s>EPA
United States
Environmental Protection
Agency
Industrial Environmental Research
Laboratory
Research Triangle Park NC 27711
Technology Transfer
Summary Report
| - '
i
Sulfur Oxides Control
Technology Series:
Flue Gas Desulfurization
i
i
Magnesium Oxide
Process
-------
-------
Technology Transfer ] EPA 625/8-81-005
I .
Summary Report
i
Sulfur Oxides Control
Technology Series:
Flue Gas Desulfurization
. i
Magnesium; Oxide
Process
April 1981
This report was developed by the
Industrial Environmental Research Laboratory
Research Triangle Park NC 27711
-------
1
Magnesium oxide storage silos
-------
I
Introduction
The magnesium oxide (MgO) flue
gas desulfurization (FGD) process
(Figure 1) is a sulfur dioxide (SO2)
recovery system that uses a recircu-
lating MgO slurry to remove SO2
from stack gas. The slurry reacts with
SO2 to form magnesium sulfite
(MgS03), which is then heated to
regenerate MgO. The concentrated
SO2 released during regeneration
can be converted to sulfuric acid
(H2S04) and other products.
Major advantages of the MgO FGD
process include ^the ability to:
Recover sulfite salts easily from
the slurry i
I
Regenerate the absorbent, MgO
Alleviate the problem of solids
disposal
An analysis of these benefits,
however, must consider the cost of
installing and operating the MgO
FGD process, or any other relatively
complex SO2 recovery system.
The U.S. Environmental Protection
Agency (EPA), working with a group
of chemical and utility companies,
funded two MgO FGD demonstra-
tion plants. The first plant was
installed on a 150-MWoil-fired boiler
at Boston Edison Company's Mystic
Station in Everett, Massachusetts;
Key
Flue gas/off-gas
Cleaned flue gas
Absorption liquor
_____ Sulfur products,
L , ^ Other systems
Desulfurized
flue gas
Power.
plant
Figure 1.
Magnesium Oxide
FGD Process Followed by SO2 Conversion
I
-------
Scrubber surge tank
the second was designed to de-
sulfurize one-half the flue gas
from a 190-MW coal-fired boiler at
Potomac Electric Power Company's
(Pepco's) Dickerson Generating
Station in Frederick, Maryland. The
magnesium sulfite formed in the
process at these installations was
regenerated, and the byproduct
SO2 stream was transported
to Rumford, Rhode Island, and con-
verted to sulfuric acid at Essex
Chemical Company's plant.
The MgO FGD process can be
evaluated objectively on the
basis of operational experience.
Boston Edison's installation
operated for 27 months from 1972
through 1975, whereas Pepco's
system operated for 15 months in
1974 and 1975. Results of tests
conducted at both plants on such
factors as design considerations,
cost, and environmental impact
have established the MgO FGD
process as a feasible means of SO2
emission control.
This report summarizes the MgO
FGD process and provides a basic
understanding of FGD technology.
-------
Process Description
The MgO FGDiprocess consists of
four major processing steps:
1. Flue gas pretreatment
2. SO2 absorption
3. Solids separation and drying
4. MgSO3 regeneration
Sulfur dioxide processing may be
considered a fifth step because it
is often associated with the
MgO FGD process.
Figure 2 illustrates the process flow
for a typical MgO FGD system.
In the first step, water scrubbing
cools and saturates the boiler flue
gas and removes fly ash and
chlorides upstream of the absorber.
As a rule, flue gas from oil-fired
boilers does not require pretreat-
ment, but this step is necessary
in coal-fired applications.
In the absorber (Step 2), S02 is
removed from the flue gas by contact
with a recircula;ting slurry of
MgO, MgS03, and magnesium
sulfate (MgS04). Flue gas SO2
diffuses into this slurry and reacts
with MgO to form MgSO3, some
of which reacts with oxygen present
in the flue gas to form MgS04.
Additional MgSO4 is formed when
flue gas sulfur trioxide (S03) reacts
with MgO.
I
Desulfurized flufe gas leaves the
absorber, is reheated if necessary,
and is exhausted through the stack.
The sulfurized scrubbing liquor
flows to a sump and is recycled to the
absorber after aj continuous bleed
stream has been withdrawn from
the recirculatiorj loop. Fresh MgO
added to the recirculation loop
replaces any magnesium removed
from the scrubbing liquor by the
bleed stream, j
In the third step, the bleed stream
is routed to a centrifuge for
processing into 60 percent solids by
weight, and the!mother liquor is
recycled to the absorber recirculation
loop. The strearh of 60 percent
solids by weight flows to a dryer
where surface moisture and most of
the water of hydration are removed
to produce a dry powder of MgS03,
MgS04/ unreacted MgO, and inerts.
Calcination of the dry powder
in the regeneration processing stage
(Step 4) converts MgSO3 and MgSO4
to MgO, which is recycled to the
absorber recirculation loop. Calcina-
tion also produces an S02-rich
byproduct stream that may be
processed further to form sulfuric
acid or elemental sulfur.
Thus, the MgO FGD process not only
regenerates the essential absorbent,
MgO, but also produces sulfur
dioxide at concentrations practical
for conversion to sulfuric acid or
elemental sulfur.
Pretreatment
Pretreatment of flue gas in oil-fired
systems is almost unnecessary
because fly ash levels are minimal.
Particles generated by such systems
are usually carbonaceous and are
consumed in the regeneration step,
where they serve as a reducing
agent for magnesium sulfate.
Unlike oil-fired systems, coal-fired
systems require flue gas pretreatment
to remove fly ash upstream of
the absorber and to prevent large-
scale contamination of the recir-
culating slurry used in the absorption
step. Pretreatment is necessary
for the following reasons:
Fly ash from coal contains com-
pounds of vanadium and iron,
which catalyze the undesirable
reaction of magnesium sulfite
with oxygen to form magnesium
sulfate.
Fly ash is separated more easily
from the water used in the
pretreatment step than from the
slurry mixture of the absorption
process. Without pretreatment,
the replacement of large quantities
of magnesia removed from the
slurry with the fly ash would
increase the operating cost of the
system significantly.
-------
I
Sulfuric acid
or
sulfur
Key
Flue gas/off-gas
llf'::"..< Cleaned flue gas
milllH Absorption liquor
I I Sulfur products
I ,. ...i , Other systems
Figure 2.
Magnesium Oxide FGD Process With Regeneration and SO2 Conversion
-------
Pretreatment cools and saturates
the flue gas and reduces the
concentrations of corrosive
chloride ions entering the
absorber. Typical flue gas tem-
peratures range from 290° to
310° F (145° to 155° C) before
pretreatment and from 125° to
135° F (50° to 55° C) following
pretreatment.1
Fly ash removal may occur in the
first stage of a double-stage venturi
scrubber (Figure 3). Flue gas enter-
ing the irrigated throat of the
first stage is sprayed with water.
The impact of the dispersed water
droplets removes particles from the
flue gas stream. Particle collection
efficiencies of 99 percent or
higher are achieved.
The ash-laden liquor from the
first stage is recycled to the venturi
converging section. A bleed
stream is continuously withdrawn
from the recycle line and routed to
the thickener, where the particles
settle. Thickener underflow is
pumped to an ash disposal pond,
and overflow is routed to a transfer
tank where makeup water is added
to the system. The liquor is pumped
back to the venturi first stage.
In a two-stage venturi scrubber,
flue gas passes through an annular
mist eliminator before entering
the second stage of the venturi. The
mist eliminator removes residual
ash-laden droplets small enough
to be carried with the gas.
Absorption
Venturi scrubbers were used at both
EPA demonstration plants. Flue
gas from an oil-fired boiler may be
routed to a single-stage venturi
for removal of sulfur dioxide. Coal-
fired installations may use a double-
stage venturi, which is capable
of pretreatment as well as absorption.
Single- and double-stage venturi
scrubbers are compared in the
section on Design Considerations.
In the S02 absorption and venting
operation (Figure 4a), flue gas
entering the venturi scrubber comes
into contact with the recirculating
MgO slurry. The'liquor is atomized
by the high-velopity gas, and
flue gas SO2 is absorbed rapidly into
the finely divided slurry. Desul-
furized flue gas then passes through
spray-washed mist eliminators and is
discharged through the stack.
i
Removal of SO2 jfrom flue gas
involves absorption and the following
chemical reactions:
Mg(OH)2 + S02 f*
MgSC-3 + H2O!
MgSO3 + H20 +|S02
Mg(HS03)2 I
Mg(HS03)2
2MgS03
MgO
H26
(1)
(2)
(3)
A 5-percent excess of MgO in the
slurry is necessary to ensure the
completion of the reaction in
Equation 3 and thus to guarantee
that the slurry leaving the absorber
contains no mag'nesium bisulfite
[Mg(HS03)2]. I
Additional reactions produce
magnesium sulfate:
MgS03 + 1/aO2 -* MgSO4 (4)
I
. MgO + SO3 -» MgS04 . (5)
Studies indicate jthat most of the
magnesium surface formed in the
scrubber results from sulfite
oxidation by excess oxygen in the
flue gas. Air oxidation of sulfite
is a light-activate'd, free radical reac-
tion catalyzed by| certain metallic
ions, including iron and vanadium,
and is inhibited by free radical
scavengers.2 Oxidation of sulfite
occurs at every stage in the process
until the salts are calcined to
regenerate MgO.j Surface oxidation
of magnesium su|lfite crystals can
become excessivje even when the
material is stored or shipped
off site for calcination. It is advan-
tageous to limit sulfite oxidation
throughout the process because a
higher temperature is required
for thermal decomposition of MgSO4
than for MgSO3. For this reason,
more energy is needed to regenerate
MgO from sulfate than from sulfite.2
Hydrated crystals of MgS03 and
MgS04 are formed in the. venturi
scrubber and sump. The slurry is
recirculated from the sump to
the spray nozzles in the scrubber.
A bleed stream is withdrawn
constantly from the absorber recir-
culation loop to maintain the
desired concentration of magnesium
solids in the recirculating slurry.
The bleed stream, which contains
MgSO3 and MgS04 crystals as well
as unreacted MgO and magnesium
hydroxide [Mg(OH)2], is routed to
a dewatering system.
Fresh magnesium hydroxide slurry
is added to the recirculation loop
to compensate for magnesium
withdrawn in the bleed stream. The
slurry is prepared by slaking
regenerated and makeup MgO in an
agitated, heated tank:
MgO + H20 - Mg(OH)2
(6)
The treated flue gas is demisted
to remove entrained liquid and solids
and may be reheated before it is
exhausted to the stack. Reheating
desaturates the gas, increases
buoyancy, and aids in dispersing any
remaining stack constituents.
Solids Separation and Drying
Figure 4b illustrates a typical solids
separation and drying system, which
requires a centrifuge, dryer, and
dust collector. The bleed stream from
the absorber recirculation loop is
routed to a centrifuge that con-
centrates the slurry from about 10
percent solids by weight to
approximately 60 percent solids by
weight. The mother liquor is recycled
to the absorber recirculation loop.
-------
Key
Flue gas/off-gas
Cleaned flue gas
Absorption liquor
Sulfur products
II ' J Other systems
Flue gas
SO2-rich calomer off-gas
to H2S04 plant
Figure 3.
Pretreatment Step in Magnesium Oxide FGD Process
-------
Key
Flue gas/off-gas
Cleaned flue gas
Absorption liquor
Sulfur products
I ''-'' J"3 Other systems
(a)
S02-nch calcmer off-gas
to H2S04 plant
Regenerated
MgO silo
(MgO .to FGD system)
Figure 4.
i
Magnesium Oxide FGD Process: (a) SO2 Absorption and Venting, (b) Solids Separation and Drying, and (c) Calcina-
tion/Regeneration
-------
From the centrifuge the wet solids
are conveyed to a dryer, typically
a rotary kiln, for exposure to hot
gases produced by combustion of
fuel oil or natural gas. This direct
firing removes surface moisture and
water of hydration from the centri-
fuged solids to produce dehydrated
MgS03, MgSO4, MgO, and inerts
(e.g., ash).
Off-gas from the dryer contains
entrained solids, which are removed
in a cyclone dust collector. The
collected dust and the dryer product
are stored in a silo for the regenera-
tion processing step. The cleaned
off-gas is routed to the venturi
scrubber where it mixes with
entering flue gas and is cleaned
of remaining particles.
Regeneration
The dryer product is calcined, de-
composing the magnesium salts
to regenerate MgO and liberate SO2.
Calcination temperatures are
adjusted within the range of 750° to
1,850° F (400° to 1,000° C) for
optimum reactivity and surface area
of the product MgO. Calcination
temperatures above 2,200° F
(1,200° C) will "hard burn" the
MgO, whereas temperatures above
2,900° F {1.600° C) will "dead burn"
the material. Magnesium oxide,
when it has been hard burned or,
dead burned, is unreactive and
useless in FGD systems.
Several studies performed on thermal
decomposition of magnesium sulfite
revealed that some decomposition
occurs even at temperatures as
low as 570° F (300° C). In the
range of 570° to 1.100° F (300° to
600° C), decomposition of MgS03
can yield a relatively high percentage
of MgS04. Even at temperatures
as high as 1,650° F (900° C), MgSO4
will form in unacceptable amounts
unless gas phase compositions
are controlled carefully. At tempera-
tures above 1,850° F (1,000° C),
MgO and SO2 are the primary
reaction products over a wide range
of gas phase compositions.
i
In a typical calcination/regeneration
operation (Figure 4c), dry magnesium
salts from the FGD system dryer
and cyclone are fedifrom an interim
storage silo to an oil-fired fluid
bed calciner. The calciner contains
a single calcination ibed designed
to operate at 1,400? to 1,600° F
(760° to 870° C).
Under ordinary conditions, calcina-
tion at this temperature produces
high levels of magnesium sulfate.
A reducing atmosphere may be
achieved by adding a proportionate
amount of carbon (coke) to the
calciner feed. In the calciner,
a reducing atmosphere prevents
MgSO4 production and ensures
decomposition of any magnesium
sulfate present in the calciner feed.
The following reactions occur
in the calciner:
C + y&z CO T (7)
CO + MgS04 -
CO2 + MgO + SO2 T . (8)
MgS03 -» MgO + SO2 T (9)
Off-gas from the calciner contains
about 10 percent sulfur dioxide.
The gas stream is partially cleaned of
particles in a cyclone and is routed to
a sulfuric acid production unit.
The collected particles join the
incoming calciner feed stream and
return to the calciner. A storage
silo holds the regenerated MgO for
reuse in the absorption system.
SO2 Processing
Calciner off-gas contains approxi-
mately 10 percent S02 by volume,
which is sufficient for the production
of dry, compressed sulfur dioxide,
sulfuric acid, elemental sulfur, and
fertilizer materials.'
Converting S02 to sulfuric acid
requires prior treatment of the
S02-rich gas from the calciner. Cool-
ing, drying, and cleaning take place
in a weak acid venturi scrubber
in series with a packed tower. The
venturi cools and cleans the gas; the
packed tower further cools the
gas by contact with an acid stream.
The latter step is necessary because
the gas is saturated with water
vapor when it leaves the venturi.
The packed tower reduces the gas
temperature to below 100° F
(38° C), causing water to condense
from the gas stream and producing
a dried gas that is suitable for
treatment in an acid plant.
Sulfuric acid production requires a
feed stream containing 8.4 to 9.0
percent O2 by volume. The feed
gas first passes through a drying
tower where it is dried by contact
with 93 percent H2S04. Any S02
absorbed by the acid is removed in
an acid-stripping tower and returned
to the main gas stream. A small
amount of air added to the gas with
the S02 serves to adjust the
oxygen concentration.
A blower draws the dried gas through
a series of heat exchangers to
raise its temperature to 815° F
(435° C). This step supplies the
heat necessary to sustain the sub-
sequent reactions. As the hot gas
enters a series of converter beds,
it reacts with a vanadium pentoxide
(V206) catalyst, and the S02 is
oxidized to form S03:
S02+1/202
V,
so, t
(10)
The hot, SO3-rich gas produced
by this exothermic reaction is routed
to the heat exchangers where it
preheats the acid plant feed stream.
The stream then enters the ab-
sorber and reacts with a weak
8
-------
Surge hopper and slaking tank
-------
Flue gas
Key
Flue gas/off-gas
Cleaned flue gas
HHHI Absorption liquor
HI: - :: I Sulfur products
!S1 Other systems
> Steam o
gas/air
S02-rich calciner off-gas
to HS0 plant
Figure 5.
General Magnesium Oxide FGD System
10
-------
H2SO4 solution to form a more
concentrated solution:
S03 + H20 H2S04 (11)
Typical SO2-to-S03 conversion
efficiencies for catalytic oxidation
range from 95 to 98 percent Plants Integrated System
with efficiencies in the lower end
of this range have high SO2 emis-
sions. Where regeneration and acid
production are conducted on site,
the acid plant tail gas can be
recycled to the |SO2 absorber to
achieve a higher conversion
efficiency.3 ,
The foregoing steps are part of the
integrated system. Figure 5 shows
the interrelationship of the
processes of a typical MgO FGD
system applied to a coal-fired boiler.
11
-------
Design Considerations
Although a complete discussion
of design considerations involved
in the construction iand operation of
an MgO FGD system is beyond the
scope of a summary report, the
following discussion contains
sufficient information to permit a
general understanding of the process.
Pretreatment and Absorption
Particle and SO2 removal efficiencies
are the major concerns in the
design and operation of pretreatment
and absorption equipment. The
venturi scrubber is preferred for
fly ash removal because it attains
high particle removal efficiency at.
minimum, cost.4 The absorber
used' to remove sulfur dioxide from
flue gas also may be of venturi
design, though other scrubber
systems, including mobile bed con-
tactors or spray towers, are
effective.2 Packed towers or tray
absorbers are not used often with
slurry systems because they tend
to become plugged.
The single-stage venturi scrubber
used at Boston Edjson's installation
(Figure 6a) was designed by
Chemical Construction Company
(Chemico), which is now known as
Chemico Air Pollution Control
Company. A slightly different
Chemico venturi Was used in
Pepco's coal-fired system (Figure 6b).
The Pepco scrubber had two
stages and was capable of flue
gas pretreatment as well as
absorption. ;
When flue gas enters a venturi
scrubber, it comes into contact with
a recirculating MgO slurry. At Boston
Edison's installation, MgO slurry was
introduced into the scrubber
at three points (Figure 6a):
Directly on the cone in the venturi
throat
On the annulus surrounding
the cone
Around the periphery of the
venturi throat
The absorbing slurry entered the
second stage of the Pepco venturi
through tangential spray nozzles
(Figure 6b). Both the Boston Edison
and Pepco systems thus provided
uniform slurry irrigation to the
converging surfaces of the scrubbers.
Important factors to consider in the
design and operation of any
scrubber used in the MgO FGD
process include such process
variables as liquid-to-gas (L/G) ratios,
pressure drop, and slurry pH.
Based on pilot and operating tests
with a venturi absorber, the optimum
L/G ratio for S02 absorption is
in the range of 20 to 40 gal/1,000
stdft3 (2.7 to 5.4 I/normal m3).
Operating the absorber at an L/G
ratio below 20 gal/1,000 stdft3 (2.7
I/normal m3) lowers the SO2
removal efficiency. In contrast,
operating with the L/G ratio above
40 gal/1,000 stdft3 (5.4 I/normal m3)
results in a higher pressure drop
across the absorber with marginal
improvement in S02 removal
efficiency.1
As the flue gas accelerates through
the venturi throat, a pressure
drop is established. The venturi
must be operated at a minimum
pressure drop to keep energy con-
sumption low. In coal-fired applica-
tions, a fixed pressure drop is
essential to maintain first-stage
efficiency. Maximum efficiency of
particle removal results from
specific droplet size distribution,
which is directly related to the flue
gas velocity. Second-stage efficiency,
however, can be maintained over
the range of pressure drops indicated
in the next paragraph. Gas absorp-
tion, unlike particle removal, requires
a constant specific surface area
of the dispersed liquid, which,
in the venturi, is self-regulating over
a range of gas flows.5
At both the Boston Edison and
Pepco installations, S02 removal
efficiencies of over 90 percent were
12
-------
attained with a pressure drop of
6 to 11 inches H2O (1.5 to 2.7 kPa).6
Particle removal efficiency at Pepco's
first-stage venturi was optimized
at a pressure drop of 11 inches
H2O (2.7 kPa). Experiments were
conducted at Pepco's installation
to test the particle removal efficiency
of a boiler electrostatic precipi-
tator (ESP) operating in conjunction
with the prescrubber. Particle
removal efficiency was greater than
99 percent when the flue gas was
taken directly from the boiler and
not preconditioned by the ESP.7
Tests conducted at Boston Edison's
Mystic Station indicate that sulfur
dioxide removal efficiency is a direct
function of the slurry pH.6 The
scrubbing liquor pH is controlled by
the addition of fresh MgO slurry
to the absorber and the concurrent
withdrawal of spent scrubbing liquor
from the recirculation loop. Fresh
MgO slurry raises the scrubbing
liquor pH and increases the SO2
removal efficiency of the absorber.
At basic pH values, the absorbent
solubility decreases, and higher
L/G ratios are necessary to maintain
removal efficiency.2 At Boston
Edison's installation, the recycle
slurry pH was maintained between
6.8 and 7.5; Pepco's recycle pH
was optimized at 7.0.
The MgO slaking operation
presented problems at both Boston
Edison's and Pepco's demonstration
plants. Initial designs for the MgO
FGD process specified an agitation
tank for MgO slaking but did not
account for differences in slaking
between regenerated and fresh MgO.
Because regenerated MgO was
not slaked as easily as fresh
MgO, problems developed in the
absorber that resulted in reduced
SO2 removal efficiencies.1 Modifying
the slaking equipment to incorporate
premix tanks and steam heaters
alleviated the slaking problem.
(a)
Flue gas
inlet
Venturi throat
I Clean gas
! to stack
Slurry inlet
A (cone wash)
Slurry inlet
(tangential wash)
Annulus
Key
(b)
Mitt eliminator Wff///////.
i
Flue gas/off-gas
Cleaned flue gas
Absorption liquor
It?" il Sulfur products
I' H bther systems
]
i Flue gas
! inlet
! Scrubbing water
! inlet
O
To slurry recycle system
Gas flow
Scrubbing water
inlet
Mist
eliminator
To scrubbing
water
recycle system
Dryer exhaust
gas inlet
Slurry inlet
Clean gas
to stack
Liquid level
To slurry recycle system
Gas flow
Figure 6. |
(a) One-Stage Vfenturi Scrubber at Boston Edison's Installation and (b) Two-
Stage Venturi Scrubber at Pepco's Installation
13
-------
Reheating
Reheating of desulfurized flue
gas may be necessary to prevent
condensation of water vapor as gas
is ejected from the stack to the
atmosphere. Reheating may be
accomplished by heat exchange with
high pressure steam. A typical
reheater'consists of a series of
shell-and-tube heat exchangers that
can raise the temperature of the gas
to 175° F (80° C). At the Boston
Edison and Pepco installations,
however, no reheating was provided
for the flue gas other than that
accomplished by mixing with hot,
untreated gas from the boiler.
Solids Separation and Drying
Off-gas from the dryer contains
entrained solids that are removed
in a cyclone dust collector. At
Boston Edison's installation, the
collected dust and the dryer product
were transferred separately to a
silo for storage until regeneration
processing. Because the Pepco
system used a cocurrent dryer, dust
recovered from the off-gas was fed
directly to the dryer product, and the
two were conveyed together to
the storage silo. At both installations,
the cleaned off-gas was routed back
to the venturi absorber where it
mixed with entering flue gas and was
cleaned of remaining particles.
The major factor affecting the
solids separation and drying
equipment is the type of MgSO3
crystals to be processed by the dryer.
Both magnesium sulfite trihydrate
(MgSO3 3H2O) and magnesium
sulfite hexahydrate (MgSO3 6H2O)
can be formed in the absorber:
MgS03 + 3H20 -"
MgS03-3H2Ol
MgS03 + 6H20 -»
MgS03 6H20 i
(12)
(13)
Based on initial pilot tests, the
Boston Edison scrubber was
designed to process MgS03 6H20
crystals. Shortly after operation
began, however, it was discovered
that MgSO3 3H2O crystals were
forming in the absorber. Because
the MgSO3 3H2O crystals are much
finer than MgS03 6H20 crystals,
excessive dusting and accumulation
of solids occurred in the dryer.
Drying ability decreased, and such
modifications as installation of
screens and lump breakers were
necessary to alleviate the problem.
An EPA-sponsored study recently
established parameters for predicting
whether trihydrate or hexahydrate
crystals will form iri the absorber.
Hydration formation is a function
of such variables as slurry density
and volume, seed crystal size
and composition, solution composi-
tion, and temperature.8 Familiarity
with these parameters will help
prevent further problems in the
drying operation.
Regeneration
Two important design considerations
for regeneration processing
equipment are the physical location
of the regeneration facility and
calciner design.
Both the Boston Edison and Pepco
demonstration units regenerated
MgO off site at Esslex Chemical's
sulfuric acid plant in Rumford,
Rhode Island. Several factors,
however, must be considered in the
selection of on-site versus off-site
regeneration facilities.
On-site facilities are more eco-
nomical in terms of fuel and
transportation costs. At an on-site
facility, the dryer product enters
the calciner at approximately 400° F
(200° C), whereas the calciner
feed stream cools to ambient tem-
peratures en route to an off-site
regeneration facility. Fuel require-
ments for calcinatipn increase as
a result of:
Heat loss
Surface oxidation of MgSO3
crystals during storage and
transportation
In addition, trucking solids to and
from an off-site regeneration facility,
increases operating costs for the
FGD system.
Construction costs for regeneration
facilities also must be considered.
Only minor capital expenditure
was required to modify the existing
Essex Chemical Company plant for
the MgO FGD process. In other
instances, it may be more cost effec-
tive to send the solids produced by
several power plants to a single
regeneration and acid production
facility. Benefits from operating such
a centralized plant might offset
increases in fuel and transportation
costs.
Two basic calciner designs must
be considered for use in magnesia
regeneration. Essex Chemical's
Rumford plant uses a rotary calciner,
although fluid bed systems also are
effective. Several companies concur
that fluid bed regeneration eventually
will be the more attractive alternative
because its estimated overall
operating costs are lower.2
Essex Chemical's rotary calciner
required minor modifications during
the Boston Edison demonstration.
A friction seal was installed
to prevent air leakage into the
calciner's firing hood and maintain
the proper reducing conditions for
the conversion of MgS04 to MgO.
A fan and short stack also were
added to the calciner to prevent
hydrocarbon startup vapors from
entering the sulfuric acid towers
and blackening the acid.6
SO2 Processing
Because SO2 conversion is a
processing step that varies depend-
ing on the end product desired,
only cursory design considerations
are presented here.
Cost, storage, and transportation
are some of the aspects that must
be considered in the production of
14
-------
sulfur, sulfuric acid, and dry,
compressed sulfur dioxide. The
capital costs of sulfur production
are high because the operation is
relatively complex. In addition, raw
material and utility costs are
approximately twice as high for
sulfur production as for sulfuric acid
production.
The major consideration in sulfur
dioxide production is the limited
market for dry, compressed SO2.
Although liquid SO2 is used in the
food processing and paper industries,
its marketing potential is low com-
pared with those of sulfur and
sulfuric acid. For this reason, produc-
tion of dry, compressed S02 is not
a feasible alternative.
Although sulfuric acid is a highly
marketable product, its generation
requires considerable storage
and transportation capacity. In
addition, H^SO^ requires more care
in handling because of its corrosive
nature.
A trade-off exists between the
reduced raw material and utility costs
of acid production and the lower
storage and transportation costs
associated withjsulfur production.
Site-specific considerations, such
as raw material availability, also
will affect the final S02 conversion
process decision. ATennessee Valley
Authority (TVA) computerized
marketing study of sulfur and sulfuric
acid production concludes that
FGD byproduct sulfur is not yet
competitive with FGD byproduct
sulfuric acid. Nevertheless, a
relatively small reduction in total
FGD byproduct sulfur costs could
make byproduct sulfur production
competitive.9 To date, all MgO FGD
demonstration systems in the United
States have selected H2SO4 produc-
tion as the conversion process.
15
-------
Eddystone magnesium oxide scrubber
16
-------
Environmental
Considerations
The MgO FGD process has demon-
strated continued ability to remove
sulfur dioxide from flue gas. System
operation tests with both fresh
and regenerated MgO at Boston
Edison's Mystic: Station and
Pepco's Dickerspn Station con-
sistently have resulted in 90 percent
or greater SO2 removal.
j
I
As a regenerate' process, MgO FGD
eliminates the rf^ajor waste disposal
problems associated with such
throwaway systems as lime/lime-
stone and dual ajkali FGD. Impurities
accumulate in the closed system,
however, and MgO FGD must
provide for fly ash and chloride
removal. Flue gas is pretreated to
keep fly ash andlchlorides out of the
slurry. In cbal-firpd boiler applica-
tions, pretreatment removes over
99 percent of th;e particles in
the flue gas stream. Boston Edison's
oil-fired demonstration unit
did not pretreat the flue gas and
achieved particlejremoval efficiencies
of only 50 to 70; percent. Approxi-
mately 70 perceht of the chlorides
can be removed'in a venturi
prescrubber; the remaining chlorides
can be withdrawn in a spray
chamber.10
Magnesium compounds are also
important emissions from the MgO
FGD process. Fugitive emissions
of magnesium compounds include
magnesium salts entrained in
flue gas vented from the absorber,
scrubbing liquor spillage, and dust
that escapes from the handling
and transfer of dry magnesium
compounds.
Measurements taken during a
13-day test at Boston Edison's
Mystic Station to estimate mag-
nesium losses from the system
indicated a daily loss of 750 Ib
(340 kg) of magnesium from absorp-
tion and solids separation/drying.
These decrements amount to
approximately 3.5 percent of the
MgO circulating in the FGD system.
Further losses of 4.7 percent'1,000
Ib/d (450 kg/d)were measured
during regeneration and acid
production. These losses were
caused primarily by the removal and
disposal of large lumps of mag-
nesium compounds from the calciner,
however, and they can be reduced
significantly by the addition of
pulverizing equipment.1
17
-------
Status of Development
The MgO FGD technology was
developed initially for use in the
pulp and paper industry, which
employs a magnesium-based liquor
in the pulping operation. The pulp-
ing liquor is burned in a recovery
furnace, producing MgO powder and
an SO2-rich stream.: The MgO
is slaked and routed to a series of
venturi scrubbers where SO2 is
removed from the recovery furnace
off-gas, thus regenerating the
pulping liquor.
Efforts to apply the:magnesium-
pulping process to SO2 scrubbing
began in the 1930's. Research
in Russia, Japan, Germany, and
the United States to develop an FGD
process using magnesia as the
absorbent resulted in the following
three major variations of the
MgO wet scrubbing FGD process:
A basic slurry of MgO and MgSO3
A slurry of MgO; MgSO3, and a
scrubbing reaction activator,
manganese dioxide (Mn02)
An acidic solution of Mg(HS03)2,
MgSO3, and MgSO4
All three variations are feasible,
but only the first two are capable
of removing 90 percent SO2 from flue
gas. The high vapor pressure of sulfur
dioxide over the solution of sulfites
used in the third method lowers
removal efficiencies to 80 to
85 percent.
Basic slurry is the process variation
described in this report. The
process was developed in the United
States by Babcock and Wilcox
Company of Barberton, Ohio, and by
Chemico-Basic, a joint company
formed by Chemico of New York and
Basic Chemicals of Cleveland, Ohio.
Russia and Japan also have con-
centrated on the basic slurry process,
whereas Germany has investigated
the scrubbing process that uses
an MnO2 activator.2
Three commercial scale MgO FGD
systems have been installed on
power plant boilers in the United
States, and three more are being
planned (Table 1). Although the
EPA-sponsored demonstration
programs at the Mystic and Dickerson
Stations have been terminated,
the testing has established the
MgO FGD process as a feasible
system for control of sulfur dioxide
emissions.2 Philadelphia Electric
Company (Peco) collected
valuable data at its Eddystone
plant before the system was
terminated. As a result of MgO FGD
system performance, additional
units are under construction at
Peco installations.
18
-------
j ;
Table 1 . j
Planned and Completed Magnesium Oxide FGD Systems in the United States
i
Utility company, station, and location3
Installed:
Potomac Electric Power, Dickerson 3: Frederick MD . . .
Philadelphia Electric, Eddystone 1 A: Eddystone PA ...
Planned: Philadelphia Electric:
Eddystone 1 B* Eddystone PA
Eddvstone 2- Eddystone PA
FGD
unit
(MW)
150
95
105
150
240
334
I
Gas volume
| treated
(1 .0'OO stdft3/min)
I
! 307
| 213
! 237
]
i 300b
] 480b
! 670b
i
Fuel
Type
Oil
Coal
Coal
Coal
Coal
Coal
%S
2.5
3.5
2.3
3.0
2.6
2.5
% S02
removal
(design)
90
NA
90
NA
NA
NA
Startup
date
1972
1973
1975
1983
1982
1982
Status
Terminated
Terminated
Terminated
Under construction
Under construction
Under construction
aAII units shown are retrofit. ;
bEstimated: stdft3/min = 2,000 X MW rating. j
Note.NA = not available. . " !
SOURCES: Smith, M., M. Melia, and N. Gregory, EPA Utility FGD Survey:\April-June 1980, EPA 600/7-80-029c, Research Triangle Park NC, July 1980.
Smith, M., M. Melia, and N. Gregory, EPA Utility FGD Survey: October-.December 1979, EPA 600/7-80-029a, NTIS No. Pb 80-1 76-811, Research
Triangle Park NC, Jan. 1980. Sommerer, D. K., Magnesia FGD Process Testing on a Coal-Fired Power Plant, EPA 600/2-77-1165, NTIS No. Pb 272-952,
Research Triangle Park NC, Aug. 1977. Koehler, G., Magnesia Scrubbing Applied to a Coal-Fired Power Plant, EPA 600/7-77-018, NTIS No. Pb
266-228, Research Triangle Park NC, Mar. 1977. Koehler, G., and J. A. Burns, The Magnesia Scrubbing Process as Applied to an Oil-Fired Power Plant,
EPA 600/2-75-057, NTIS No. Pb 247-201, Research Triangle Park NC! Oct. 1975. Isaacs, G. A., Survey of Flue Gas Desulfurization Systems,
Eddystone Station, Philadelphia Electric Company, EPA 650/2-75-057;f, NTIS No. Pb 247-085, Research Triangle Park NC, Sept 1975.
Magnesium sulfite kiln heater
19
-------
System Requirements
Raw Materials and Utilities
Compared with a liitie/limestone
process, the MgO FGD process
has a relatively low raw material
requirement and a relatively
high FGD process energy require-
ment. Although the regeneration
of MgO minimizes the cost of
raw materials, the separation/drying
and regeneration processing steps
require substantial quantities
of fuel oil, resulting in higher energy
requirements for the process.
In terms of ground-to-ground
energy requirements, MgO FGD
compares more favorably with the
lime/limestone process. A major
factor is the energy credit for
the byproduct sulfuric acid. Sulfuric
acid usually is produced from
elemental sulfur that is mined by
the Frasch method, an energy-
intensive operation. Byproduct acid
production conserves the energy
consumed by mining, transporting,
and converting sulfur to sulfuric acid.10
Table 2 presents the estimated
raw material and utility requirements
for three different MgO FGD
systems. The information is taken
from a 1980 TVA study.
This information is based on
converting the SO2 stream from the
regeneration processing area
to H2SO4 in a conventional contact
sulfuric acid plant. This conversion
process requires a catalyst but
generates a heat credit. Agricultural
limestone is used to neutralize
the chloride-rich bleed stream from
the venturi prescrubber. These
and other raw material and utility
requirements vary for different
conversion processes.
Installation Space and Land
Calculations have been made of the
installation space required for an
MgO FGD system applied to a new
500-MW boiler burning 3.5 percent
coal. The total estimated require-
ment for the FGD unit and sulfuric
acid plant is 2.34 acres (0.95 ha).
Approximately 0.78 acre (0.32 ha),
or one-third of this space, is
required for the pretreatment and
absorption processing equipment.
An estimated 0.85 acre (0.34 ha)
is required for the regeneration
operation and 0.71 acre (0.29 ha)
for the sulfuric acid plant. Space
for the pretreatment and absorption
equipment is the most critical
Table 2.
Estimated Raw Material and Utility Requirements for Magnesium Oxide
FGD Process
Component
Size of new
coal-fired plant (MW)
200
500
1,000
Raw materials:
Catalyst3 (ft3/yr) ;
Utilities: :
Fuel oil (1 06 gal/yr).
Steam (1 09 Btu/yr)
Electricity (1 Oe kWh/yr)
Heat credit (1 09 Btu/yr)
600
26
1 ,330
2.6
206
965
26
55
1,470
64
3,240
6.3
503
2,359
62
136
2,840
120
6,260
12
973
4,561
119
262
"Catalyst for sulfuric acid plant.
Note.Base: 3.5% sulfur coal; plant operating time of 7,000 h/yr; meets emission regulation of
1.2 lbS02/106 Btu. ;
SOURCE: Anderson, K. D,, J. W. Barrier, W. E. O'Brien, and S. V. Tomlinson, Definitive SOX Control
Process Evaluations: Limestone, Lime, and Magnesia FGD Processes, EPA 600/7-80-001, NTIS
No. Pb 80-196-314, Jan. 1980.
20
-------
Centrifuge
To MgO recovery
plant
120 ft
From MgO recovery
plant
Key
Flue gas/off-gas
Cleaned flue gas
Absorption liquor
t£^I Sulfur products
Other systems
Figure 7. I
Retrofit Scrubbing System for Boston Edison's Installation
requirement in a utility application
because the equipment must be
located near the boiler and the
stack, whereas the regeneration
operation and the acid plant can be
located at a remote site. Although
these space requirements are
approximate, retrofit installations
usually require more space than new
designs.
Figure 7 shows an example of
a retrofit installation forthe 150-MW
Boston Edison installation. In
general, retrofit installations require
more and longer piping and ducting
but do not require a prescrubber,
because most existing boiler
plants already haye particle
controls. j
Flue gas pretreatment is necessary
for coal-fired plants. Land require-
ments for fly ash disposal have
been estimated at 76 acres
(31 ha). The basis for this estimate
is a new 500-MW boiler burning
coal containing 12 percent ash.
A 30-year plant life with an average
capacity factor of 48.5 percent
is assumed.4
21
-------
Costs
Estimated and actual costs for an
FGD installation can vary widely
depending on the assumptions
made, options included, degree of
redundancy, and conditions of
operation. Sampleicost estimates
prepared by TVA are presented in
this report.
Table 3 delineates the capital and
annual operating costs for MgO
slurry FGD systems installed on
different sizes and types of boilers.
The costs are subject to variation,
depending on site-specific factors.
Specific cases may be evaluated
in terms of the bases used in Table 3.
Table 4 lists specific components
of the annual operating costs for
a typical MgO FGD system on a new
boiler and provides examples
of the contribution of each com-
ponent to the annual operating cost.
MgO FGD is an equipment-intensive
process with higher capital require-
ments than lime/limestone
processes. Regeneration of the
spent magnesia, including process-
Table 3.
Estimated Capital and Operating Costs for Magnesium Oxide FGD Process3
System characteristics
Size
(MW)
200
200
500
500
500
500
500
500
1,000
1,000
Application
Existing
New
Existing
New
Existing
New
New
New
Existing
New
Fuel
Type
Coal
Coal
'Coal
Coal
Oil
Coal
Coal
Coal
Coal
, Coal
%S
3.5
3.5
3.5
2.0
2.5
3.5
3.5
5.0
3.5
3.5
Plant
life
(y)
20
30
25
30
25
30
30
30
25
30
%SO2
removald
S
S
S
S
R
S
90
S
S
S
Total capital
investment1"
$106
35.12
34.44
66.84
53.70
42.64
65.91
68.62
75.81
103.64
101.35
$/kW
176
172
134
108
85
132
137
152
104
101
Annual operating
costs0
$106
9.81
9.27
18.31
14.66
12.18
17.79
18.47
20.41
28.81
27.74
Mills/kWh
7.01
6.62
5.23
4.19
3.48
5.08
5.28
5.83
4.12
3.96
aMidwest plant location. Stack gas reheat to 1 75° F. Investment and revenue requirement for fly
ash removal and disposal excluded.
bProject beginning mid-1977, ending mid-1980. Average cost base for scaling, mid-1979.
Minimum in-process storage; only pumps are spared. Disposal pond located 1 mile from power
plant. FGD process investment estimate begins with common feed plenum downstream of
electrostatic precipitator. No overtime pay.
C1980 revenue requirements. Power unit operating 7,000 h/yr.
dS = meets emission regulation of 1.2 Ib S02/106 Btu. R = meets allowable emission of 0.8 Ib
S02/106 Btu.
SOURCE: Anderson, K. D., J. W. Barrier, W. E. O'Brien, and S. V. Tomlinson, Definitive SOX
Control Process Evaluations: Limestone, Lime, and Magnesia FGD Processes, EPA 600/7-80-001,
NTIS No. Pb 80-196-314, Jan. 1980.
22
-------
Table 4. j
Annual Operating Costs for a Magnesium Oxide FGD System on a 500-MW Coal-Fired Boiler
Component
Direct costs:
Delivered raw materials:
Maanesium oxide
Catalyst
Agricultural limestone
Total raw materials
Conversion costs:
Operating labor and suoervision
Utilities:
Fuel oil
Steam
Process water
Electricity
Heat credit
Maintenance, labor and material
Analyses . . . . :
Total conversion costs
Total direct costs
Costs
Annual
quantity , , Annual
Unit ($) operating
($1,000)
23RQ?X1O6nal n 1 9/1 n-3 rial OQQ m
-
Indirect costs:
Capital charges: 1
Depreciation, interim replacements, and insurance at 6.0% of total
depreciable investment . ' o oK1 Qn
Average cost of capital and taxes at 8.6% of total capital invest-
Overhead:
Plant, 50% of conversion costs less utilities
Administrative, 10% of operating labor
Marketing, 10% of byproduct sales revenue
Total indirect costs
Gross averaae annual ooeratina costs
'
Byproduct sales revenue, 1 00% sulfuric acid
Net annual operating costs
Mills/kWh
0.126
0.001
0.014
0.141
0.170
0.718
0.288
0.081
0.512
(0.077)
0.705
0.041
2.438
2.579
1.103
1.620
0.458
0.017
0.077
3.275
5.854
(0.771)
5.083
Note.Midwest plant location, 1980 revenue requirements. Remaining life of power plant, 30 yr. Power unit on-stream time, 7.000 h/yr 1 500 100
tons/yr coal burned, 9,000 Btu/kWh. Stack gas reheat to 175° F. Meets emission regulation of 1.2 Ib S02/106 Btu. Investment and revenue require-
ment for removal and disposal of fly ash excluded. Total direct investment, $35,354,000; total depreciable investment, $64,365 000- total capital
investment, $65,911,000. ! '
SOURCE: Anderson, K. D., J. W. Barrier, W. E. O'Brien, and S. V. Tomlinson, Definitive SOX Control Process Evaluations: Limestone Lime and
Magnesia FGD Processes, EPA 600/7-80-001, NTIS No. Pb 80-196-314, Jan. 1980.
23
-------
Second-stage slurry ducting to scrubber
ing, drying, and calcination,
requires a capital investment of
almost $9 million. The recovery
system may also require chloride
removal before the SO2 absorber,
necessitating an additional
$5 million. Sulfuric acid production,
storage, and shipping increase costs
another $7 million. These capital
requirements of approximately,
$21 million exceed by over
$14 million the savings gained by
eliminating disposal of solid
waste in ponds.10
Under certain conditions, MgO FGD
can be economically competitive
with throwaway processes. Oil-
fired installations, for instance,
would not require a chloride
prescrubber. At sites where disposal
of slurry waste in ponds is not
practical, the increased costs for
fixation and landfill would also
favor MgO FGD as an alternative.10
24
-------
1
References
1U.S. Environmental Protection
Agency. Flue Gas Desulfurization
and Sulfuric Acid Production
Via Magnesia Scrubbing. EPA-
625/2-75-007,|NTIS No.
Pb 258-817. Research Triangle
Park NC, EPA, 1 975.
|
2McGlamery, G. JG., R. L. Torstrick,
J. P. Simpson, 'and J. F. Phillips,
Jr. Conceptual 'Design and
Cost Study. Sulfur Oxide Removal
from Power Plant Stack Gas.
Magnesia Scrubbing, Regenera-
tion: Production of Concentrated
Sulfuric Acid. EPA R2-73-244,
NTIS No. Pb 222-509. May 1 973.
i
30ttmers, D. M.,|jr., E. F. Aul, Jr.,
R. D. Delleney, ;G. D. Brown,
G. C. Page, and D. O. Stuebner.
Evaluation of Regenerable Flue
Gas Desulfurization Processes.
Rev. rep. 2 vols' EPRI Contract
No. RP 535-1. Austin TX,
Radian Corporation, July 1976.
4McGlamery, G. <3., R. L. Torstrick,
W. J. Broadfoot; J. P. Simpson,
L J. Henson, S.j V. Tomlinson,
and J. F. Young1,,Detailed Cost
Estimates for Advanced Effluent
Desulfurization Processes.
EPA 600/2-75-006, NTIS No.
Pb 242-541. Jap. 1975.
i
5Koehler, G. Magnesia Scrubbing
Applied to a Coal-Fired Power
Plant. EPA 600/7-77-018,
NTIS No. Pb 266-228. Research
Triangle Park NC, Mar. 1977.
6Koehler, G., and J. A. Burns.
The Magnesia Scrubbing Process
as Applied to an Oil-Fired Power
Plant. EPA 600/2-75-057, NTIS
No. Pb 247-201. Research
Triangle Park NC, Oct. 1 975.
7Sommerer, Diane K. Magnesia
FGD Process Testing on a Coal-
Fired Power Plant. EPA 600/2-
77-165, NTIS No. Pb 272-952.
Research Triangle Park NC,
Aug. 1977.
8Lowell, Philip S., Frank B.
Meserole, and Terry B. Parsons.
Precipitation Chemistry of
Magnesium Sulfate Hydrates in
Magnesium Oxide Scrubbing.
EPA 600/7-77-109, NTIS No.
Pb 277-086. Research Triangle
Park NC, Sept. 1 977.
90'Brien, W. E., and W. L. Anders.
Potential Production and
Marketing of FGD Byproduct
Sulfur and Sulfuric Acid in the
U.S. (1983 Projection). EPA
600/7-79-106, NTIS No. Pb 299-
205. Research Triangle Park NC
Apr. 1979.
10Anderson, K. D., J. W. Barrier,
W. E. O'Brien, and S. V. Tomlinson.
Definitive SOX Control Process
Evaluations: Limestone, Lime and
Magnesia FGD Processes. EPA
600/7-80-001, NTIS No. Pb 80-
196-314. Jan. 1980.
25
-------
26
This summary report was prepared jointly by the Radian Corporation of
Austin TX and the Centec Corporation of Reston VA. Jack M. Burke and
Elizabeth D. Gibson of Radian are the principal contributors. Michael A.
Maxwell is the EPA Project Officer. Photographs were taken at Philadelphia
Electric Company's Eddystone facility in Philadelphia PA.
Comments on or questions about this report or requests for information
regarding EPA flue gas desulfurization programs should be addressed to:
Emissions/Effluent Technology Branch
Utilities and Industrial Power Division
IERL, USEPA(MD 61)
Research Triangle Park NC 27711
This report has been reviewed by the Industrial Environmental Research
Laboratory, U.S. Environmental Protection Agency, Research Triangle Park NC,
and approved for publication. Approval does not signify that the contents
necessarily reflect the views and policies of the U.S. Environmental
Protection Agency, nor does mention of trade names or commercial products
constitute endorsement or recommendation for use.
COVER PHOTOGRAPH: Eddystone Unit No. 1 hardware and coal conveyor
fr U.S.GOVERNMENT PRINTING OFFICE: 1981-758-896
------- |