling liquid temperature on the evaporator bottom
surface was regulated at a constant to obtain a certain Cl
concentration.

Solidification Facility

The liquid secondarily concentrated with evaporator is mixed
and kneaded with cement and flyash and disposed off as solid.
As mixer-kneader, a vibro mixer was employed.

107-5

-------
Results of Verification Test

Following items were confirmed with the verification test.

Filtration Equipment

The external element surfaces of this filtration equipment are
pearlite coated to facilitate collection and removal of waste
water contaminants. With the collection of contaminants, that
isr with the accumnlated SS loading, there is pressure loss
increase in the filter with time. The filter is therefore
periodically back washed to let the pearlite layer regain its
full capacity. To confirm performance of this filtration
equipment, together with the SS removal capacity it is
important to grasp the correlation of SS loading and pressure
loss.

SS concentration at the respective filter outlets over the 9-
month operation period met with the planned values and were as
shown below.

1st stage filter outlet < 1 ppm
2nd stage filter outlet < 0.2 ppm

During the test period the SS removal percentage has been
always 99.6% which thus met the planned value of 99%.

Electrodialyzer Tests

While carrying out concentration of gypsum saturated FGD
system waste water, following are important:

•	Establishing feasible membrane electric current density
and membrane washing method from the point of scale
prevention on the dialyzer membrane

•	Comprehension of concentration performance change with
time

•	Comprehension of permeability of the ions contained in the
liquid.

In the following explanation xs made about the results
obtained from the pilot test facility and fundamental
research.

Setting of electric current density. A current density of
0. 8A/dm2 was adjudged to be appropriate but as gypsum scaling
was apprehended even with this value periodic washing of the
dialyzer membranes was ascertained to be a necessity. In
Figure 6, CI concentration of the concentrated liquid with
change in electric current density is shown. Both in the
elementary research and pilot plant tests, the CI
concentration of the concentrated liquid with an electric
current density of 0.8A/dmz was approx. 70,000 ppm. Moreover,
within the range of these conditions, the CI concentration of
the concentrated liquid was not affected by diluted liquid

107-6

-------
flow velocity of 4.3 to 6.0 cm/s through, the membrane surface.
With membrane washing, gypsum scale causing Ca and S04 ions
collected within the membranes were discharged outside of the
membranes.

Change of concentration with time. Since the dissolved salts
in this FGD system waste water were mainly calcium chloride,
magnesium chloride and sodium chloride, the electric current
flow efficiency tin based on the permeability of CI ions
through the membranes as standard was used as a measure of the
concentration capacity. Electric current flow efficiency
was defined as follows:

1 CI

_ Qc.Cc.F ^ (io~3)

I.A.MW,

cl

[%I

Where?
Ici 1

Qc
Cc

F
I
A

Non-

standard electric current flow

efficiency of chlorine ions	[%]

j Flow amount of concentrated liquid	[kg/hr]
: Chlorine ion concentration in

concentrated liquid	[ppm]

Faraday constant	[A.hr/Eq]

Electric current density	[A/dm2]

Membrane area	[m2]

Chlorine ion molecules	[g/Eq]

Change of electric current flow efficiency ijCI with time is
shown in Figure 1.

As can be seen from the Figure 7, the electric current flow
efficiency remains practically same as that at the initial
operation stage even after a lapse of 5,400 hours without any
deterioration.

Permeabiflity of each type of ion. As stated before, the main
ions contained in waste water are: Ca2+, Mg2+ and Na* as
cations; and Cl" and S042" as anions. The permeability of these
ions through the dialyzer membranes depends on the particular
characteristics of the ion concerned. This characteristics of
an ion when expressed in a determinate quantity after being
compared with that of the Cl" ions is called the selective
permeability coefficient of the ion and is defined as follows:

rrt 1 _ ^"Cl

Tcl	—

'"Cl

Where;

1%! : Selective permeability coefficient of i component

[	]

Ci : Concentration of i component	[ppm]

Ccl : Concentration of chlorine ions	[ppm]

107-7

-------
Suffixes C and D represent concentrated liquid and diluted
liquid respectively.

Selective permeability coefficient of various components with
change in electric current density is shown in Figure 8. It
can be understood from this figure that the respective
selective permeability coefficient for the cations Ca2%
Mg2+, and Na+ is not affected by the electric current density
and remains almost constant. It also became clear from this
figure that the ion having the largest selective permeability
coefficient, that is, the ion which can permeate most easily
through the membranes among these ions is the Ca2+ ion while as
for the Mgz+ and Na2+ ions, their permeability is almost same as
each other.

With regard to the sulfuric acid SO*"2 ions, the object of
using these membranes in this equipment are:

*	to prevent gypsum scaling of the downstream evaporator

•	to prevent gypsum scaling of the dialyzer membrane itself

Since these membranes are of the univalent anion selective
type, their selective permeability coefficient was small
varying between 0.017 to 0.040.

Secondary Concentration with Evaporator

In this research a wiped film evaporator was used for
secondary concentration of the waste water as the purpose of
this evaporator was to further lower with concentration the
amount of liquid primarily concentrated with the
electrodialyzer in order that the quantity of auxiliary agent
used in the fmal solidification process can be reduced•

In a wiped film evaporator, chlorine concentration in its
concentrated outlet liquid is controlled by regulating the
pressure reduction degree and the concentrated liquid
temperature in side the evaporator but as a change in chlorine
concentration also largely affects the boiling point and
viscosity of the concentrated outlet liquid, it is necessary
to know the correlation of chlorine concentration change and
the change in boiling point and viscosity with this control.
Further, as it is also considered a possibility that with
scale generation there will be deterioration in the heat
transfer capacity of the evaporator with time, a discussion
regarding the change in heat transfer capacity of the
evaporator with continuous operation and physical properties
of the concentrated liquid is made in the following.

Change in heat transfer capacity with time. The change in
overall heat transfer coefficient of the evaporator with time
is shown in Figure 9. Overall heat transfer coefficient Ut was
calculated with the equation below:

Where;

Ut : Overall heat transfer coefficient [W/m2.K]

107-8

-------
Qh ; Overall heat transfer amount	[J/hr]

ATln : Logarthimic mean temperature difference [°C]
A i Heat transfer area	[m2]

Overall heat transfer amount was found from the product of the
condensed water quantity in the condenser used for condensing
the generated steam from the liquid subjected to secondary
concentration and the total of sensible heat and latent heat
found from the evaporation temperature. On the other hand,
the logarthimic mean temperature difference was found from the
temperature of the concentrated liquid and the heating steam.

As can be seen in Figure 9, the overall heat transfer
coefficient was almost constant at 872 [W/mz.K] throughout the
5,000 hours of operation period without showing any large
change with time.

Further, during the inspection of heat transfer surface on
completion of the operation, no scaling on it could be found
which confirmed the long-term stability of the concentration
process of this research carried out with the combination of
electrodialyzer and the wiped film evaporator.

Physical properties of concentrated liquid. During the
continuous operation shown in Figure 9, the liquid

concentration conditions were so adjusted as to let the
chlorine concentration of the concentrated liquid become
between 250,000 to 280,000 ppm as CI. Incidentally, it must
be mentioned that with higher chlorine concentration as the
physical properties of the liquid also change largely proper
attention must be paid in the selection of each type of
pITOCOSS equipment.

In the following mention is made about the changes in the
principal physical properties of the liquid.

(1)	Increase in boiling point

Along with the increase in chlorine concentration, the rise in
boiling point becomes larger; for example, at chlorine
concentration of 250,000 [ppm as CI] it becomes about 16°C (at
water vapor partial pressure Pw=10,133Pa). This means, as the
value of concentrated chlorine concentration becomes higher,
evaporation becomes difficult.

(2)	Specific gravity

Similar to the rise in boiling point, the specific gravity of
the liquid also increases with the increase in chlorine
content of the concentrated liquid. At a concentration of
250,000 ppm as CI, the specific gravity gets to be about 1.4.

(3)	Viscosity

107-9

-------
Viscosity also increases with the increase in concentration of
the concentrated liquid. It is characterized by the fact that
it increases rapidly particularly in the high concentration
region.

Solidification Facility

Purpose of solidification facility is to mix and knead
solidifying agents with the concentrated liquid from the
evaporator and produce a disposable solid. For the production
of an acceptable disposable solid, there are the following
requisites s

•	appropriate selection of solidifying agent mixing
proportions

•	easy handleability of the solidified product

•	appropriate leaching characteristics and strength of the
solidified product to meet environmental regulations

Solidified product obtained from the pilot plant facility is
shown in Figure 10.

Mixing proportions. As solidifying agent , a powdery mixture
of flyash and cement was used. In the selection of mixing
proportion of cement and flyash although good handleability of
the produced solid is an important consideration but in order
to reduce the operation cost it is necessary to decrease the
proportion of cement and replace it with flyash. However, as
the property and composition of flyash change largely
depending on the kind of coal and combustion conditions,
fundamentally it is necessary to select the mixing proportions
on the basis of the kind of coal used. As an example, for the
flyash recovered from the Hekinan Thermal Power Station, the
appropriate mixing proportion for it with concentrated liquid
ana csiDsiit# wss ss xoxxows •

Concentrated liquid s Cement : Flyash = 1.0 : 0.67 : 1.1

Handleability of the produced solids. For satisfactory
handling during the conveying process, the produced solids
should be free of fluidity and should not get bonded with each
other.

Leachability and compressive strength of the toxic components.

In order to be able to handle the produced solxds as
industrial waste, it is necessary to meet the standard values
prescribed in the Prime Minister's Office Ordinance (*1).

The solids produced from this pilot test facility met the
standard values both in respect of strength and leachability.
(*1) Prime Minister's Office Ordinance No. 5

[Ordinance issued by the Prime Minister's Office_to prescribe

107-10

-------
adjudication s tandards for industrial. wastes containing metals

etc. ]

Capital Cost and Operation Cost

In order to evaluate economic aspect of the Wastewater
Concentration and Solidification (WCS) system, investment
costs and operation costs of the system for a coal fired power
generating plant having two boilers of 1,000 MW capacity were
estimated together with the conventional waste water treatment
system under the same process design criteria and economic
criteria.

For the scope of complete wastewater treatment system, the
investment cost and the operation costs (total cost of
utilities, operation and maintenance, finance and management)
of WCS system are found to be 50 while these of the
conventional system are 100.

Conclusion

A waste water solidification system was newly developed for
the desulfurization (FGD) plants to replace the existing waste
water treatment systems. In this system, the waste water from
an FGD system after being concentrated by means of
electrodialysis and evaporation is solidified by mixing with
flyash and cement to produce an environmental pollution-free
disposable product for land reclamation use. In comparison
with the present waste water treatment systems, the newly
developed system is simple, requires less equipment space and
more economical from the point of construction and operation
costs. From the results of the present research it is firmly
believed that this newly developed waste water solidification
system will make large contribution as one of the
technological environmenta1 protection measures for the
thermal power stations hereafter.

107-11

-------
Year

1990

1991

1992

1993

Elementary research









Pilot facility design
and installation









Continuous
demonstration test









Figure 1 Development Milestone

Figure 2 Overall View of the Pilot Facility

NaCi Solution Tank

¦vasal

NaCt

Vaetmn*
Pump I

nl

~d

R«« Liijwi
Rrtidut! Filtrate.

Cake

Evgpsratar

A

- Condenser
4—|

Condensed Water
/Tobe recycled for use *s \
^FGD process cleaning water /

Primary
FiltnrteTank

Diluted liquid Tank Primarily Concentrated
U
-------
Figure 4 Electrodialyzer

:< A K A K A KA

Cathode

Na +













Ca + +













CI- —7



ci-



~ • «



CI-



Jf

SO4-









D

C

D



D

c

S04

7

D

Na +
Ca + ¦

Anode

1 )

I



I



,

,













A :
K :

C :
D :

Anion exchange membrane

K :

C :

D :

Diluted liquid	Concentrated

Liquid

Figure 5 Schematical Construction of Electrodialyzer

107-13

-------
100,000

CI concentration
of C liquid
[ppm]

50,000

Ccc« = 117201 + 65910 tppm]

I: Electric current density
(Data in between the current density of 1 to 3.2 [A/dm2] are
approximated with the least square method)

Measured values in the pilot facility

Pilot facility

Item

Other conditions

Cl concentration of 0 liquid {ppm]

10,800-18,400

Flow velocity of D liquid through
membrane chambers [m/sl

4.3 ~ 6.0

0.5 1.0 1.5	2.0 2.5 3.0

Electric Current Density [A/dm2]

3.5

Figure 6 Cl Concentration of Concentrate with Change
in Electric Current Density

100

Standard electric
current efficiency
of chlorine ions
[%]

~—e	©	®	—32	e			

[Concentration conditions]

Cl concentration of D
liquid [ppm]

13,610-18,400

Temperature of 0 liquid^

40 + 1

Flow velocity of D liquid
through membrane
chambers [cm/s]

6.0

J	1	1	1	L

°0	1000	2000	3000	4000	5000

Operation time Jhr]

Figure 7 Standard Electric Current Efficiency of Cl Ions

107-14

-------
1.0 -

Selective
permeability
Coefficient
Corresponding
to Ci

0.5 •

CD



Key

Component

O

Ca

~

Mg

A

Na

Mg • Na

&
D

O

Symbols with vertical line
indicate values measured
in the pilot facility



CI concentration of 0 liquid
Ippml

10.800- (8.40C

Temperature oi 0 liquid

rci

40 ±1

Flow velocity of 0 liquid
through membrane chambers
[cm/sl

4.3 ~ 5.0

0	'	1	1	'	'	1	1	

0	0.5 1.0	1.5 2.0 2.S	3.0	3.5

Electric current density [A/dm2]

Figure 8 Selective Permeability Coefficient of Components

Overall heat

transfer
coefficient
iW/ma-K]

1000 -

500

Item

Set Value

Pressure drop degree

[mm Hg abs]

100~120

Temperature [®C]

57 ™ 60

Processing amount [kg/hr]

70

CI concentration of concentrated
liquid |ppm as CI]

250,000-280,000

1000

2000

3000

4000

Operation time [hr]

5000

Figure 9 Overall Heat Transfer Coefficient of Evaporator

107-15

-------
7 8 9 023 123* 56789 TOl t 2 3 4 5 6 7 S S j5

Figure 10 Solidified Product

107-16

-------
FGD WASTE COMPOSITIONS AS LANDFILL LINER

C.L. Smith
Conversion Systems, Inc.

200 Welsh Road
Horsham, Pennsylvania 19044

Abstract

Coal burning utilities are and will be generating more flue gas desulphurization (FGD)
wastes at a time when many states are developing more stringent solid waste
regulations. In more and more cases, generators of these wastes are being required
to use some type of low permeable liner to prevent migration of waste into the
environment and degradation of groundwater. These liners, either clay or a
geomembrane, can be inordinately expensive.

Current developments in sulfo-pozzolanic chemistry offers the potential to process the
FGD wastes into a landfill liner, meeting presently promulgated environmental
regulations and providing the same protection from contamination to the environment
as other liners. These compositions contain sufficient compressive strengths and
flexural strengths which adequately compensate for the plasticity exhibited by clay and
geomembrane liners. FGD liners also exhibit a self-healing property, known as
autogenous healing, which seals stress induced fractures similar to clays.

Introduction

The Clean Air Act Amendments (CAAA)of 1990 will generate millions of tons per year
of FGD sludge, primarily calcium sulfate dihydrate and calcium sulfite hemihydrate, as
some utilities will install scrubbers to obtain compliance. Unless a market materializes
to reuse all of the FGD sludge and flyash (collected by the particulate removal systems
on utility stacks), a significant percentage of these waste products will require
disposal -in a solid waste facility. Furthermore, as many states are re-examining their
solid waste regulations to comply with and implement EPA's Subtitle D landfill
requirement, regulations at the state level are becoming more stringent than were
previously in place. In many cases, liners, composed of earthen material,
geomembranes or a combination thereof, are required in the design of a solid waste
disposal facility handling exclusively coal combustion by-products.

108-1

-------
These liner systems serve as a low permeable barrier to protect the ground-water from
contamination caused by the ieachate of the FGD wastes. The capital cost of these
liner systems will cost the utility millions of dollars. Certainly, a liner system that is
more cost effective yet provides equal protection to the environment should be
welcomed by both the utility industry and the regulators.

Regulatory Perspective

The € AAA requires utilities to reduce sulphur oxide emissions, in phases, to eventually
1.25 pounds per million of sulfur oxides (S02) British Thermal Units (BTU's) by the
year 2000, Phase I of the CAAA requires 110 coal burning utilities to obtain an
emission level of 2.5 pounds of S02 per million BTU's by the year 1995 (for the
purposes of this discussion, the extension provisions of the CAAA will not be
discussed since this provision allows the compliance deadline to be extended by two
years, however, it only shifts the starting waste generation date by two years without
affecting the disposal requirements). Utilities that choose stack scrubbing as their
compliance option will be faced with the disposal problem of a sulfate/sulfite filtercake
stream.

The Resource Conservation and Recovery Act (RCRA) has excluded coal combustion
by-products from regulation as a hazardous waste pending the results of studies and
final published rulemaking on the part of the United States Environmental Protection
Agency (EPA). EPA's final ruling is anticipated by August 1993. Unless and until EPA
publishes final regulatory requirements that drastically differ from present practices,
most coal combustion electrical generating facilities will be required to comply with
current sate solid waste disposal regulations.

EPA has also published regulations regarding disposal of municipal solid wastes (40
cfr 258). These regulations, commonly referred to as Subtitle D, require states to
adopt EPA technical requirements by October 9, 1993. (Note: There is a possibility
that EPA will extend the compliance deadline until April 1994).

Many states have extended the basic subtitle D regulations to the industrial sector.

Table I provides a limited selection of some relevant state regulations along with EPA's
Subtitle D requirements.

108-2

-------
TABLE I

SELECTED STATE REQUIREMENTS FOR LININGS16'



California |

Florida |

Illinois

Pennsylvania |

Virginia

Primary

* Landfill underlain

* Composite liner:

• Compacted

• Subbase (6" soil

• Compacted

Design

by natural

60-mil

earth

K<1X105)

clay

Liner

geologic

geomembrane

• 5 feet clay with

• 50 mil

• 2 feet clay with

System

material with

over 3 feet clay,

K <.1X107

geomembrane or

K <.1X107



K <.1X10 6

K <.1X107



1/4" clay mat





• Sufficient

• 24-inch-thick



• Plus 1 ft. clay





thickness to

protective layer



(K<_lX10"e)





prevent leachate

leachate



• Soil cover 18"





movement to

collection



(K> 1X102)





groundwater









Acceptable

• 2 feet of clay

• Double liner,

• Composite

• (Options for

* 60-mil

Alternative

with K <.1X10 6

Lower 60-mil

lining 60-mil

alternative designs

geomembrane

Design



geomembrane

geomembrane,

are included in the

over prepared

Liner



on 6" subbase

3 feet clay with

primary design)

base, or

System



with KXlXlO5

K <.1X107 or



• Other





• Leachate

• Other as



augemented





collection

approved



clay or soils





between liners;





manipulated in





Kj>lX10'e





place; or





* Upper 60-mil





• Double liner,





geomembrane





leachate





• 24" protective





collection





layer, leachate











collection







-------
Waste Quantities Arid Disposal Requirements

To place the potential effect of the impending regulatory changes in perspective, we
should review the volume of waste under consideration. There are approximately 450
coal burning power plants in the U.S. There are approximately 618 operating disposal
sites serving these utilities, while an additional 120 sites will be required by the end of
the decade. To complete the picture, there are approximately 759 sites already
closed.

The breakdown of the 618 operating sites are given in Table II.

TABLE IT"

Surface Impoundment

321

Landfill

273

Minefill

22

Waste Pile

2

Current coal combustion by-product generation is on the order of 90 million tons per
year. As part of Phase I of the CAAA, approximately 17 utilities have committed to
adding scrubbers to control S02 emissions. This will add approximately 4 million tons
per year of sludge to the existing disposal sites. Many utilities have begun planning
for Phase II compliance. While it is unknown as to how many additional scrubbers will
be installed, utilities will require disposal options, such as landfills.

Liners

A liner is a low permeable membrane which minimizes the amount of uncontrolled
leachate available to the environment. Ultimately, the liner, regardless of its
construction, will allow minimal amounts of leachate to enter the sub-soil and
groundwater. It is the job of the landfill designer to incorporate risk based factors into
the landfill design, and to minimize the exposure a receptor will receive at some
designated compliance point hydrologicaily downgradient from the landfill. Some of
the factors the landfill designer must incorporate into the design include the following:

1)	Compatibility of the liner with the materials being disposed of in the liner;

2)	The hydraulic conductivity of the liner;

3)	The ability of the liner not to crack (span sub-surface voids, settling, etc);

4)	The ability of the liner not to contribute to the quality or quantity of the leachate
generated;

108-4

-------
5) Risk based factors, such as distance to receptors, that will ensure that

groundwater will not exceed Maximum Contaminant Levels or Drinking Water
Standards (whichever is required at a particular site) at a compliance point
designated by the appropriate regulatory agency. Contaminate transport time
to the receptor may be a factor to consider in this subcategory.

Liners are composed primarily of either earthen material (such as clays) or
geomembranes (such as high density polyethylene (HDPE). Typical clay liners are
constructed by importing a sufficient quantity of material to properly compact and
construct a liner at least 5 feet thick with a hydraulic conductivity of
1 x 10'7 cm/sec or less over the entire area to be used for landfilling. If the utility is
fortunate to have in-situ clay, the cost to construct the liner is minimized. More
common is the practice of obtaining clay from a nearby source which requires
obtaining mineral rights to a property, mining the vein, transporting the clay to the site
and constructing the liner. The cost for this activity varies depending on the local
geological conditions, property values and landfill size, and reasonably has a range
between 5 and 20 million dollars for a typical disposal site.

A geomembrane only liner is usually between 30 to 60 mil thick and may be either a
single or double layer depending upon state requirements. These liners are effective
in preventing the passage of leachate, however, they are susceptible to puncture and
seam leakage and require expert installation to ensure proper construction. These
liner systems may be more cost effective given local availability of sources of clay and
the price to mine and transport the clay.

Composite liner systems are composed of a 30 mil layer of a geomembrane and
between 2 to 5 feet of low permeable earthen materials. In addition to the
construction concerns for the individual material, it is imperative that good contact be
made between the geomembrane and the earthen portions of the liner. Poor contact
between the two can allow for increased lateral flow from any hole in the
geomembrane increasing the leakage rate from the landfill as a whole.

FGD Liners

Liners can be manufactured by properly blending FGD scrubber sludge, flyash and
other additive(s). The blend initially results in a damp soil consistency near optimum
density eventually curing into a cementitious mass. Placement of the material is
comparable with clays. After placement, compaction is performed to allow the
pozzolanic and sulfo-pozzolanic reactions to take place bringing about a reduction in
the hydraulic conductivity of the blend. Formulations in the laboratory have
demonstrated permeability coefficients in the ranges of 10'7 to 10'8 cm/sec.
Permeability values of the fully cured stabilized FGD blends can reach into the 10"9
cm/sec range. The cost of these liners tends to be a function of the desired

108-5

-------
permeability. These liners will cost less than their clay or geomembrane equivalent
since approximately 95% of the liner will be constructed from waste materials being
generated by the utility.

As the liner is constructed, the nature of the pozzolanic bonding allows for increasing
the thickness without compromising the liner's integrity due to cold joints. Pozzolanic
reactions continue to take place over a long time. The placement of new liner grade
material on top of already placed liner grade material will result in vertical as well as
horizontal chemical bonding. The final liner effectively behaves as a monolith.

Furthermore, the landfill designer can vary the thickness of the liner providing for
equivalent to greater protection (as compared with common lining materials) of the
groundwater. The greater the thickness, the greater the vertical barrier preventing
percolation into the groundwater, the greater the amount of time it will take a
contaminant to reach a potential receptor.

Another beneficial property of stabilized FGD materials is the eventual development of
the liner into a cementitious mass with an unconfined compressive strength in the 500
to 1100 psi range. The corresponding tensile strength values are approximately 16%
of the compressive strengths or 80-175 psi. These strengths serve to bridge sub-
surface voids and support the weight of landfilled material (including heavy
equipment).

Cracking Of FGD Liners

Since stabilized FGD wastes will form a cementitious product, a brief discussion of the
potential for the stabilized FGD liner to crack thus providing a pathway for leachate to
enter the groundwater is appropriate. A close examination of the mechanism of
concrete cracking and comparison to the properties of the stabilized wastes should
satisfy the concerns of the designer that the liner will not crack under normal
operating conditions.

Concrete cracks for a variety of reasons, including shrinkage from moisture loss and
temperature changes of the mass. Both of these factors cause significant stresses
within the quick curing (28 days) concrete mass. The inability of the concrete mass to
handle the internal stresses causes the external surface to crack to relieve the
stress/strain buildup.

Stabilized- FGD materials are placed as part of the liner at a high percent solids
(approximately 75% or more) unlike concrete which is typically wet placed (poured).
While some of the residual moisture is continually consumed in the slow curing
pozzolanic reaction, very little water is lost to evaporation. Furthermore, since the
stabilized FGD wastes results in a mildly expansive reaction, similar to shrinkage
compensated concrete, cracking caused by shrinkage does not occur. Stresses

108-6

-------
normally caused by temperature differentials do not occur because the slow curing of
the stabilized FGD mass allows the liner to absorb the stresses.

The above is confirmed with field data from Conversion Systems, Inc. (CSI) operation
of the landfill at the Seminole Electric Cooperative, Inc. in Palatka, Florida. (Figure 2)
CSI, as part of the operating contract, regularly cored the landfill and prepared reports
on the condition of the landfill. Analysis of the cores showed little decline in the
moisture content of the cores over time. The cores were also visually examined for
cracks and none were recorded.

Performance of Stabilized FGD Waste Compositions

From 1985 to 1989, four water storage ponds at the Palatka facility that were
previously lined with plastic were found to be leaking; all four were replaced with a five
foot thick lining of the FGD waste composition. The four ponds, each of which
averages about 1/4 acre area, have performed well to date. This is a more rigorous
test as compared to a landfill lining, in that hydraulic gradient builds up slowly via
rainfall in a typical landfill, while the hydraulic gradient was applied to the ponds within
a very few weeks. Figures 3, 4, and 5 show two of these ponds.

After 6 months, cores were removed from segments of the pond lining, in a manner
not compromising the integrity of the linings. The average permeability coefficient
values (after 6 months of cure) was 1.8xi0'8 cm/sec. Unconfined compressive
strength values averaged 679 psi. Hence, the performance and durability of FGD
composition liners have been demonstrated.

Fissures

Small fissures which may develop are corrected by the autogenous healing
characteristic of sulfo-pozzolanic materials. Due to differential settling or uneven
loading in the initial months, some fissures may develop. The autogenous healing
characteristics of the sulfo-pozzolanic materials are capable of bonding the segments
and restoring the hydraulic conductivity initially developed while increasing the
strength of the liner mass.

CSI empirically evaluated the autogenous healing characteristic by preparing a wide
range of stabilized FGD compositions at optimum moisture and density. These
Proctor cylinders were cured at 100% relative humidity for 100°F for seven days.

These cylinders were fractured and then recombined and returned for curing. The
data in Figure 1 shows that the sulfo-pozzolanic chemistry effectively regenerates
strength significantly beyond the original cylinders tested. Table III provides the same
data as in Figure 1, but includes specific strength values and provides the number of
data points used in obtaining the averages shown in Figure 1.

108-7

-------
AUTOGENOUS HEALING

o
<2°

oo

2 3 4 5 6 ?

10 11 12

EQUIVALENT FIELD CURE (Months)

FIGURE 1: Strength Regeneration in FGD Compositions

-------
TABLEffl

FIELD
EQUIVALENT
CURE
(Months)

CURE AFTER
"BREAK"
(Months)

UNCONFINED
COMPRESSIVE
STRENGTH
(AV PSD

DATA POINTS
IN

AVERAGE

1

0

453

43

1

1

949

42

3

0

1042

60

3

3

1195

60

6

0

1151

39

6

6

1271

39

Environmental Impact Of FGD Liners

EPA, as well as many sates, evaluate the potential hazards of a landfill through the
use of fate and transport models. These models will calculate potential exposure to
hazardous chemicals released from a landfill. The use of these predictive techniques
is appropriate in evaluating the environmental impact of an FGD liner system for utility
combustion wastes. Using EPRI's Fossil Fuel Combustion Waste Leaching (FOWL)
model in conjunction with EPA's Hydroiogic Evaluation of Landfill Performance (HELP)
and Composite Model for Landfills (CML), a thorough evaluation of the site design can
be performed demonstrating to the regulators that the final site design will have a non-
adverse environmental impact on a potential receptor.

Conclusions

Stabilized FGD materials can be successfully utilized as a liner for a landfill of coal
combustion by-products. Key design parameters, such as hydraulic conductivity and
thickness, can be varied to provide the utility a cost effective design without an
adverse environmental impact. Design adjustments can be driven by modelling results
thus assuring a potential receptor is not exposed to hazardous materials above the
appropriate state standard. The sulfo-pozzoianic chemical reactions taking place
within the chemical matrix will continue to lower the permeability over time while
providing sufficient compressive and flexural strengths to avoid breakage of the liner
from spanning sub-surface voids and differential sub-surface settling.

108-9

-------
Reference

1.	"Utilities May have To Clean Up Old Coal Waste Sites", Power Engineering,

May 1993, p. 10

2.	R.G. Knight and E.H Rothfuss. FGD By-Product Disposal Manual. Palo Alto,
CA: EPRI. 1993, pp 7-56-7-61.

3.	Hairy M. Freeman (editor). Standard Handbook of Hazardous Waste Treatment
and Disposal. New York, New York: McGraw - Hill 1991, pp. 10.68 - 10.70.

4.	Robert A. Corbitt. Standard Handbook of Environmental Engineering. New
York, New York: McGraw-Hill 1990, pp. 8.123-8.125

5.	R.H. Cook and D.C. Kocunik, "Examination of State Regulations for Dry
Disposal Facilities for Coal-Fired Plant Wastes," Paper No. 54-1, presented at
the American Power Conference, Chicago, IL(April,1993)

6.	C.L. Smith, "Physical Aspects of FGD By-products" International Journal of
Environmental Issues in Minerals and Energy Industry Vol 1 No. 1, p. 37(1992).

7.	Code of Federal Regulations, Title 40, Part 260

8.	Code of Federal Regulations, Title 40, Part 258

9.	J, P. Giroud and R. Bonapack, "Leakage Through Liners Constructed with
Geomembranes," Part I: Geomembrane Liners, Geotextiles and
Geomembranes, 8 (1) (1989) 27-67.

10.	Foundation Engineering Handbook; edited by Hans Winkerton and Hsai-
Yang Fong; Van Nostrand Reinhold Company; 1975.

11.	M.Anderson; G. Sharpe; D.Allen; H. Southgate; and R. Dean; "Laboratory
Evaluations of Stabilized Flue Gas Desulfurization Sludge and Aggregate
Mixtures;" presented at the Seventh International Ash Utilization and Exposition;
Orlando, FL; 1985.

12.	R. Smock; "UtilitiesMay Have to Clean Up Old Coal Waste Sites," Power
Engineering, May 1993, Page 10.

108-10

-------
Figure 2: Florida Utility With 8+ Million Ton Monolithic FGD
Waste - Fly Ash-Lime Landfill

Figure 3; Water Storage Ponds Lined With FGD Waste-Fly Ash-Lime
Composition

108-11

-------
Figure 4: Closer View of Pond Lined With FGD Waste-Fly Ash-Lime
Composition

Figure 5; Comparable View of Another Fond

108-12

-------
TECHNICAL REPORT DATA

(Please read Instructions on the reverse before completi II11 III II 1|||| 111 III 1 1 1 111

1. REPORT NO. 2.

EPA-600/R-95-015d

3. f III HII || Hill 11mi ii 1111

PB95-179255

4. TITLE AND SUBTITLE

Proceedings: 1993 SOg Control Symposium, Volume
4. Sessions 7, 8A, and 8B

S. REPORT DATE

February 1995

6. PERFORMING ORGANIZATION CODE

7. AUTHOR(S)

Miscellaneous

8. PERFORMING ORGANIZATION REPORT NO.

TR-103289-VI, -V2, and -V3

(EPRI)

9, PERFORMING ORGANIZATION NAME AND ADDRESS

See Block 12

10, PROGRAM ELEMENT NO.

11. CONTRACT/GRANT NO.

NA (Inhouse)

12. SPONSORING AGENCY NAME AND ADDRESS

EPA, Office of Research and Development

Air and Energy Engineering Research Laboratory
Research Triangle Park, NC 27711

13. TYPE OF REPORT AND PERIOD COVERED

Final; 8/93-8/94

14. SPONSORING AGENCY CODE

EPA/600/13

is. supplementary notes j^eERL project officer is Brian K. Gullett, Mail Drop 4, 919/541-
1534.

i6. abstract repOI»fc documents more than 100 presentations at the 1993 SG2 Control
Symposium, in Boston, MA, August 24-27, 1993. The presentations covered a wide
range of topics: industry's strategies for dealing with the Clean Air Act Amendments
of 1990, including Phase I strategies, the emission allowance trading system, and
retrofit construction; additiv|.es, materials, and operating issues for wet flue gas
desulfurization (FGD); clean 'coal demonstration programs; the effect of FGD sys-
tems on air toxics; dry FGD technologies of spray drying and furnace sorb en t injec-
tion; applied sulfur dioxide (S02) control research results, and emerging acid rain
control technologies; and waste disposal issues. The presentations covered results
obtained from full-scale demonstration/operation to pilot and bench scale work.

\o

17. KEY WORDS AND DOCUMENT ANALYSIS

a. DESCRIPTORS

b.IDENTIFIERS/OPEN ENDED TERMS

c. COSATI Field/Group

Pollution Toxicity
Sulfur Dioxide Spray Drying
Additives Furnaces
Flue Gases Sorb en ts
Desulfurization Waste Disposal
Coal

Pollution Control

Stationary Sources
Acid Rain

13 B 06T
07b 13 H

11G 13 A
2 IB

07A.07D 15E
2 ID

18. distribution STATEMENT

Release to Public

19. SECURITY CLASS (ThisReport}

Unclassified

21. NO. OF PAGES

526

20. SECURITY CLASS (Thispage)

Unclassified

22. PRICE

EPA Form 2220-1 (9-73]

-------