EPA-R2-72-021
July 1972
Environmental Protection Technology Seriei
STUDY OF THE CHARACTERIZATION
AND CONTROL OF AIR POUUTANTS
FROM A FlUIDIZED-BED B0ll||
THtS02 ACCEPTOR PROCESS
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EPA-R2-72-021
STUDY OF THE CHARACTERIZATION
AND CONTROL OF AIR POLLUTANTS
FROM A FLUIDIZED-BED BOILER-
THE S02 ACCEPTOR PROCESS
by
J.S. Gordon, R.D. Glenn, S. Ehrlich,
R. Ederer, J.W. Bishop, andA.K. Scott
Pope, Evans and Robbins, Inc.
320 King Street
Alexandria, Virginia 22314
Contract No. CPA 70-10
Program Element No. 1A2013
Project Officer: D. Bruce Henschel
Control Systems Laboratory
National Environmental Research Center
Research Triangle Park, N. C. 27711
Prepared for
OFFICE OF RESEARCH AND DEVELOPMENT
U.S. ENVIRONMENTAL PROTECTION AGENCY
WASHINGTON , D. C. 20460
July 1972
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This report has been reviewed by the Environmental Protection Agency and
approved for publication. Approval does not signify that the contents
necessarily reflect the views and policies of the Agency, nor does
mention of trade names or commercial products constitute endorsement
or recommendation for use.
ii
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Ill
TABLE OF CONTENTS
Page
NOMENCLATURE 1-4
1. SUMMARY 1-1
1.1 General 1-1
1.2 Tests in the Fluidized-Bed Colum .... 1-4
1.3 Tests in the Pilot Plant (Fluidized-
Bed Module) 1-5
1.4 Designs and Cost Estimates 1-11
2. CONCLUSIONS 2-1
3. RECOMMENDATIONS 3-1
4. INTRODUCTION 4-1
4.1 Description of a Fluidized-Bed Boiler. . 4-1
4.1.1 The Need for a New Form of Combustion. . 4-1
4.1.2 Defining a Fluidized-Bed Boiler 4-2
4.1.3 What a Fluidized-Bed Boiler is Not . . . 4-3
4.1.4 A Simplified Description of a Fluidized-
Bed Boiler 4-3
4.2 Air Pollution Control Potential of
Fluidized-Bed Boilers 4-7
4.2.1 What is Air Pollution 4-7
4.2.2 Pollution Control Potential of Fluidized-
Bed Boilers 4-8
4.2.2.1 Discussion 4-8
4.2.2.1.1 General 4-8
4.2.2.1.2 Sulfur Oxides 4-8
4.2.2.1.3 Nitrogen Oxides 4-10
4.2.2.1.4 Others 4-11
4.2.2.2 Conclusions 4-14
4.3 Pope, Evans and Robbins' Prior Work 4-15
4.3.1 General 4-15
4.3.2 Once-Through Limestone Injection .... 4-16
4.3.3 Regenerative Limestone Process 4-17
4.3.4 Pollutants Other Than SO- 4-18
4.3.5 The Carbon-Burnup Cell 4-20
4.3.5.1 Tests in the Fluidized-Bed Column. . . . 4-20
4.3.5.2 Tests in the Fluidized-Bed Module. . . . 4-23
4.3.6 Recommendations Based on Prior Work. . . 4-24
4.4 Specific Objectives of This Work .... 4-25
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IV
5. APPARATUS
5.1 Pilot Scale Combustor, FBC ....... 5-1
5.2 Full-Scale Boiler Module, FBM ...... 5-7
5.3 The Pilot Scale Carbon-Burnup Cell, CBC . 5-16
5.3.1 The Initial CBC ............ 5-17
5.3.2 The Modified CBC ............ 5-20
5.4 The Pilot Scale Bed Regenerator, REG . . 5-20
5.5 Instrumentation ............. 5-25
5.6 Materials ............... 5-33
6. RESULTS OF BENCH AND PILOT SCALE TESTS 6-1
6.1 Reduction of Emissions of S0_ ...... 6-1
6.2 FBC Tests .......... ...... 6-1
6.3 FBM/CBC Tests ............. . 6-21
6.3.1 Test with the Short CBC Regenerator. . . 6-21
6.3.2 FBM/CBC Extended Run (Run 168H) - Tall
CBC Regenerator ............ 6-27
6.3.2.1 Test Operation and Burnup Results. . . . 6-27
6.3.2.2 Gaseous Emissions ............ 6-37
6.3.2.3 Boiler System Heat Balance ....... 6-42
6.3.2.4 Bed Specific Gravity .......... 6-42
6.3.2.5 Sulfur Analyses ............. 6-45
6.3.3 FBM/CBC Run with Coal Feed to CBC
(Run 169H) .............. 6-46
6.4 FBM/CBC/REG Extended Run (Run 171H) . . . 6-48
6.4.1 Background ............... 6-48
6.4.2 Extended Run Operation ... ...... 6-50
6.4.3 Arsenic Analyses ............ 6-63
6.5 FBM/REG Test (Test No. 172 H) ...... 6-64
6.6 Bed Particle Size Determinations .... 6-65
7. PRELIMINARY FLUIDIZED-BED BOILER DESIGNS
HEAT BALANCES AND COST ESTIMATES ... 7-1
7.1 Preliminary 30 Megawatt Boiler Designs,
Heat Balances and Cost Estimates . . . 7-1
7.1.1 30 Megawatt Packaged Boiler Concepts . . 7-1
7.1.2 30 Megawatt Boiler Cost Estimates, Flui-
dized-Bed(RV-III vs Pulverized Coal) . 7-8
7.1.3 Overall 30 Megawatt Power Plant Cost
Estimates, Fluid Bed vs Pulverized
Coal ................. 7-8
7.2 Preliminary 300 Megawatt Fluidized
Boiler Concept, Heat Balances and
Cost Estimates ............ 7-8
8. SULFUR RECOVERY ............. 8-1
9. REFERENCES ............... 9-1
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LIST OF FIGURES
No. Page
1 Schematic of Fluidized-Bed Boiler 4-4
2 Fluidized-Bed Column, Construction
Detail, Front View 5-2
3 Fluidized-Bed Column, Construction
Detail, Side View 5-3
4 Air Distribution Grid Button 5-4
5 Section Through Fluidized-Bed Column
Showing Insulated Steel Liner and
Adjustable Cooling Surface 5-5
6 Schematic of FBC Air and Exhaust Gas
Ducting Showing Sampling Points .... 5-7
7 Fluidized-Bed Module, Internal Con-
struction 5-10
8 Fluidized-Bed Module, Internal Con-
struction with Tall CBC Showing
Numbered Thermocouple Locations . . . .5-12
9 Section Through Fluidized-Bed Module
Initial Configuration 5-13
10 Section Through FBM and CBC (Revised
Configuration) 5-14
11 Schematic of FBM Test System Showing
Various Subsystems 5-15
11A Physical Interconnections between FBM,
CBC and REG 5-22
11B Flow Schematic for FBM and REG 5-23
11C Flow Paths of Bed Material for FBM, CBC
and REG 5-24
12 Schematic of Gas Transfer System for
Continuous Monitoring of Sulfur Dioxide
Nitric Oxide, Carbon Dioxide and
Hydrocarbons 5-26
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VI
List of Figures (continued)
No. Page
13 Schematic of the FBM Gas Sampling System 5-29
14 Schematic of CBC Gas Sampling System . . . 5-30
15 Bed Sulfur Accumulation vs Time FBC Run
C-321 6-5
16 Bed Sulfur Content vs Time FBC Run C-322 6-11
17 Bed Sulfur Content vs Time FBC Run C-323 6-15
18 Bed Sulfur Content vs Time, FBM Run B-18
Illustrating Regeneration 6-26
19 FBM Steam Rate vs Carbon Burnup, Run 168H 6-33
20 FBM-CBC System Carbon Burnup vs CBC SO_
Output, Run 168H 6-34
21- Coal Particle Size Analysis, Run 168H. . . 6-38
22 FBM Flue Gas SO_ Concentration vs Bed
Sulfur Content 6-39
23 FBM Flue Gas NO Concentration vs Bed
Sulfur Content 6-41
24 Heat Balance - Run 168H 6-44
25 CBC Carbon Burnup and Tandem Cycle
Burnup vs Oxygen Level 6-47
26 Coal Fired Fluidized Bed Utility Boiler
Factory Assembled - 300,000 Ib/hr 1270
psig 925° FTT - FBP-I 7-2
27 Boiler Coal Supply System Detail 7-3
28 300,000 Ib/hr Packaged Fluidized-Bed
Boiler 1270 Psig, 925° F for High Sul-
fur Coal, FBP-III 7-7
29 300 M.W. Utility Boiler Concept, 1,900,000
Ib/hr, 2400 Psig, 1000°F, Reheat
1,600,000 Ib/hr, 650 1000°F 7-10
30 Gas, Fly Ash and Bed Regeneration
Schematic 7-14
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Vll
LIST OF TABLES
No.
1
2
3
4
5
6
7
8
9
10
11
12
13
14
15
16
17
18
19
20
21
1
FBC Test Data Summary (Test C-321) ....
FBC Test Data Summary (Test C-322) ....
FBC Test Data Summary (Test C-323 ....
FBC Test Data Summary (Test C-324) ....
FBM-CBC Test Summary (Test B-18) ....
FBM/CBC System Log - 168 Run . .
Summary of Burnup Data , Run 16 8H
Heat Balance Data Run 168H
FBM-CBC Run with Coal Feed to CBC(Run 169H)
FBM/CBC/REG System Log - Run 171H. . . .
Fly Ash Characteristics - Run 171H. . . .
Calcium Balance, FBM/CBC/REG Test No. 171H
Heat Balance Data - Run 171H
RV-ll Heat Balance and Surface Summary
RV-III Heat Balance' and Surface Summary
Incremental Cost, 300,000 Lb/Hr Fluidized-
Bed Boiler Installation
30 Megawatt Power Plant Cost Estimates. .
300 Megawatt Fluidized-Bed Boiler Single
Bed Level, Natural Circulation Heat
Balance and Surface Requirements. . . .
Incremental Cost of 300 MW Boiler ....
Overall 300 MW Fluidized-Bed Power Plant
Capital Cost Estimate .
Capital i Operating Costs
Page
6-3
6-8
6-12
6-17
6-24
6-28
6-35
6-43
6-49
6-51
6-58
6-61
6-62
\J \J £*
7-6
7-9
"7 TO
/ — L£.
7-13
7-17
7-18
7-19
8-3
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VI11
APPENDICES
No. Page
A FBC Specifications A-l
B FBM Specifications B-l
C CBC Specifications C-l
D Laboratory Apparatus D-l
E Backgound Information on the "Chemical
Attriter Process" E-l
F Sulfur Balances F-l
G Particle Size Determinations G-l
H FBM Run B-18 Log (Table) H-l
I Run 168H - Condensed Data - CO and Hydro
Carbon Emissions 1-1
J Run 168H- Condensed Data - CO and Hydro
Carbon Emissions J-l
K CBC Fly Ash Analyses, Rurt 168H K-l
L Cost Estimate for RV III Design L-l
M Letter from CE Industrial Boiler Operations M-l
N 300 MW Boiler - Weight and Cost Estimate
Data N-l
0 Prorating for Case I Fluidized Bed. . . . O-l
p Further Information on Recovery Processes P-l
Q Arsenic Analysis Data Q-l
R Conversion Factors R-l
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1-1
1. SUMMARY
1.1 General
Work is described on studies leading to low-cost
environmentally-clean power generation from coal. A
regenerative limestone process for multi-cell fluidized-
bed coal combustion-desulfurization is described.
Pope, Evans and Robbins (PER), in the final phase
of a project* sponsored by the Office of Research and
Monitoring, Environmental Protection Agency, has studied
the air pollutant emissions reduction capability of
fluidized-bed boilers for the combustion of coal. In
earlier work with sintered ash beds, it was found that
sulfur-oxide emissions could be markedly reduced by
injecting finely divided limestone on a once-through
basis into a coal-burning, atmospheric fluidized-bed
operating at 1500°F to 1600°F&, and with about 3% residual
oxygen or more remaining in the flue gas.** At 3%
residual oxygen or less in the flue gas, a significant
percentage of the input coal's fuel value would appear
as carbon in the boiler fly ash. A comprehensive search
for methods to reduce the elutriation loss of fuel, which
ranged to 15% for the high air rates (superficial velocity
about 9 ft/sec) used in the PER designs, led to the inven-
tion of the Carbon-Burnup Cell.*** The alternative,
elutriation minimized by operation at low gas velocity,
was shown to be uneconomical due to increased boiler
size. The Carbon-Burnup Cell was simply a region of
'Thisreport describes the results of experiments
carried out between November 1970 and August 1971.
Primary emphasis has been placed on bituminous
coal combustion.
** Reference 2 describes the details of the sulfur
control program.
*** U.S. Patent 3,508,506.
# Although it is EPA policy to use metric units, certain
non-metric units are used in this report to reflect
actual test results. Please use the conversion factors
in Appendix R if you are more familiar with the metric system.
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1-2
a fluidized-bed boiler into which carryover from the primary
boiler bed was injected, and in which the amount of heat
transfer surface per unit heat release was lower than in the
adjacent coal-fired regions. In this way a temperature in
the 1900°F to 2050°F range could be maintained. Design criteria
for an effective Carbon-Burnup Cell were produced in the
previous phase of this program*. Reinjection of carbon-
bearing fly ash into the primary boiler bed was shown inferior
to the Carbon-Burnup Cell technique, since the fly ash
carbon must compete with fresh coal for the available
oxygen.
It was desired to improve upon the sulfur capture
results obtained with once-through limestone injection
into a coal-burning, sintered ash fluidized bed. Tests
PER ran in May 1969 showed that limestone was a suitable
bed material in place of the sintered inert ash. Lime
bed operation was found promising for SO2 capture. A
continuous capture/regeneration operation of the sorbent
was visualized. PER also found that injection of salt
(sodium chloride) aided lime SO2 capture effectiveness
and also increased combustion efficiency. The major
goal of the study described in this report was to inves-
tigate these types of process operations.
Therefore, the program described in this report
consisted of two major tasks: laboratory scale, batch
type coal combustion experiments using limestone beds
for sulfur capture, with and without salt additive to
modify lime effectiveness; and pilot plant (boiler system)
experiments in which the limestone from the primary
combustion zone is made to undergo continuous regeneration
in a regeneration section. The primary benefit of the
regenerator technique is to minimize makeup limestone
requirements at continuous high sulfur capture levels
Reference 1.
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1-3
in the boiler. Without regeneration, large quantities
of limestone would have to be cross-flowed through the
boiler and discarded. A second advantage of this regener-
ation technique is that the regeneration zone produces
a SO_-rich flue gas stream, which is volumetrically
only a small fraction of the total boiler gas flow; this
SO_-rich gas may be fed to a sulfur recovery operation.
The regeneration scheme of the SO_ Acceptor Process
works as follows: sulfated lime from the 1500 F pri-
mary zone continuously circulates to the 2000 F regen-
eration zone where carbon-bearing fly ash (MK I) or
coal (MK II) is burned with low excess air. As des-
cribed in Refs. 1 and 2, SO- is driven off in high
concentration and lime is recovered for reuse by the
reaction:
CaSO4 + C + JsO2 -»• CaO + SC>2 + C02
In the SO. Acceptor Process MK I configuration, the
regeneration is performed in the Carbon-Burnup Cell and
the process consists of two zones. In the MK II config-
uration, separate regenerator and Carbon-Burnup Cells
are used, and the process consists of three zones.
The goals of the present pilot plant study were:
a. 90% or better SO- removal from primary cell
b. 3-4% or more SO- in regenerator flue gas
c. 98% or better carbon burnup in the system
In these tests, the regenerator was fly ash or coal-
fired. Use of fuel gas for CaSO. regeneration has been
avoided for economic reasons. Because of current natural
gas shortages, there is little possibility that a gas-
fired regenerator could be guaranteed a fuel supply.
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1-4
Use of pulverized coal fuel has been avoided in
order to achieve high combustion efficiency, as well
as economy of coal preparation. Superficial gas veloci-
ties have been above 6 ft/sec for economic reasons, i.e.,
to minimize boiler size.
1.2 Tests in the Fluidized-Bed Column
The purpose of the FBC tests was to explore SO-
capture and salt addition on a bench scale prior to
a major expenditure of resources in FBM testing. The
first portion of the effort, to explore limestone bed
effectiveness for SO_ capture, was carried out in the
Fluidized-Bed Column (FBC) which, during these tests,
had a plan area of 0.86 square feet. FBC operation
in this program was nonregenerative. In previous
programs, the FBC had been used in cyclic sorption-
regeneration tests.
Coal was burned in this device over a wide range
of conditions, summarized as follows:
Bed temperature: 1470°F to 1650°F
Bed depth (static): 14" to 22"
Air rate: 700 to 800 Ib/hr ft2
Fuel rate: 47 to 81 Ib/hr ft2
The method of FBC operation was: the initial bed
for each test was fresh limestone at a low bed level
(about 8 inches). The bed was then brought to calcining
temperature by the gas burner and coal feed. More
bed was then added to create the desired depth. A
calcining period of several hours thus precedes the
start of SO- release buildup. If the gas burner light-
off had been attempted with the full bed depth, prep-
aration time would have been much longer.
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1-5
Flue gas analyses, sulfur balances, and bed particle
size vs. time analyses were compiled for these tests.
Portions of the operation utilized salt addition to
provide lime activity enhancement as well as combustion
catalysis in a nonregenerative operation.
Typical results of FBC tests are: carbon burnup
in one pass is typically 90 to 96%; bed sulfur content
increases with time since the operation is nonregenerative;
as the bed deactivates, the gas S0~ level also increases
from 200 ppm to 1000 or more in the absence of salt;
S0_ levels with the aged lime bed are reduced dramatically
by salt addition. The fate of the salt is to produce
more fly ash, since NaCl is believed to be converted
to Na2Ca(SO.)_ and the gravimetric factor, CaNa-fSO.)-/
NaCl, is 4.7. The chlorine fed is believed to be tied
up by lime as CaCl-. No equipment corrosion due to
salt has been observed. The vapor pressure of NaCl
as a function of temperature is listed in various sources;
e.g., the JANAF Thermochemical Tables. At 1500°F, it
is significant and NaCl acts as a homogenous vapor
phase combustion catalyst. Also, the isokinetic samples
in the cooled flue gas stream display a high salt fog
content. Solid salt feeder operation was unreliable.
Good results were obtained with aqueous salt solution
injection, but the.water vaporization thermal penalty
is undesirable.
1.3 Tests in the Pilot Plant (Fluidized-Bed Module)
Several tests of longer duration were conducted
in the Fluidized-Bed Module (FBM), a boiler with an
2
air distributor grid area of 9 ft . . Its rating is
5000 Ib/hr steam, 800 Ib/hr coal input. Two different
designs for Carbon-Burnup Cells were appended
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1-6
to the FBM to determine the problem areas in operating
a regeneration-cycle fluidized bed with two distinct
temperature zones (1500°F in FBM, 2000°F in CBC) . A
vertical coal feeder, with flow splitter and two horizon-
tal opposed outlets, was used in the FBM. Some operating
problems occurred with this feeder. Coal feeding to
a f luidized-bed boiler is an area which requires further
development effort.
As discussed in Reference 1, an experimental Carbon-
Burnup Cell was appended to the FBM, to bring system
carbon burnup into the 98-100% range. Without the CBC,
the FBM provides 90% carbon burnup in one pass. CBC
emissions of hydrocarbons and carbon monoxide were
essentially zero when operating in the Carbon-Burnup
Cell mode. It was found that for the small CBC (^1.1 ft )
an opening in the barrier between the FBM and CBC regions
of just 2 square inches permitted the desired temperature
difference to be achieved. An interchange rate between
the FBM and CBC of 12,000 pounds of bed material per
hour per square foot of opening was estimated from two
transient heating tests described in Reference 1. When
the CBC was used as a regenerator, interchange was some-
times supplemented by screw feeding bed material from
the FBM to the CBC, by which means the CBC flue gas
S0_ content was increased.
Two CBC configurations were used in the current
program, both incorporating bed particle "knockouts" or
"baffle screens". Figures 7 through 10 show most of the
details of the two CBC's. The baffle screens consisted
of cylinders arranged in a triangular array. With the
short CBC configuration*, without a baffle screen, large
This original short configuration resulted from the
FBM steam drum layout, it is shown in Figures 7 and 9.
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1-7
quantities of bed material had been expelled from the
CBC. With either baffle screen, the quantity of bed
material carried over was small. With the short CBC (Figure
9) configuration, when the baffle screen was made up of
water-cooled tubes, an initial* heat transfer coefficient
of ^35 Btu/ft hr°F was measured. This was approximately
twice the value that would be predicted from simple
radiation and convection, indicating that an active heat
transfer region exists above the dense phase of a high
velocity (hence turbulent) fluidized bed. With the tall
2
CBC shown in Figure 10, the coefficient was only 20 Btu/ft
hr°F where the temperature is the log mean temperature
differential**. The lower coefficient is caused by the
fact that the baffle screen in the tall CBC was far
above the fluid bed.
When all operational problems has been overcome, the
FBM with the tall CBC achieved an overall carbon combustion
efficiency of 99%.
When the CBC was used as a regenerator, the rate
of regeneration of bed material was shown to be increased
by coal feed to the CBC; but this procedure reduced
carbon combustion efficiency of the system below 98%.
The integrated FBM/CBC boiler system was operated
extensively to investigate process parameters of the
two-cell "SO2 Acceptor Process," labelled the MK I
system. First, a shakedown test of about 28 hours'
duration was made. An extended run of about 80 hours'
duration was then made. In the MK I system, coal is
burned in the FBM where an active lime bed (-8+20 mesh)
at the optimum sorption temperature (1550°F) absorbs
* Prior to onset of noticeable fouling buildup.
** Temperatures and rates of flue gas and cooling
water are measured entering and leaving the
cooler.
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1-8
better than 90% of the SO- emitted by coal combustion.
Carbon-bearing fly ash from the FBM is collected and
introduced to the high-temperature CBC where further
carbon burnup occurs. As a function of CBC fuel and
air rates, temperature and combustion efficiency, FBM
bed material circulating through the CBC is regenerated.
Three benefits result:
(1) A relatively SO_-rich flue gas stream is
synthesized in the CBC which volumetrically is
only a small fraction of the total system flue
gas.
(2) Lime containing CaSO. is regenerated,
minimizing CaCCU makeup and calcining require-
ments in the FBM.
(3) FBM bed density is minimized*, aiding
efficient distribution of coal within the
aerated bed.
Coal containing 4.5% sulfur was used, limiting
the CBC gas SO_ content to a maximum of about 3%**.
Additional process parametric variations were studied
during these tests: coal addition to the CBC/Regenerator
to supplement reducing conditions provided by fly ash
feed; salt addition to the FBM to enhance SO_ capture
and combustion efficiency.
Following these tests, two additional tests of
about 8 hours' duration each were made in an effort
to reduce CBC air rate and hence increase SO- concentra-
tion. CBC operation was on coal only, to aid regenera-
tion at very low CBC air rates. Fly ash feed was restricted
to the CBC to minimize CBC air rate requirements; salt
**
A bed of calcium sulfate is denser than one of
calcium oxide.
The SO_ concentration in the CBC off gas is set by
the mass balance and could come close to the thermo-
dynamic equilibrium concentration of 8 to 11%.
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1-9
was added to the FBM to enhance carbon burnup, thereby
reducing the CBC carbon burnup requirement and associated
air rate.
The conclusions regarding two-cell operation
of the SO- Acceptor Process, MK I, drawn from these
tests are:
(1) The tandem FBM/CBC system operates in the
S0_ Acceptor Process mode continuously and with
stability.
(2) The SO- removal from the primary cell (FBM)
flue gas, as determined by continuous infrared
analysis, is easily maintained in the 90-95%
range.
(3) Limestone makeup requirements are remarkably
low, about 5% of coal input by weight. The same
bed was used in a month of testing, with additions
needed to make up for analytical samples with-
drawn and attrition losses which are very modest
with the 1359 limestone. No limit on the sorp-
tion-regeneration cycles tolerated by the lime
bed has been observed, and activity remains at
a high level.
(4) The coal combustion efficiency of the over-
all system can be maintained for an extended
period at 98% or better using FBM/CBC burnup
conditions comparable to testing earlier in
this Contract*. Under these conditions, a mini-
mum amount of, or preferably no, coal is fed
to.the CBC, O2 levels greater then 1% are main-
tained in the CBC off-gases, and SO- levels
in the CBC gas are below 3%, using the 4.5% S
coal. Under these conditions, sulfur balances
Reference 1
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1-10
(by analysis of gas streams, ash, and bed material)
have not been of high accuracy, and incomplete
regeneration of bed material is probably occurring
due to the need for 0_ levels of more than 1%.
the CBC gas as imposed by burnup demands.
(5) The SO2 concentration in the CBC off-gases,
also as determined continuously by infrared, can
be maintained at levels of 3-4%, corresponding
to tandem operation in the SO- Acceptor Process
MK I mode, for an extended period using less
oxidizing conditions in the CBC (0.5% O_ or less).
Good regeneration occurs at temperatures of
1850°F or higher. Under these conditions, coal
combustion efficiency of the present FBM/CBC
overall system is less than 98%. Regeneration
is aided by coal addition to the CBC. Sulfur
balances are good when coal is fed to the CBC.
(6) At low bed sulfur levels, salt addition
to the FBM acts primarily to increase combustion
efficiency. When SO_ capture is already 90%
or better, the improvement in capture caused
by salt addition is difficult to detect. Acting
as a vapor-phase, homogeneous combustion catalyst,
the amount of salt required is very small (less
than 1% of coal weight). This enhancement is
important, since higher carbon burnup in the
FBM leads to a possible reduction in CBC cross-
sectional area, and consequently lower CBC air
rates, making higher SO_ levels possible in the
regeneration off-gas, at a given value of 0~.
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1-11
Using the results of the tests in the FBM-CBC
system, a regeneration section was designed and an FBM-
CBC-REG three-vessel system built and tested to
investigate the MK II concept. In the CBC, high O- yields
high carbon burnup. In the REG (coal-fired), low 0-
yields high levels of bed regeneration and a low volumetric
flow of high-SO_ process gas (up to 10% SO-). Various
process arrangements are possible. FBM bed may be fed to
the REG. REG bed may be flowed to the FBM or CBC or both.
REG high-carbon fly ash may be fed to the CBC or discarded.
Studies on the optimum arrangement were not performed.
1.4 Designs and Cost Estimates
This report contains a section on boiler design con-
cepts and costs estimates prepared in 1971. Inflation since
that time makes the absolute values presented here obslete
However, cost studies done in 1973 show that the cost
of a fluidized-bed boiler relative to the alternative
furnace design is still favorable.
NOMENCLATURE
CBC Experimental Carbon-Burnup Cell
d Weight mean particle diameter
FBC Fluidized-Bed Column
K Equilibrium constant
eq
ppm Parts per million, volume basis
T Bed temperature, °F
ym Micrometers
1° primary
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2-1
2. CONCLUSIONS
Based on an analysis of the experimental work carried
out under this program, the following conclusions were
drawn:
a. A multicell fluidized-bed pilot scale boiler has
been constructed which can operate at high air rates and
achieve 99+% combustion efficiency. Because of the high
air rate and combustion intensity, cost estimates for com-
mercial boilers show that fluid-bed units will be less
costly than conventional units.
b. The desired level of combustion efficiency may
be achieved by recycling collected carbon-bearing fines
from the primary combustor to a region of the fluidized-
bed boiler called the Carbon Burnup Cell, in which the
bed temperature is in the ranqpe 1950-2050°F and the 02
is above 3%. The Carbon Burnup Cell requirements for
bed depth, firing rate, and air rate have been determined.
The tandem boiler-CBC operation was shown superior to
fly ash reinjection to the primary cell, since carbon in
fly ash is lass reactive than coal.
c. Fuel costs are minimized; i.e., the system is
optimized, when the boiler system is designed and operated
so that approximately 90% of the fuel value, fed as coal
to the fluidized-bed boiler, is consumed in the primary
cells and 10% in the Carbon-Burnup Cell region.
d. This boiler has shown advantages in control of
S0», NO , hydrocarbons, carbon monoxide and other pollutants,
^ X
e. Calcium sulfate formed in the low temperature
(i.e., coal-burning) regions of the fluidized-bed boiler,
either by use of particulate lime beds or by injection
of fine limestone into inert ash beds, will not decompose
in the high temperature Carbon-Burnup Cell if the residual
oxygen level is maintained at above 3.5%, when no coal
is added to the Carbon-Burnup Cell.
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2-2
f. Decomposition of this calcium sulfate may be
achieved by operation at low levels of residual oxygen,
and/or by feeding coal to the CBC. However, the SO_
concentration achievable with fine lime in a high velocity
apparatus was relatively low, %%%. SO- levels of 3 to 4%
were achieved with the existing CBC using -8+20 mesh lime
bed material but not with 3% O_ in CBC flue gas. The
existing prototype burnup cell is not optimized for
simultaneous use as a regenerator, in the SO- Acceptor
Process MK I configuration.
g. Injection of fine limestone does not appear to
lend itself to a regenerative SO_ control process for a
second reason: no easy method of separating the regenerated
fine lime from the burned-out coal ash is apparent;
reinjection of coal ash to the boiler is undesirable.
h. Using lime beds in either the two-cell or three-
cell systems, simultaneous achievement of 90% or better
capture of SO- in the FBM and 98+% carbon burnup is
relatively easy. Alternatively, 90% SO- capture and
synthesis of a 3% S02 regenerator gas can simultaneously
be achieved. But all three results: high capture,
high burnup and high regenerator SO concentration simul-
taneously required use of the three-cell system during
this program. At the gas velocities in use, simultaneous
achievement of all three was realized using a three-
reactor system in which the high temperature functions
(carbon burnup and S0_ regeneration) occur in separate
zones at different 0_ levels (the SO- Acceptor Process,
MK II). A Carbon-Burnup Cell could be built using lower
gas velocity and larger bed area in which the simultaneous
98% burnup-3% SO- synthesis conditions could be met.
Alternatively, the Carbon-Burnup Cell could operate with a
relatively deep bed and achieve the same results. It is
believed that a two-cell system may be technically feasible,
although experimental proof is required.
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2-3
i. Using the three-cell SO_ Acceptor Process, MK II
regenerator flue gas containing up to 10% SO_ by volume
was achieved. Such a flue gas is believed suitable for
economical sulfur recovery, «'.. SO, synthesis cr lime scrubbing
treatment.
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3-1
3. RECOMMENDATIONS
Near term optimization of the three-cell FBM/CBC/REG
steam generating system is recommended. The regenerator
would operate at low excess O_ levels and return active
bed material to the boiler. It is presently sized to
yield a flue gas containing 5% or more S0_ by volume*.
It is not anticipated that the steady state SO_ level
in the regenerator flue gas will exceed 10%, although
the upper limit value is not firmly known at this time.
The FBM typically captures 90 to 95% of the sulfur re-
leased by coal combustion, using an active lime bed.
The tandem FBM/CBC system will combust 98+% of the carbon
fed in coal, using burnup techniques already established.
In order to realize the air pollution control potential
of the multicell fluidized-bed boiler as rapidly as
possible, the following further actions are recommended:
a. A set of cost and performance goals for fluidized-
bed boilers should be established. This would include
stringent air pollution control goals on SO , NO , hydro-
carbons, CO, halogens, particulates, and plume opacity.
b. Pope, Evans and Robbins, together with one of the
major boiler manufacturers and a public utility, should
perform a detailed engineering desigh--.for a large
coal-fired, multicell, fluidized-bed boiler which may
meet these goals.
c. Based on the questions which arise about that
particular design and related designs, an experimental
program should be conducted to answer those questions.
Operation at over 10% SO- is not visualized.
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3-2
d. If, based on the assessment of an actual design,
EPA's goals will be met, a prototype boiler should be
constructed. A unit capable of producing 10 to 40 MW(e)
would be adequate to provide the needed operating experience
and design data for much larger units.
e. Pending a decision to proceed with the rational
plan outlined above, the following experimental work
on the air pollution control aspects of the multicell,
fluidized-bed boiler should be carried out:
(1) Studies are required to optimize the regener-
ation section of the multicell, fluidized-bed
boiler.
(2) Tests of the SO- Acceptor Process should
be conducted which will determine, for geographically
matched limestones and coals, and various coal sizes,
the required sorbent circulation and makeup rates.
(3) Experiments to indicate the directions in
which reduced emissions of oxides of nitrogen
may be achieved should be initiated.
(4) A definition of the particulate and plume
opacity control requirements of a multicell,
fluidized-bed boiler should be found in coopera-
tion with a leading manufacturer of dust collec-
tion apparatus.
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4-1
4. INTRODUCTION
4.1 Description of a Fluidized-Bed Boiler
4.1.1 The Need for a New Form of Combustion
Since 1962 PER has carried out design and experi-
mental studies aimed at reducing the cost of utilizing
coal as a boiler fuel, initially under the sponsorship
of the Office of Coal Research, Department of the Interior,
and later under EPA sponsorship as well.
These studies were concerned primarily with improve-
ments in the design of plants and boilers for industrial
steam generation; i.e., systems which would be large
enough to supply power and/or process heat to a factory,
but too large to be used by a laundry or apartment house,
and too small to be used by a large electric utility.
Boilers in this size range were selected for develop-
ment for a number of reasons which can be summarized
by a single statement -- a novel boiler could be commer-
cially successful in this size range with less develop-
ment expense. Success at this level could then lead
to scaling both downward and upward.
Conventional methods of firing coal, as a fixed-
bed on a stoker grate, or as a suspension flame in a
pulverized fuel burner, were not found to hold promise
for major reductions in size and cost, regardless of
development effort.
It was found that the high cost of a coal-fired
industrial boiler compared to comparable oil or gas-
fired boilers was primarily due to differences in fur-
nace size. Oil and gas could be burned in a smaller
furnace than could coal. For industrial-sized boilers,
this difference in furnace size meant that oil and gas-
fired boilers could be assembled in a factory and shipped
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4-2
to the user's site by rail, while coal-fired boilers
of equivalent capacity could not. The need arose for
a method of firing coal which would reduce the size of
the furnace so that higher capacity coal-fired boilers
could also be factory-assembled with consequent cost saving,
compared to on-site construction.
An evaluation of alternatives led to the selection
of the fluidized-bed boiler as the most promising method
of achieving the goals of our sponsor. Later, the air
pollution control potential interested EPA.
4.1.2 Defining a Fluidized-Bed Boiler
A fluidized-bed boiler is defined as a system
which meets all of the following criteria:
a. The system's primary function is the genera-
tion of steam. Therefore, the materials of con-
struction, the mode of operation, the arrangement,
auxiliary power requirements, etc., are consistent
with existing practices and economics in the con-
ventional boiler field.
b. The fuel is added to and burned within a
turbulent, aerated bed which has been termed a
"fluidized bed."
c. A significant fraction of the heat released
by the burning fuel is immediately extracted by
heat transfer surface in contact with the turbulent
bed.
A fluidized bed, in turn, is defined as a mass
of particulate solids held in suspension by an upward
current of fluid such that the bed has zero-angle of repose
and exhibits certain other properties of a liquid. Among
the liquidlike properties of a fluidized-bed which are
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4-3
important to the boiler designer, is that the bed becomes
well-mixed, with sufficient agitation, and the bed material
can be caused to flow about or out of the system without
the aid of mechanical devices.
4.1.3 What a Fluidized-Bed Boiler is Not
A number of systems have been conceived in which
partial or complete combustion is carried out within a
fluidized bed, the most successful being the regeneration
section of the fluid catalytic cracker developed for
the petroleum industry. However, unless the system
meets each of the criteria listed above, it would not
be classified as a fluidized-bed boiler.
A fluidized-bed is not an exotic new system re-
quiring the establishment of a new industry such as
required for nuclear power development. The boilers will
be built by the companies now making conventional boilers,
in their existing shops. A few months of experience
will provide a skilled boiler operator with a sufficient
understanding of the new form of combustion to perform
his job.
4.1.4 A Simplified Description of a Fluidized-Bed Boiler
A fluidized-bed boiler consists, in its simplest
form, of an enclosure containing both boiler tubes and
a bed of granular solids. The bottom of the enclosure
is perforated, and air is forced into the enclosure to
fluidize the solids and react with coal which will be
added to the bed. Such a system is shown schematically
in Figure 1. When the temperature of the bed is raised
by an auxiliary means to above about 800 F, bituminous
coal added to the bed will ignite*. The temperature
*The ignition temperature of anthracite was shown to
be well above 800°F.
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4-4
FLUE GAS TO
STACK
\
COAL SUPPLY
FLUIDIZED-BED
COMBUSTION ZONE
COAL
METERING
FUEL INJECTION
AIR
FAN
COMBUSTION
AIR
STEAM
FEEDWATER
>~BREECHING
BOILER TU3E£
IN FORM OK
WATER WALLS
AIR DISTRIBUTION
GRID
PLENUM
CHAMBER
FIGURE 1. SCHEMATIC OF FLUZDIZEO-BED BOILER
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4-5
will then rise until the system achieves a thermal
equilibrium; i.e., the energy added to the bed by the
burning fuel precisely equals the energy extracted by
the boiler tubes touching and viewing the bed and the
gases and dusts leaving the bed. At steady state, the
fluidized bed consists almost entirely of inert particles
with a small quantity of reacting coal.
This equilibrium temperature may be as low as
1200°F or as high as 2500°F, although a narrower range,
1450°F to 2050°F, is of practical interest. Below about
1450°F, it is difficult to completely consume carbon
monoxide, while above 2050°F particles of certain fusible
bed materials may couple and the bed collapse.
The depth of the fluidized bed may also have a
very wide range — from a few inches to many feet. Again,
however, practical atmospheric systems must operate in
the range of 12 to 48 inches. Below about 12 inches,
the combustion efficiency degrades; above about 48 inches,
the power required by the blowers becomes excessive*.
In specifying the operating temperature and bed
depth, the energy release rate still remains unspecified.
Rather than specify an energy release rate, however,
it is more appropriate to specify an easily measurable
equivalent, the mass flow rate of combustion air. The
air rate may vary over the range of 100 Ib/hr ft of
bed plan area to well over 1000 Ib/hr ft of bed plan
Fluidized-bed boilers may be incorporated into com-
bined cycles in which the hot flue gas at high pressure
is passed through a turbine which is used to compress
the combustion air as well as drive an alternator.
In this system the economics are not so sensitive to
bed depth. The pressure cycle leads to important size
reduction possibilities. The reduced size of fluidized-
bed boilers compared to pulverized fuel furnaces is an
important advantage in pressure vessel design. Availa-
bility of turbines designed to run on fly ash remains
limited, however.
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4-6
area. Practical systems, operating at atmospheric
pressure, will operate in the range of 500 to 900 lb/hr ft
2
The coal mass rate is then about 50-60 lb/hr ft .
From a specification of the air rate, the fuel
rate, the bed depth, and bed temperature, it is possible
to compute the apparent volumetric heat release rate,
the superficial gas velocity and residence time and other
parameters of interest. The air rate selected determines
the size and density distribution of the particles which
will make up the fluidized bed. Particles above a certain
maximum size and density will sink to the air distributor.
Particles below a certain minimum size and density will
tend to be entrained in the gases leaving the bed and
be carried out of the system*. The fluidized bed, in
addition to being an averaging device, naturally selects
those particles which it wishes to retain and rejects
those whose properties fall outside the desired range.
The particles which make up the bed may be any non-com-
bustible, granular solid which is sufficiently tough**
to retain its shape and size in a bed over an extended
period. In some coals, the adventitious or "non-inherent"
mineral matter leaves an ash which meets these criteria
and the coal can be burned in a fluidized-bed composed
of its own ash. For many coals, however, the ash is
too fine to be retained in a fluidized bed unless it is
deliberately sintered.
* It appears characteristic of beds which are not
monodisperse that fines added to the bed do not
elutriate immediately but rather a pulse of fines
is emitted with a decay curve corresponding to a
zero order process.
** Methods of quantatively specifying "toughness" were
beyond the program scope.
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4-7
For these coals, a material must be added with
the coal to make up for particles of starter bed which
are lost to the system. This material, besides meeting
the properties listed above, must also be inexpensive.
In many areas, limestone may be the material of choice,
because of its cost and relatively low density after
calcination, which permits a relatively deep bed to be
used with a moderate pressure drop through the bed.
Limestone also possesses other properties of interest
to the designer of a fluidized-bed boiler and these are
discussed below.
4.2 Air Pollution Control Potential of Fluidized-Bed
Boilers
4.2.1 What is Air Pollution?
The question is not facetious, for it has been
seriously suggested that all the products of combustion
of a fossil fuel constitute air pollution. Disasters
have been predicted because carbon dioxide may increase
to levels at which the earth's solar energy balance and
global ecology are disturbed. However, the CO« absorp-
tion capacity of seawater and vegetation appear unlimited.
While water-vapor release into the lower levels of the
atmosphere has not been judged to be dangerous, the steam
plume which is visible atop many stacks on cold, humid
days arouses numerous calls to a community's air quality
wardens.
The nitrogen which leaves the stack of a large
power station at temperatures above ambient may affect
the micrometeorology of the plant site. It is of value
to sailplane enthusiasts. However, most responsible
authorities would not classify carbon dioxide, water
vapor, and warm nitrogen as pollutants. Classified
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4-8
(by law) as pollutants are products of incomplete combus-
tion (carbon monoxide and a variety of hydrocarbons),
gaseous oxides of sulfur and nitrogen, gaseous halides,
and particulates. To these might also be added heavy
metal vapors and natural radioactive isotopes in the
40
particulate (e.g., K ).
4.2.2 Pollution Control Potential of Fluidized-Bed Boilers
4.2.2.1 Discussion
4.2.2.1.1 General
Fluidized-bed combustion of coal is, in itself,
not remedy of any of the pollutants. A fluidized-bed
boiler is not per ae a pollution control device. However,
certain properties of fluidized-bed combustion and of
a properly designed fluidized-bed boiler can be exploited
to produce a steam supply which is "clean" at less cost
than available alternatives.
By careful design, partial combustion products
can be consumed within the fluidized bed and in the
freeboard. Although no work appears to have been done
on methods of reducing CO and C H (beyond decreasing
the coal: air ratio), emissions on the order of 200 ppm
or less for each can be anticipated. Additional research
possibly a search for low cost combustion catalysts,
may provide methods for a further reduction in emissions.
Such an approach may be economical in a fluidized-bed
boiler, but is less likely to be practical for conven-
tional boilers.
4.2.2.1.2 Sulfur Oxides
By the use of an attrition-resistant limestone*
as the bed material, it is possible to absorb virtually
Some limestones have been found to possess inadequate
attrition resistance. Note that hardness and attri-
tion resistance are not synonymous.
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4-9
all the sulfur released by the coal. To achieve very low
levels of emission, a very large excess of active lime
is required in the bed and freeboard at all times*.
Economics, in turn, require that the limestone be kept
active by continuously stripping off much of the sulfur
in a separate zone so as to regenerate the limestone.
The off-gas of the regenerator can then be processed to
recover sulfur or sulfur products or be scrubbed, using
a portion of the lime produced by the boiler, or CBC
fly ash which is naturally rich in lime. The scrubber
effluent containing CaSO. must then be disposed of.
Elemental sulfur is difficult to obtain by the use of
alkaline earths as the principal sulfur acceptor. The
use of alkali-based sorbents may prove more effective.
The presence of less excess sorbent may be feasible
and elemental suflur may be more readily obtained.
The property of a fluidized-bed boiler which makes
effective in situ sulfur control possible is the relatively
low temperature (about 1500°F) of the medium in which
the combustion occurs. The combustion bed may be kept
cool by immersing heat exchange surfaces therein, or,
alternatively, by circulating the sorbent bed through
a separate heat exchanger. In addition, the refluxing
Pressurized boiler operation aids SO- capture but
poses additional regeneration problems. For example,
pressure favors formation of other calcium-sulfur
compounds:
CaSO. + S03 -»
CaS04 + SO2 -»
Thermochemical data are lacking for the pyrosulfate
and dithionate type materials. Attempts to regenerate
bed material may form gaseous sulfur oxides which
then are resorbed by the reactions shown. Current
difficulties in decomposing sulfated lime under pressure
may be due to this mechanism.
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4-10
of sorbent particles in the freeboard provides a degree
of concurrency to an otherwise well-stirred system.
Additional development to provide countercurrency in a
multi-bed device (stacked beds) with bed regeneration
may provide emission control of the order of 99+%.
4.2.2.1.3 Nitrogen Oxides
Nitrogen oxides are produced in any boiler through
the fixation of atmospheric nitrogen and oxidation of
nitrogenous compounds in the coal*. Because of the
relatively low temperature of the combustion medium
("v>1450-1600°F) , it had been anticipated from equilibrium
data that nitrogen oxides from a fluidized-bed boiler
would be very low when compared to conventional boilers.
However, this has not been found to be the case. Measured
values of NO from some benchscale fluidized-bed combustors
have approached levels reported for the "worst conventional
boilers"; measurements of NO in which tend to be on the
low side due to the reactivity of NO in sample systems.
While little work has been done on methods of
reducing NO emissions from a fluidized-bed boiler, the
techniques suggested for conventional boilers provide
a starting point. One such technique is the recircula-
tion of flue gas. While this method may lead to unstable
combustion in the conventional boiler, this problem
will not occur in a properly designed fluidized-bed boiler.
Experiments we ran (2) in which less than the normal supply
of oxygen was sent through the fluidized bed, reduced
nitrogen oxide emissions to below 200 ppm.
Argonne National Laboratories, in a well-designed
experiment utilizing an "artificial air" where argon
was substituted for nitrogen, demonstrated that
nitrogen oxides can arise from the nitrogenous compounds
in the coal.
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4-11
Catalysts which promote the reduction of nitric
oxide to nitrogen in the presence of reducing gases might
also be applied when a fluidized-bed heat exchanger follows
a fluidized-bed combustor. A search for catalysts which
can be produced from the mineral matter in coal is a
promising area for research. A goal of 50 ppm for NO
would be set for this approach.
4.2.2.1.4 Other
While the mineral matter in coal may prove useful
in controlling nitrogen oxides, it is usually considered
a source of pollution; i.e., particulates. Here the
fluidized-bed boiler possesses certain inherent advantages
over most conventional boilers. Since the coal is crushed
rather than finely ground, much of the adventitious mineral
matter remains in the bed and is not entrained in the
gas. Particles which are entrained may be removed by a
relatively inefficient and inexpensive mechanical dust
collector, compared to the requirements imposed by pul-
verized coal firing.
Inherent ash, on the other hand, may be very fine
and will be entrained by the gas regardless of coal
particle size. Some of this will pass through a mechanical
collector and must be removed by an electrostatic precipi-
tator. This final collection may be difficult when there
is little sulfur trioxide in the gas, due to lime bed
operation, since the electrical resistivity of most of
the ash matter (silica, alumina and ferric oxide) is high
at conventional collection temperatures. However, in a
fluidized-bed boiler in which a Carbon-Burnup Cell is
applied, the carbon content of the fly ash from the
primary combustor (>40%) should act as a natural condi-
tioning agent. This would allow efficient dust removal
at conventionally low temperatures from the gas leaving
the 1° cells.
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4-12
The conventional measure of particulate emissions
(as well as gaseous emissions) has been weight of pollutant
per unit energy input. Emissions of 0.02 grains/standard
cubic foot, the most stringent standard yet proposed, are
the equivalent of 0.04 lb/10 Btu of fuel input. This
should be achievable from a fluidized-bed boiler equipped
with a high efficiency electrostatic precipitator.
A second standard for particulates is concerned
with the opacity of the plumes issuing from the power
plant's stacks. Standards now proposed will require
that the emission be "invisible". While the opacity of
the plume is related to the weight of particulates, it
is not congruent with weight. A relatively opaque plume
may result, for example, when firing conditions result
in incomplete combustion causing micron-sized carbon
particles to form from cracked hydrocarbons. Gas turbines
and many incinerators smoke through this mechanism.
This will not occur in a properly designed fluidized-bed
boiler. A second source of light-scattering particles
is sulfur trioxide droplets formed as the dew point is
reached in the gas stream. When the bed of the fluidized-
bed boiler contains limestone, the quantity of sulfur
trioxide has been shown to be vanishingly small.
A third source of small particles is the inorganic
fumes and smokes which result when mineral matter is
volatilized during combustion. Because of the low tem-
perature of the combustion medium in a fluidized-bed
(^15.00 F) boiler, some of the mineral matter may not
volatilize, and a fraction of that volatilized out of
the coal matrix will condense on the bed material and
fly ash. As much as 99% of certain volatile species
could be retained in a fluidized-bed operating at
about 800°C.
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Chlorine, fluorine, arsenic and selenium are
four minor or trace constituents in coal which may be
released in flue gas. Other trace constituents include
sodium, potassium, phosphorous, titanium and manganese.
While these have not yet been considered for control
from coal-fired power stations, it may be anticipated
that, as other pollutants are brought under control,
attention may be turned to these. A fluidized-bed boiler
can be designed to control even these pollutants.
Chlorine and fluorine might be controlled by adding some
potassium carbonate to a fluidized-bed combustor. Arsenic
and selenium will be quantitatively removed by lime when
present in the combustor bed, along with part of the
chlorine and fluorine.
Coal, like seawater, contains every element in
the periodic table below Z=92, and every natural isotope
of these elements*. Since, on a worldwide basis, about
g
3x10 tons of coal are burned each year, tonnage quanti-
ties of even the trace constituents enter the atmosphere
from the stacks of coal-burning power plants. Fortunately,
the majority of these elements add little to the natural
background levels and are of no concern.
Some elements are present in sufficiently high
concentration and are sufficiently toxic to be of con-
cern. Among these would be barium, cadmium, lead, and
mercury**. The natural radioactive elements—radium
for example—might be added to this list. It would be
anticipated that a large fraction of the lead and uranium
would be tied up by lime. Cadmium and mercury may be
*Seawater now contains some of the higher elements
such as plutonium.
** Up to 25 ppb Hg in bituminous coal.
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4-14
sufficiently volatile to appear in the flue gas, un-
combined, and add measurably to the background concen-
tration downwind of the plant stack. For each element,
with the exception of mercury, the enormous surface area
and low temperatures of a fluidized-bed boiler or heat
exchanger may reduce the quantity of these elements which
enter the atmosphere in dangerous forms. A systematic
search for methods of removing cadmium and mercury from
flue gas is indicated. Possibly the fluidized-bed boiler
may prove useful in applying the method economically.
Even if the undesirable elements are fixed in the ash,
they may reappear in the biosphere if the methods of
ash disposal are inadequate. There is no evidence today
that ash from fluidized-bed boilers operating with lime
beds will be utilized in ways that fix the soluble consti-
tuents*. No investigations have been made of the poten-
tial uses to which the coarse stream (impure lime) or
the fines stream (coal ash containing partially sulfated
lime) from a fluidized-bed boiler could be put.
However, because these streams have not been vitri-
fied by exposure to high temperature, there is some hope
that the ash of a fluidized-bed boiler may be a useful
raw material for processing into commercial materials
in which the impurities become permanently fixed.
4.2.2.2 Conclusions
It is seen that a fluidized-bed boiler may be
designed to reduce air pollution to low levels. The
pollution control potential follows from certain inherent
characteristics: the low temperature of the bed, the low
e.g., CaSeO. solubility 9 parts per 100.
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4-15
temperature of the combustion, the high carbon content
of the 1° cell fly ash, the stability of the combustion,
and the general properties of a fluidized-bed. While
relatively little work has been done on the control of
pollutants other than sulfur oxides, the development of
a system in which pollution control and economical steam
generation are simultaneously optimized appears to be
an attainable goal.
4.3 Pope, Evans and Robbins' Prior Work
4.3.1 General
Pope, Evans and Robbins, under contract with the
Office of Research and Monitoring of the Environmental
Protection Agency (successor to the National Air Pollution
Control Administration, Department of Health, Education,
and Welfare), has characterized emissions from a fluidized-
bed boiler developed for the Office of Coal Research,
Department of the Interior. Three reports have been
prepared describing this work (References 1, 2, and 3).
The pollution control aspects of a fluidized-
bed boiler had been considered as early as 1965, when
it was discovered that if coal distribution were uniform,
smdceless combustion was achieved at attractively low
excess air levels. Early in 1966, a test was conducted
in which dolomite was mixed with the coal, and the
sulfur oxide control potential of the process was demon-
strated. A literature and patent search revealed that
this approach to the control of sulfur oxides from a
coal-fired boiler was novel, though a patent had been
granted earlier on the use of limestone in a shale-burning
process.
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4.3.2 Once-Through Limestone Injection
Based on this early work, the studies conducted
for EPA first investigated the use of relatively coarse
limestone and dolomite.
It was determined that coarse stones would not
retain more than about 30% of the sulfur released by the
coal at reasonable Ca/S ratios. A "sulfate shell" theory
was invoked to explain this.
A dolomite identified as 1337 was found to decrepi-
tate rapidly and would not be retained in the bed. A
limestone, 1359, was retained in the bed but could not
be fully converted to CaSO. if sufficient stone were
added to significantly reduce sulfur oxide emissions.
An unexpected result was that the SO- capture
rate decreased with increasing bed temperature. Bench
scale studies by others using simulated flue gas mixtures
had indicated more rapid capture of S0_ at 1800 F than
at 1600°F, yet the SO_ capture process, which was then
assumed kinetically limited as well as irreversible,
performed far better at 1600°F than at 1800°F. It was
then understood that proper simulation required combustion-
generation of SO™ within the sorption bed.
In order to enhance the process rate by increased
surface area, the stone was finely ground. Despite the
bed's inability to retain indefinitely particles under
about 30 U.S. Standard Mesh*, the efficiency of sulfur
capture increased with decreasing particle size, implying
among other things that the finest particles calcine faster.
As noted earlier, economically feasible fluidized-
bed boilers operate at air rates between about 400
and 1200 Ib/hr per- ft2 of bed surface which at 1500°F
corresponds to superficial gas velocity of 5.6 to
16.7 feet per second.
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It was determined that about an 80% reduction in SO2
emissions could be achieved by injecting fine 1359
limestone at Ca/S = 2.5. Raw stone and hydrate both
performed equally well, while stone precalcined by
the supplier performed poorly. The behavior of the
precalcined stone was considered anamalous, since it
would have been expected to be better; perhaps it was
dead-burned.
4.3.3 Regenerative Limestone Process
Toward the end of the pollution characterization
effort (May 1969), it was discovered that beds composed
almost entirely of coarse limestone could be made to
release the accumlated sulfur that had been retained
in a batch operation by increasing the coal-feed rate
so as to increase the bed temperature and decrease the
oxygen content of the flue gas. A regenerative cycle
was devised and a patent application prepared for the
"SO2 Acceptor Process."** This discovery was felt to
provide an explanation for the anomalous temperature
behavior noted earlier. Sulfur retention in a fluidized-
bed combustor was now seen as a reversible process
sensitive to temperature and oxygen partial pressure.*
The reversibility at the temperatures of interest exceeded
thermodynamic predictions and may depend upon contamin-
ation of the CaO by coal ash ingredients. Alternatively,
localized reducing conditions might be responsible.
Pope, Evans and Robbins also demonstrated in a
coal-fired, fluidized lime bed combustor that the injection
of small amounts of sodium chloride greatly increases the
sulfur removal capacity of the lime. We also determined
* It was not determined whether the apparent effect
of oxygen was direct, or whether the effect of oxygen
on the reducing gases present in the bed governed the
retention-release of sulfur.
** U.S. Patent 3,717,700, February 20, 1973
-------
4-18
in bench scale experiments that salt dissolves CaSO.
at the temperature of the fluidized-bed combustor.
This activity enhancement effect was thought to be due
to the removal of the "sulfate shell" from the lime
particles, increasing the activity of calcium oxide.
See Appendix E for further background information.
4.3.4 Pollutants Other Than SO,,
Pollutants other than sulfur dioxide were charac-
terized—nitrogen oxides, hydrocarbons, and particulates.
Oxides of nitrogen were found to be dependent
on the coal: air ratio; i.e., nitrogen oxides would
increase as the amount of oxygen remaining in the flue
gas increased*. It was not determined if this resulted
because more nitrogen was fixed, or because less nitric
oxide was decomposed as the reducing gases in the bed
were consumed. However, experiments in which some air
was diverted from the base of the bed to a port above
the bed resulted in a marked reduction in NO. This
might indicate that the reducing gases in the bed were
decomposing the NO. These results might also have
indicated that conditions near the air distributor were
governing. Tests with a variety of distributor designs
were not conducted. Nitrogen oxides were not materially
and reproducibly affected by the injection of fine lime-
stone, although Corkers at Argonne have shown a reduction
in NO with limestone injection (Reference 4). An
increase in NO with limestone injection has also been
predicted on the basis of experiments at Esso Research
and Engineering (Reference 5).
A dependence of NO output on coal particle size has
also been observed. This dependence on coal size tends
to suggest that at least some NO is due to thermal
fixation of N_ from air because the large coal particles
are known to Be hotter then the rest of the bed particles,
-------
4-19
NO levels of less than 400 ppm were found in
PER's boiler. 275 ppm was typical.
Hydrocarbons and, presumably, carbon monoxide*
were found to be sensitive to the quality of the fuel
distribution and to the coal:air ratio. With between
3% and 4% oxygen remaining in the flue gas, the hydro-
carbons were reduced to below 100 ppm. Hydrocarbons
were not affected by fine limestone injection.
The major fraction of the mineral matter in
the coal appears as fly ash. The larger particles of
adventitious matter remain in the bed. The bed material
itself adds some particulate matter to the flue gas
and when fine limestone is injected, essentially all
of this appears in the fly ash stream.
At the high dust loadings, when fine limestone
injection was used, an inexpensive, low pressure-drop,
mechanical collector proved remarkably efficient
(^90% removal).
The particles not removed by the collector were
all smaller than 20 ym. These particles would have
to be removed by an electrostatic precipitator, bag
collector, or wet scrubber. Because of the high carbon
content (^40%) , the resistivity of the FBM fly ash may
be sufficiently low to permit efficient electrostatic
collection at low gas temperatures.
There was no reason to believe, on the basis
of tests performed, that a fluidized-bed boiler would
be unable to comply with the most stringent regulations
governing particulate emissions at equal or lower cost
than any other boiler.
HC was recorded continuously. CO was determined
intermittently by Orsat and was not over 0.1%, the
limit of detection; no continuous record available.
-------
4-20
4.3.5 The Carbon-Burnup Cell
4.3.5.1 Tests in the Fluidized-Bed Column
Development of staged combustion was required
to reduce ash carbon content to the 0-10% range. The
major portion of the effort to produce the design
correlations of fly ash combustion tests was carried
out in the Fluidized-Bed Column (FBC) which, during
these tests, had a plan area of 0.86 square feet.
Fly ash was burned in this device over a wide
range of conditions, summarized as follows:
Bed temperature: 1750°F to 2140°F
Bed depth (static): 10" to 22"
Air rate: 385 to 1000 Ib/hr ft2
Fuel rate: 48 to 350 Ib/hr ft2
Carbon concentration
of fuel: 28% to 65%
Heat removal rate: 15% to 40% of heat release
The key result of the test program was this
performance model, which may be used to predict the
combustion efficiency in a Carbon-Burnup Cell from the
parameters which were found to control performance.
This model equation is as follows:
Combustion efficiency, % = (1)
-13.78
+0.05193 (bed temperature, °F)
+0.03973 (air rate, Ib/hr ft2)
+0.3831 (static bed depth, inches)
2
-0.7514 (carbon feed rate, Ib of carbon/hr ft )
-0.1638 (inert feed rate, Ib of inert/hr ft2)
+0.0020 (carbon feed rate x inert feed rate,
Ib2/hr2ft4)
-------
4-21'
Tests of this model, using data obtained in runs
not used in the model derivation, indicated that the
model was not limited to simply reproducing itself but
was a useful prediction -tool.
The residual oxygen content of the flue gas could
also be predicted, although somewhat less accurately
than the combustion efficiency. This model equation
is as follows:
Residual oxygen,'% = (2)
22.91
-0.007353 (bed temperature, °F)
+0.01118 (air rate, Ib/hr ft2)
rt
-0.1390 (static bed height, inches)
-0.1521 (carbon feed rate, Ib carbon/hr ft )
-0.0151 (inert feed rate, Ib of inert/hr ft2)
+0.0002653 (carbon feed rate x inert feed rate,
Ib2/hr2ft4)
A parametric study was performed to determine
the optimum split in duty between the primary cells
and the Carbon-Burnup Cell. The lowest fuel costs
are realized if the Carbon-Burnup Cell burns about
10% of the input fuel. This is equivalent to saying
that the primary cells should operate with about 3%
residual oxygen in the flue gas. This coincides with
a good trade-off between higher hydrocarbon at low O-,
and lower thermal efficiency at high excess air.
When finally divided (-325 mesh) No. 1359 lime-
stone was injected into the FBM in this series of tests,
less sulfur was removed than in the tests described in
Reference 2 (^70% at Ca/S = 2.5 compared to -^80% at
Ca/S = 2.5 measured previously). The reasons for the
less favorable performance are unknown but may have been
the result of a different injector design.
-------
4-22
When partially sulfated fine limestone entered
the Carbon-Burnup Cell along with the carbon-bearing
fly ash, its sulfur content could be released as SO_ if
the CBC operated at a temperature of ^2000°F with residual
oxygen _<_ 2.0%. The highest S0_ values, ^5000 ppm,
were measured when the residual oxygen level was ^0.2%.
On the other hand, when the residual oxygen in the Carbon-
Burnup Cell was in the range 3.5 to 6.0%, the fine sulfate
did not decompose at CBC temperatures. In fact, in one
test at 6% O_ , the partially sulfated limestone appeared to
still be reactive at a bed temperature of 1980°F. An
alternative explanation for this result appears to involve
dynamic exchange of sulfur*. An important result was
that the sulfur oxide emission from a Carbon-Burnup
Cell may be as low as 350 ppm.
When no sorbent was present, the sulfur in the
fly ash would burn with an efficiency equal to that
of the carbon, or greater.
The data gathered on nitric oxide emissions from
a Carbon-Burnup Cell were correlated less accurately
than other parameters, though the highest emissions
were detected at the highest bed temperatures. A mean
value of 539 ppm was measured for all fly ash combustion
tests in which the bed temperature was above 1900 F**.
No efforts were made to reduce nitric oxide emissions and
this remains one of the major areas requiring research.
It was noted that the addition of coal to the fly ash
feed would reduce the NO level, suggesting one line
of potential research.
'More about this in Section 6.2, page 6-20.
** At typical boiler operating conditions, about one-
third the total NO emission is CBC; two-thirds FBM.
-------
4-23
Particulate emissions for a Carbon-Burnup Cell
were found to decrease with increasing bed temperature,
possibly through agglomeration of ash matter to the
sintered ash bed particles.
Considering that in some tests the fuel was
^65% ash, particulate emissions were remarkably low.
The pollution control potential of retaining a large
fraction of the ash in the bed of a Carbon-Burnup Cell
is an avenue of research which should be explored.
Hydrocarbons and carbon monoxide emissions were
essentially nil when operating in the Carbon-Burnup
Cell mode.
4.3.5.2 Tests in the Fluidized-Bed Module
Some tests were conducted in the Fluidized-Bed
Module (FBM), an actual boiler with a grate area of
-^9 ft2.
A simulated Carbon-Burnup Cell was appended to
the FBM to determine the problem areas in operating a
fluidized-bed boiler with two distinct temperature regions,
This device was designated the CBC. The low freeboard CBC
design and was necessitated by the FBM steam drum. It
2
was found that for the small CBC (^1.1 ft ) an opening
in the barrier between the two regions of just 2 square
inches permitted the desired temperature difference
to be achieved. An interchange rate of 12,000 pounds
of bed material per hour per square foot of opening was
estimated from two transient heating tests.
A new coal feeder design also added at the same
time performed very poorly until a number of minor
alterations corrected the problems. Coal feeding to
a fluidized-bed boiler is an area which requires further
development effort.
-------
4-24
Two fly ash feeders were tested in the CBC. One
design, termed a mushroom feeder, performed well, giving
a relatively even fuel distribution.
Two bed particle "knockouts" were also tested
in the CBC. These both consisted of cylinders arranged
in a triangular array. Due to the inadequate freeboard,
without a baffle screen, large quantities of bed material
had been expelled from the CBC; with either baffle screen,
the quantity of bed material carried over was small.
When the baffle screen was made up of water-cooled tubes,
an initial heat transfer coefficient of ^35 Btu/ft hr°F
was measured. Here the temperature refers to the
fluidized-bed temperature. This was approximately twice
the value that would be predicted from simple radiation
and convection, indicating that an active heat transfer
region exists above the dense phase of a high velocity
(hence turbulent) fluidized-bed.
When all operational problems had been overcome,
the FBM/CBC achieved an overall combustion efficiency
of approximately 99%.
4.3.6 Recommendations Based on Prior Worjc
The program was ended with recommendations for
further study of the most promising results:
(1) A study of the regenerative mode of utilizing
a limestone sorbent. If the sorbent were circulated
rapidly enough, it appears possible to achieve low
emission levels from the acceptor region of the boiler
and relatively high concentrations from the regenerator
region, suitable as feed to a sulfur recovery operation.
-------
4-25
(2) A study of reducing NO emissions by disturbing
X
the oxygen gradients within the fluidized bed. Recircula-
tion of flue gas, tmrning natural gas and coal together,
and studying tne effect of air distributor design were
recommended.
(3) Modification of limestone's sulfur capture
capability by additive injection.
The first and third suggestions formed the basis
of the experimental work conducted in the period November
1970 through July 1971, which is discussed in this report.
4 . 4 Specific Objectives of This Work
As noted in Section 1, SUMMARY, the two primary
objectives of this work were to (a) investigate the
pollution control potential of the concept of coal
combustion in a fluidized bed of lime particles, with
continuous regeneration of the sorbent by circulating
the partially reacted lime between the primary combus-
tion bed and a connecting high temperature zone in which
carbon bearing fly ash from the primary bed is burned
(i.e., the 2-cell system); and (b) investigate the
potential of salt injection into a fluidized-bed coal
combustor as a means of increasing the sulfur sorption
capacity of limestone, as well as combustion efficiency.
The major operating, design, and economic factors
of interest in the "SO- Acceptor Process" included:
investigations of interaction of superficial velocity,
temperature, fly-ash carbon burnup, and SO- production;
verification of better than 90% sulfur capture using a
better than 4% sulfur coal; demonstration of 98+% carbon
burnup in the system; demonstration of high (3-4%) SO_
concentration in the gases off the regeneration zone;
consideration of rate of makeup limestone required;
verification that discard of used bed material is uneces-
sary for SO_ capture maintenance; and superheat production
-------
4-26
from the burnup cell flue gas cooler. The approach taken
to accomplish these objectives was to first conduct
batch nonregenerative limestone bed and salt tests in
the FBC on a small scale; then to arrange components
and demonstrate the regenerative process in long duration
tests on a larger scale in the FBM/CBC. It was not
found practical in the limited time available to study
the effect of CBC bed levels different than the FBM
bed level; nor was it practical to vary CBC grate area
to collectively optimize gas velocity, carbon burnup
and regenerated SO_ concentration. Testing was also
to be performed in a three-reactor FBM/CBC/REG con-
figuration: the SO- Acceptor Process MK II.
The effects of the burnup cell and regeneration
functions on the design and economics of the PER modular
fluidized-bed combustor in the 30MW and 300MW sizes
were to be identified.
-------
5-1
5. APPARATUS
5.1 Pilot Scale Combustor, FBC
The small scale nonregenerating tests were conducted
in a pilot scale combustor, designated the FBC for
"Fluidized-Bed Column." The FBC consists of a rectangular
combustion space, 12" x 16", having an air distributor as
shown in Figures 2 and 3. In operation, air at ambient
temperature, compressed by two blowers in series, enters
a plenum below the air distributor, and passes up through
a grid of buttons (bubble caps) and into the combustion
chamber where it fluidizes the bed material and provides
the oxygen for combustion. Fuel, coal and/or fly ash
is pneumatically injected through a port at the base
of the bed.
The air distributor contains a matrix of grid
buttons mounted in a mild steel plate. The buttons
are 303 stainless steel and designed to direct the air
slightly downward toward the grid plate. This initial
downward flow tends to eliminate stagnant areas. A
cross section of a typical button is shown in Figure 4.
Obviously, this is an anti-sifting design.
To reduce the heat loss to the waterwalls of the
column, and thereby to allow the study of bed temperature
effects independently of bed height, the unit was insul-
ated internally as shown in the partial cross section of
Figure 5. The water-cooled hood (used in CBC simulation)
was insulated in a similar manner. By this means, high
bed temperatures, in the 1800°F to 2100°F range, can be
achieved with relatively deep beds (10"-22"). By use
of water-cooling surface in the form of bayonets, also
shown in Figure 5, the temperature of the bed can be
adjusted by raising or lowering bayonets. It was found
in coal combustion tests that considerable carbon fouling
-------
5-2
LIMESTONE
FEED
SCREW
COAL
FEED SCRE'
ASH RECIR
CULATION
PORT
PLENUM
CHAMBER
WELDED SEAM DUCT
SIGHT PORT
WATER COOLED HOOD
WATER JACKETS
KAOWOOL GASKET
THERMOCOUPLE PORTS
WATER WALLED COLUMN
AIR
PROPANE
LIGHT-OFF
GAS BURNER
FUEL INJECTION
AIR LINE
2 AUXILIARY FEED PORT
(ONE RIGHT, ONE LEFT)
COMBUSTION
AIR INLET
FIGURE 2. FLUIDIZED-BED COLUMN (FBC) CONSTRUCTION DETAIL
FRONT VIEW
-------
5-3
WATER WALLED J.
COLUMN
COAL
FEED SCREW
LIMESTONE
FEED -
SCREW
INJECTION
AIR LINE
FUEL
INJECTION
PORT
. INLET
SIGHT PORT
WATER COOLED HOOD
LIGHT-OFF BURNER PORT
THERMOCOUPLE PORTS
WATER WALL
AUXILIARY FEED
PORTS (2)
PLENUM CHAMBER
\k - • » •» v • •>..»»•» . * •» •?•_«. *. » *.». ^ • . . ».
- •» • • i •*"•>•'. •»«*-.•••»• r '.» » •••-•»-•. .> •_. •>> ,
FIGURE 3. FLUIDIZED-BED COLUMN (FBC)
CONSTRUCTION DETAIL-SIDE VIEW
-------
5-4
45'
0.078" dia.
(Typical for
45'
TOP VIEW
SIDE VIEW
FIGURE 4. ASK DISIfBIBUTION GRID BUTTON
-------
5-5
COOLING
WATER IN
1 —
1
I
1
(fll
II
ll
II
'1
ll
II
ll
II
1
u
1
^ OUT
COOLING }
SURFACE 1
TRAVEL '
!
i
ft
II
II
' |
I
H
>l
I
I
1
'
1
1
1
1
U
/•
JT
A
f
*'*
DJUSTA
jL
^^
., •
^
I ANNULAR SPACE
FILLED WITH KAOWOOL
WATER JACKET
^ LINER OF ASTM 446
NOTE
INSULATED LINER AND
ADJUSTABLE COOLING SURFACE
PERMIT OPERATION AT
TEMPERATURES ABOVE
1800° F WITH BEDS 22 DEEP
FIGURE 5. SECTION THROUGH FLUIDIZED-BED COLUMN SHOWING
INSULATED STEEL LINER AND ADJUSTABLE COOLING
SURFACE
-------
5-6
of the water-cooled bayonets occurred, due to condensa-
tion and pyrolysis of coal volatile constituents. This
did not occur when the fuel was carbon-bearing fly ash.
The insulation consists of a sleeve or liner of
ASTM 446, one of the most refractory steels, backed with
1" of Kaowool, a refractory insulation. With the insul-
ated sleeve, the FBC has an internal cross section of
0.86 ft2.
The bed material consisted of 1359 limestone crushed
and screened to -8+20 U.S. Standard sieve size.
The bed is heated to coal ignition temperatures
with a premix gas burner flame directed downward into
the bed, as shown in Figure 2. The ignition procedure
involves fluidizing the bed material with minimum air
flow, raising the bed temperature to 800 F, and then
injecting coal until the combustion is self-sustaining.
About ten minutes are required for the ignition procedure.
The bed temperature is monitored with several
thermocouples spaced vertically in the combustor. Kaowool
seals were provided to prevent flue gas leakage out of
the system. Specifications for the FBC are presented
in Appendix A.
The fuel feed system is capable of delivering
^3.5 x 10 Btu/hr. The air feed system is capable of
delivering oxygen sufficient for a heat release of
^2 x 10 Btu/hr. These are both in excess of the actual
operational values.
The FBC test system is shown schematically in
Figure 6. Combustion products from the FBC pass through
a heavy gauge welded seam duct, through an optional
induced draft fan*, through a dust collector and on to
The induced draft fan was not used in this test
series.
-------
•SPRING SUPPORT
FROM CEILING
i.D- FAN
BYPASS
,-12" WELDED
\y ROUND DUCT
CONTINUOUS ANALYZER
SAMPLING POINT
* I.D. FAN
-COAL AND FLY ASH
FEED SYSTEMS
*-I.D. FAN NOT USED IN THIS TEST SERIES
FLUE GAS
"isch
PARTICULATE
I SAMPLING PT.
n
6 WELDE
DUCT
4 TUBE i
VENTURI
PITOT
GATE VALVE
iDUST
(COLLECTOR
I BAG
IDUST
RECYCLE
I LINE
WET TEST
ANALYZER
SAMPLE POINT
FORCED DRAFT FAN
FIGURE 6. SCHEMATIC OF FBC AIR AND EXHAUST GAS DUCTING SHOWING SAMPLING
POINTS
-------
5-8
the analytical system and stack. The slanted configur-
ation of the duct between the FBC and dust collector
provides gas cooling without causing wall surface
temperatures to fall below the dew point of sulfur
trioxide (^360 F). A control damper may be used to
adjust back pressure on the system.
Fluidizing combustion air, provided by two series
blowers located external to the test area, was monitored
both by a pitot tube and a venturi meter located in the
long entrance duct. A gate valve in the line was used
to control air flow to the unit.
The fuel and additive feed rates were controlled
by variable speed drives on the feed screws. The
pneumatic fuel feed system was capable of feeding 250
Ib/hr. Collected dust from the overhead cyclone was
discharged into bags and weighed. A dust recirculation
system, indicated in Figure 6, was available but was
not used in this test series. Locations of thermocouples
are described in Section 5.5 - Instrumentation. Salt
feed into the pneumatic coal feed stream was by screw
feeder or in some tests by metered aqueous solution
injection. An isokinetic sampler is provided for studying
stack particulate emissions.
5.2 Full-Scale Boiler Module, FBM
The full-scale boiler module, designated the FBM,
is a boiler unit capable of generating steam under pres-
sure up to 300 psig. In this unit, the fluidized bed
is contained in a rectangular enclosure in which each
wall is a row of vertical boiler tubes seal-welded so
as to form a gas-tight enclosure. The FBM represented
one half cell of the first multicell, fluidized-bed
boiler concept developed under contract with the Office
of Coal Research. Two modules placed back to back would
-------
5-9
comprise one cell. A number of cells placed side by
side without intervening insulation would have made
up a full-scale boiler.
A cut-away sketch of the FBM as it was at the
beginning of the program is provided in Figure 7. A
simulated Carbon-Burnup Cell (CBC), added at the rear
of the FBM, is also shown in Figure 7. The CBC is
discussed later. The FBM cross section is ^18 x 72
inches, 9 square feet, roughly seven times the unsleeved
FBC cross section. The bed is surrounded by vertical
boiler tubes which extend from two cross headers below
the grid plate to the steam drum. No other tubes are
placed in the bed. The boiler tubes are joined together
by welded fins and are backed by insulation. The fins
do not extend the full height so that flue gas passes
between the tubes at the top of the unit and around
the steam drum.
The combustion space is accessible through a
water-cooled panel at the front of the unit. The panel
contains a view port and a premix gas burner used to
fire the bed. The burner directs a flame downward onto
the front of the bed. Two pneumatic feed ports are
provided below the access panel, one for an optional coal
feed or flyash reinjection tube, and the other for the
makeup bed material feed tube. Four optional pneumatic
feed ports are provided at the bottom of one side of
the FBM, shown in Figure 9. For this test series, a
vertical, split coal feeder, shown on Figures 7 and 9,
was utilized. Other feeder designs have been used in
other programs.
-------
FBM EXHAUST
rr= CBC
EXHAUST
VERTICAL
COAL FEEDER
INLET
FBM GAS
BREECHING
STEAM
DRUM
FRONT PANEL
LIGHTOFF
BURNER
OPTIONAL
FEED
POINT
ADDITIVE OR ASH
FEED PORT
AIR PLENUM
-GRID PLATE
-HEADER
-FLUIDIZED BED
Ul
I
DOWNCOMERS
FIGURE 7. FLUIDIZED-BED MODULE (FBM) INTERNAL CONSTRUCTION
(INITIAL CONFIGURATION)
-------
5-11
From the plenum at the base of the unit, air is
directed upward through the grid and then through the
bed. The grid consists of a mild steel plate containing
buttons of the same spacing and design used in the FBC
operation. This bed material used in the FBM tests
was either 1359 limestone or sintered ash, double screened,
-8+20 U.S. Standard mesh. The static bed depth may
be varied from 6" to over 30", although the useful range
is narrower. A bed sampling pipe and valve are provided.
Thermocouples were mounted throughout the bed, as shown
in Figure 8. Detailed specifications of the FBM are
presented in Appendix B. Figures 9 and 10 show section
views through the FBM and the two CBC's.
In operation, the bed is raised to the ignition
point of coal by use of the gas burner. Combustion of
the coal begins in the vicinity of the light-off burner
flame and propagates rapidly throughout the bed. Firing
with a coal input of 800 Ib/hr, the FBM produces 200
psig steam at the rate of 5,000 Ib/hr. The energy not
absorbed by the waterwalls and steam drum leaves the
FBM as hot products of combustion. Two water-cooled
tube arrays to simulate the convection bank and economizer
of a conventional boiler system were installed in the
ducts beyond the FBM to absorb some of this energy.
A schematic drawing of the FBM test system is
shown in Figure 11. Air from an external forced-draft
fan passes through the air preheater and into the FBM
plenum. Coal feed is controlled by the speed of a
screw feeder* which drops the coal into a pneumatic feed
tube at the injection port*. Sorbent materials were
screw fed to a pneumatic injection line at a rate control-
led by a variable speed screw drive. Ash recirculation
* When salt is fed to the FBM, it is screw fed from
a weighing hopper and mixed with the coal feed.
-------
FBM GAS
BREECHING
STEAM
DRUM
DOWNCOMERS
FRONT PANEL
LIGHTOFF
BURNER
FEED PORT
FEED PORT
AIR PLENUM
-GRID PLATE
-HEADER
-FLUIDIZED BED
en
H
K)
FIGURE 8. FLUIDIZED-BED MODULE (FBM) INTERNAL CONSTRUCTION WITH TALL
CBC SHOWING NUMBERED THERMOCOUPLE LOCATIONS
-------
F8M EXHAUST
XHA&T
• >•.?'»""f-~?J ;•."." ••••.?"^'iV-?.-..' .?•"£
COAL FEEDER
(PARTIAL VIEW)
ADDITIVE ASH PORT4*!
(TYPICAL)
z-tmxe
INTERCOMMUNICATION SLOTS
ACCESS DOOR
MU3IIROOM FEEDER (FLVASH)
AM DISTRIBUTION GR*
FIGURE 9, SECTION THROUGH FLUIDIZED-BED MODULE (FBM) AND CARSON-BURNUp
CELL (CBC) (INITIAL CONFIGURATION)
-------
DOWNCOMER
SLOT FOR GRAVITY
FEED TO AIR LIFT
METERING SCREW
TO SAMPLE POINT,
AIR LIFT INLET,
FLUE GAS COOLER
AND ATMOSPHERE
INTERCOMMUNICATION SLOTS
ACCESS DOOR
MUSHROOM FEEDER (FLYASH)
AIR DISTRIBUTION GRID
CBC PLENUM
FIGURE 10. SECTION THROUGH FBM AND CBC (REVISED CONFIGURATION)
-------
STUB STACK
INDUCED DRAFT FAN
ROOF
DRAFT BALANCE
DAMPER
DUST COLLECTOR
DUST COLLECTOR
HOPPER
INCLINED SCREW
PRESSURE
ADDITIVE
HOPPER
STEAM DRUM
DOOR
SIGHT PORT
DUCT FROM
FORCED DRAFT FAN —»
STAR FEEDER
TO CBC —
ASH REINJECTION LINE
HOT AIR LINE
INLET AIR FROM PREHEATER
J
GRID PLATE-
FIGURE 11. SCHEMATIC OF FBM TEST SYSTEM SHOWING VARIOUS SUBSYSTEMS.
WATER COOLED
METERING SCREW
AIR LIFT AIR
RAVITY BED FEED
PLENUM
-------
5-16
is accomplished by pneumatic transport of fly ash from
the dust collector through a star feeder. Bed material
to be airlifted for regeneration is withdrawn from the
FBM and passes through a variable speed,water-cooled
screw feeder.
Flue gas from the FBM passes across the first gas
cooler (convection bank) above the steam drum to reduce
temperature before the gas enters the air preheater.
As the flue gas passes through the air preheater-, a
portion (the coarse fraction) of the fly ash drops out
and is collected in the hopper (see Figure 11). The
bulk of the fly ash is removed by a multi-cone collector
downstream of the air heater. During recirculation,
the coarse fraction (ash knocked down by the preheater)
is screw fed into the dust collector hopper. This
unfortunately tends to cause inhomogeneous feed compo-
sition to the CBC. From the collector, the gas flows
through a long duct to an induced draft fan and then to
atmosphere. A damper is provided in the ducting to
control pressure in the combustion chamber. The system
is operable without the induced draft fan, but is not
usually run pressurized. An isokinetic sampler is
provided for the study of stack particulate emissions.
5.3 The Pilot Scale Carbon-Burnup Cell, CBC
The pilot-scale burnup cell, designated the CBC,
was an appendage to the FBM It was to operate parallel
with the FBM, at a higher temperature but with a common
bed. The design of this system was based on the results
of the previous phase of this program (Reference 1).
The initial CBC consisted of an insulated rectangular
box fabricated of ASTM 446, a steel with a relatively
low coefficient of expansion capable of resisting oxi-
dation at 2000°F, but of low ductility.
-------
5-17
Although the internal dimensions of the CBC reactor
zone were constant throughout the test program, several
detail modifications were required before the CBC per-
formed its functions of burning fly ash efficiently and
with low bed loss; the system was developmental.
5.3.1 The Initial CBC
The initial CBC, shown in Figures 7 and 9, was
that used in the previous test program (Reference 1);
the CBC, after modification, is shown in Figures 8 and
10. As shown in Figures 7 and 9, the CBC was added
to the rear* of the FBM. Its air distributor, 10-5/8" x
15-5/8", or 1.13 square feet, was identical in design
and at the same elevation as that of the FBM. In the
initial CBC configuration, the distance from the top
of the air distributor to the roof of the CBC was 56"**.
Starting 48" above the air distributor, the exhaust
was installed as an 8" I.D. horizontal duct. The
duct extended horizontally under the FBM steam drum,
through the building wall and then ran vertically to a
dust collector installed on the laboratory roof. A
gas sample outlet was provided in the duct. Dust collected
here could be wasted, routed to the FBM via the four-
port injector described earlier, or back into the CBC.
For Run No. 171H, a weighing ash silo was provided.
After a few hours of operation, this device became
inoperative.
*The front of a unit is designated as the side in
which the access opening is located. The front of
the FBM is at the left of Figure 9. The front of
the CBC is at the right in Figure 7.
** The dimensions of the unit were set by the design
features of the FBM and its location in the laboratory.
The height was set by the location of the FBM steam
drum. The width by the location of the downcomers
(see Figure 7) and the depth by a structural wall
20" beyond the boiler supports as shown in Figure 9.
Except for the height, the dimensions were appropriate
for the expected firing rate, 500,000-750,000 Btu/hr.
-------
5-18
Fuel for the CBC was fly ash generated by the
FBM, plus its own fly ash and coal, if required, to
maintain the temperature at the 1900°F-2100°F level.
Although the first trial injector was similar to that
of the FBC (a square, horizontal tube jutting into the
unit), improved operation was obtained with a "mushroom"
feeder as shown in Figure 9. The mushroom feeder, as
the name implies, is a solid cone. Its underside might
be described as a hemi-toroid but the underside of a
mushroom with short stem is a suitable description.
It is fixed by four solid connections to the air distri-
butor. An open vertical pipe extends up through the
plenum and air distributor and ends below the stem of
the mushroom. A fuel/air suspension will leave the pipe
and be deflected by the underside of the mushroom into
an isotropic stream with horizontal inertia. While
the mushroom had been developed in 1967 for coal feeding,
it is most suitable for dry materials such as fly ash.
Fluidizing-combustion air is withdrawn after the
air heater from the main FBM air supply*. The design
of the mushroom feeder makes it capable of supplying
a significant fraction of the fluidizing air although
the plenum and air distribution grid were always used.
The FBM and CBC interface at the rear of each unit.
To provide for a common fluidized-bed, the fins between
boiler tubes making up the FBM's rear wall were removed.
Looking toward the rear of the FBM, the initial configura-
tion had five slots, 18" high and 1" wide, a 2" O.D.
boiler tube between each slot. The back of these tubes
(in the CBC) was insulated by a semicircle of insula-
tion held in place by a thin metal shield. For bed
* Provisions also exist for feeding unheated CBC air
in the unusual event a higher than desired CBC tem-
perature should develop.
-------
5-19
intercommunication, holes cut in the baffle provided
the desired opening, which was 2 square inches in this
program.
Two designs for flue gas "particulate knockouts"
were tested, based on information obtained in April 1970
from BCURA Industrial Laboratories under an agreement
between the EPA and England's National Coal Board. A
knockout was required because the limited height of
the initial CBC (48" to horizontal exhaust) resulted
in excessive carryover of bed particles. The knockout is
shown in Figure 9, labeled "baffle screen".
The first design consisted of an irregular tri-
angular array (4 rows) of 1" tubes connected by U-Bends.
Water flowed into the tube at the bottom and front of
the unit and out the tube at the top and front. Thermom-
eters installed in the inlet and outlet permitted the
temperature rise to be determined. A total of 5.2
square feet of heat exchange surface was provided by
the screen. The effect of this surface was described
in the previous report in this series (Reference 1).
The second screen design consisted of a similar
array of uncooled rods fabricated of ASTM 446. This
screen, as well as the first, was installed within the
CBC extending from the 32" level to the 39" level.
The test procedures for both the FBC and FBM/CBC
operations involved igniting the bed and stabilizing
the combustion at the desired bed temperature until
steady-state conditions prevailed. Steady state was
assumed when the bed temperature was constant and the
gas composition detectors indicated constant values.
A bed sampling pipe and valve are provided.
-------
5-20
5.3.2 The Modified CBC
As development proceeded, it became obvious that
the low CBC freeboard was placing intolerable constraints
on system operation; especially carbon burnup efficiency
and bed carryover. Therefore, a revised CBC* was built
and the FBM steam drum and downcomers modified to allow
a vertical CBC flue configuration (Figures 8 and 10).
The modified CBC cross-section is a refractory cement
box within a 446 steel shell. A flue gas cooler is
incorporated in the top of the CBC. Sample gas is
withdrawn below the cooler. FBM bed material may be
airlifted into the CBC at a point above the gas sampler.
The water-cooled tubes then act as a knockout for air
lifted bed material entrained in the CBC flue gas rather
than CBC entrained bed particles. When FBM bed material
was airlifted, it was withdrawn at a point remote from
the CBC and its rate was metered by a variable speed,
screw feeder which was water-cooled**. By use of refractory
insulation, improved heat economy is achieved and CBC
coal feed should not be necessary to maintain the desired
temperatures.
5.4 The Pilot Scale Bed Regenerator, REG
Following testing of the two-cell SO_ Acceptor
Process (in which the CBC is employed as the sorbent
regenerator), it appeared desirable to physically
separate the high temperature functions, Carbon-Burnup
Cell (high O~ with fly ash feed) and sorbent regeneration
(low 02 with coal feed). The ORM Project Monitor accord-
ingly modified the contract work statement.
*Also of 1.1 ft grid area.
** It is important to note that a fluidized-bed boiler
system using airlift bed transfer contains the basis
for an oversize rejection classifier, permitting
large lump coal utilization.
-------
5-21
The pilot-scale regenerator, designated the REG,
was an appendage to the side of the FBM. It was to
operate parallel with the FBM and CBC, at a higher tem-
perature than the FBM, but with a common bed. The design
of this system was based on the results of the previous
testing with the CBC regeneration operation. The REG
cross-section consists of a refractory cement box with
a carbon steel shell. It was felt unnecessary to use
446 alloy steel. The REG grid area is 10 x 10 inches,
2
0.7 ft , designed to produce a flue gas with a much
higher SO- content than the CBC at comparable super-
ficial gas velocity. The grid pattern is identical to
the FBM and CBC and at the same elevation. Two bed
transfer slots are provided. (See Figure 11A) . Bed material
removed from the FBM via the gravity bed feed (See Figure
11B) and blown into the REG can return via these bed
transfer slots to either the FBM or CBC. Figures 11A, 11B
and 11C each show the relationship -between bhe .three vessels,
When the air lift is not used bed material moved
between the three vessels via gravity.
A bed sample pipe and valve are provided to remove
REG bed material for analysis.
The REG coal feeder is a square, horizontal %-inch x
1-inch tube similar to the FBC operation blowing across the
FBM/REG transfer slot. This aspiration is believed to aid
bed material transfer, but may have also speeded slag
buildup in the REG. Further design, optimization and
testing are needed.
A gas sample exit, airlift entrance and tubular
flue gas cooler are provided similar to the second CBC
flue layout. In order to conserve program funds, no
REG dust collector was purchased, but its dust collected
in the FBM collector. (See Figure 11B)
-------
CARBON STEEL
SHELL
3 CASTOLITE
^ HIGH X l" WIDE
COAL INJECTOR TUBE
I" WIDE X 8" HIGH
BED MATERIAL
FLOW CONNECTION'
TRANSFER SLOT
PLENUM
FLYASH FEEDER
(FLOWS UP FROM BELOW)
PLAN
SECTION TAKEN JUST ABOVE AIR DISTRIBUTOR
(SEE FIGURES 8 AND 10 FOR CBC a FBM DETAILS.)
KEY
SFER SLOTS
IGH
-At
f— ' V
FBM
\ <
\
f
J
1
^ 1
4-1
' * ,
,i 'i '
Z *
= UJ
O _1
Q. 0
2 0
0
7- V)
1— rf
• K • ^*
< t °
(E _j
(_ ^ U-
X ? 0
UJ
l\ A . . 1. ~
v-
v' •
» *
:r;-
^
i lUJ.'.v
ft
((-'*-
\^
i
-^
V
REG
•— io"-»
.*
"0
|
'•>
''*•
."»
» '
'"•.'
•',•
•;•
;•-
[•;
^"^
M x
jl| N
^-3"CASTOLITE J j P
/ 0- P
/ Z gc O
^* 3? ^^
C/J ^
Ul
o P 2 M
f- r- r- KJ
e C *
•— levi — IM — |csi
-. IO 00
i 1
~
-------
CBC FLUE,GAS TO CBC
DUST COLLECTOR AND STACK
TO FBM
DUST
COLLECTOR
AND
STACK
AIR
HEATER
CBC (BEHIND FBM)
GAS TO SAMPLE SYSTEM
FBM
o o
GREATLY EXAGERATED
SEPARATION
•CONNECTING SLOT
DROP LEG
GRAVITY BED
FEED
AIR
GAS OUTLET
to
U)
BED SAMPLE LINE
FIGURE 11B. FLOW SCHEMATIC FOR FBM AND REG.
-------
5-24
UNCONTROLLED
FLOW PATHS
FBM
CONTROLLED
FLOW PATH
EY
BED MATERIAL MAY FLOW IN
BOTH DIRECTIONS- VIA GRAVITY,
PRESSURE DIFFERENCE, OR DIFUSSION.
RATES UNCONTROLLED
BED MATERIAL CAN FLOW
ONLY IN DIRECTION OF ARROW
WITH RATE CONTROL
FIG. II-C. FLOW PATHS OF BED MATERIAL
FOR FBM, CBC, 8 REG.
-------
5-25
The test procedure involved igniting the FBM and
CBC beds in the usual manner and then pulling fluidized
bed from the REG (through the sample outlet tube), until
coal combustion became stabilized in the REG. The desired
velocity, and gas compositions were then created.
In the demonstration runs when the REG was operated,
it was fed fine coal (-8 mesh). We recommend that tests
to demonstrate an optimal coal size for the REG be
performed.
5.5 Instrumentation
Emissions of sulfur dioxide, nitric oxide, carbon
dioxide and hydrocarbons were monitored continuously*.
Infrared analyzers (Beckman 215) were used to monitor
sulfur dioxide, carbon monoxide, and nitric oxide.
Carbon dioxide was measured, using a Beckman 7C thermal
conductivity analyzer. Hydrocarbons were detected with
a flame ionization analyzer (Beckman 109A), using methane
as the calibration gas. The signal output of each of
these units was displayed on strip chart recorders.
The gas transfer system used with these analyzers
is sketched in Figure 12. The system permitted rechecking
of calibrations and zero settings on any of the four units
at any time during the test by switching from sample gas
to reference and zero gases at the rotameter valves.
The sample gas was drawn from the hot flue gas stream
through a sintered stainless steel filter and conditioned
to remove water to a 32°F dew point. The sample gas
was again filtered before entry into the analyzers to
prevent possible contamination of the optical cells and
the hydrogen burner. Wet assay was also used for NO
(PDS procedure) and SO content of process gas samples.
Carbon monoxide was also monitored in the extended
FBM runs.
-------
SYMBOLS:
CALIBRATION REFERENCE AND COMBUSTION GAS SOURCES
o
® CONTROL VALVE FLUE GAS.
NOTE: CARBON MONOXIDE ANALYZER NOT SHOWN- THE INSTALLATION IS IDENTICAL
TO S02 AND NO EXCEPT FOR SPAN GAS
r FILTER 0.3 Urn
FILTER
0.3-Um
/DUCT
FILTER 9Um
/
1
H^
DRYE3.
/
—~ i
-FILTE
v
\
CONDENSER
RELIEF
PUMP
Cn
H2 i
N2 *
CH4
MIX
Cf
N2
ZERO
FIGURE 12. SCHEMATIC OP GAS TRANSFER SYSTEM FOR CONTINUOUS MONITORING
OF SULFUR DIOXIDE, NITRIC OXIDE, CARBON DIOXIDE, AND HYDROCARBONS.
-------
5-27
Methods to assay chloride content of flue gases are
presently under development. Chloride content with
lime beds is expected to be very small*.
The hot FBC gas sample was drawn into the instru-
ment room from the horizontal FBC exhaust duct which
extended overhead.
In sampling the FBM flue gas, special precautions
were necessary because of the possibility of infiltration
of dilution air in the duct above the unit. Also, the
poor instrument response which would result from drawing
a small sample a long distance (^60 feet) from unit
to instrument room was undesirable.
A system was devised to draw a large hot gas sample
from the FBM, just above the first gas cooler, pass it
through a dust collector, and then through a loop above
the instrument room. The sample tube was a 3" pipe
with sections screw-fitted and welded. The system was
driven with an I.D. fan located at the discharge to
atmosphere. A schematic drawing of the system is shown
in Figure 13. The CBC and REG gas conduit systems were
identical in design. The CBC sample line in the initial
CBC configuration was located at the 62" level. For
modified CBC, see Figure 10. The REG gas sample line
was located .below the airlift entrance.
Particulate emissions were monitored with the
isokinetic probe and filter system described in Refer-
ence 2. The probe design permits equalization of
Cl- and HC1 in a filtered flue gas stream would be
absorbed in aqueous NaOH. The chloride content would
be determined colorimetrically with mercuric cnloranilate.
The aosoroance at 530 nm would be correlated witn
chloride content of calibration samples. With a
15 (ju.fu. gas sample, 2 ppm uhloride sensitivity
should be achievable.
-------
5-28
of internal and external static pressures to match the
sampling velocity with the stream velocity. Locations
of sampling points in the FBC, FBM, and CBC test systems
were indicated in Figures 6, 12, and 13, respectively.
Carbon dioxide was continuously monitored, using
a self-referencing thermal conductivity analyzer (Beckman
Model 1C, range 0-25%). The span calibration gas contains
16% CO_. The reference and downscale calibration gas
is purified air. Instrument cabinet temperature is
controlled at 130°F*. Bailey oxygen analyzers (Type
OC1530A) were used as operating devices to indicate
oxygen concentration in the FBC, FBM, and CBC and REG
flue gases. During a test period, the air input rate
was held constant and the coal rate adjusted to maintain
the oxygen concentration at the desired value. The
Bailey instruments have been calibrated periodically
with 0_ , N_ , and CO- mixtures and found to be very
reliable. The flue gas oxygen was also verified, using
the standard Orsat technique, which determined also
carbon dioxide and carbon monoxide**. When the FBM/
CBC was operated, a separate oxygen analyzer served
each system. When the FBM/CBC/REG was operated, the
REG and CBC shared an 0_ analyzer intermittently. Other
instruments were shared by the three units by switching
from one sample loop to the other***.
* We found this instrument to be more sensitive to
SO than to CO_, and also to have a negative
response to CO.
** Limit of detection 0.1%.
*** This procedure was only partly satisfactory when
using the CBC as a regenerator. An S02 instrument
transient of 30 min was observed when 5000 ppm
SO- gas (CBC) was replaced with 200 ppm (FBM)
SO_ gas. NO, CO, CO2 and HC responses were
immediate, however.
-------
Flue Gas
Sample Line
Filters
3" Pipe, Welded
Connections
OOOOOOOO
O O O O O O O
O O O O O O
Loop over
Instrument
Room
Thermocouple
Insulation
Convection
Bank
Steam Drum
Analysis
(Wet Tests)
To IP.
HC
Analyzers
Dust Collector;
Welded Seams
—°a From CBC Exhaust
ui
i
ro
FBM
To
Atmosphere
Flue Gas from
Fluidized Bed
Fan
FIGURE 13. SCHEMATIC OF THE FBM GAS SAMPLING SYSTEM
-------
5-30
FBM Steam Drum
CBC Exhaust
62" Level-
Gas Sample Point
(2" pipe)
Building Wall
Access
rid Level -0"
Exhaust to Dust Collector
and I. D. Fan
\
2" pipe
•Loop over Instrument Room
Cyclone Dust Separator
To Sample Fan
•CBC
B. Schematic of Sample Flow
-1/2"
Stainless
Steel Tube
Gas to Instruments
FIGURE 14. SCHEMATIC OF CBC GAS SAMPLING SYSTEM
-------
5-31
Temperatures in the bed and at various other points
in the system were recorded remotely on a Honeywell
multipoint recorder (cycle time 2 min). A multiple switch
panel was used to connect the 24-point recorder input
to either the FBC or FBM/CBC systems, as required.
Locations of thermocouples in the systems are indicated
in Appendices A and C. West dial meters with multipoint
temperature selector switches were used as operating
indicators, in parallel with the chart recorder. The
West meter readings are usually 25°F above the chart
record.
The infrared analyzers and the hydrocarbon analyzer
were periodically calibrated with gas mixtures supplied
by vendors. The concentration of the active components
in the calibration gases was checked after delivery to
the laboratory. The methane mixture was analyzed by
the National Bureau of Standards -- a report is shown
in Reference 2. This gas, containing 1265 ppm CH.,
was used to calibrate a second methane mixture before
it was depleted.
The sulfur dioxide calibration gas was analyzed
with a peroxide absorption train. Concentrations of
2650 and 2530 ppm were used in the program. The NO
gas was analyzed by PDS technique. Analysis of the
nitric oxide calibration gas is also given in Reference 2.
The output signal of the NDIR sulfur dioxide analyzer
(0-5000 ppm) varies in a nonlinear manner with SO_ con-
centration. The calibration curve provided with the
instrument was checked by precision dilution of the
known calibration gas. The curve was found to be correct
except for a slight deviation at the low end of the
range. The calibration curve and check points are
-------
5-32
are given in Reference 2. The combined contribution of
H_O and CO_ is about 100 to 220 ppm with this analyzer
depending on the dew point achieved. The calibration curve
was used without correction since the deviation is not
more than 1% of full scale. When SO_ levels exceed
5000 ppm, N_ dilution gas can be added through a separate
flowmeter. When flows of N_ and sample are equal, a
scale deflection of "5000 ppm" can then be interpreted
as "1% SO2". In test 171 H and later tests, the SO2
analyzer was modified to have three ranges: 0-920 ppm,
0-5000ppm and 0-5%, S02.
The calibration curve for the nitric oxide NDIR
analyzer is given in Reference 2. The contribution of
water vapor to the signal output is significant with
this analyzer. The water vapor correction determined
by the vendor (180 ppm) was checked by testing a dry
gas in the analyzer for comparison with refrigerated
ambient air. A correction of 100 ppm was noted and
incorporated in the data reduction. The range of this
unit is 0-1000 ppm NO. This analyzer was inoperative
durinq Tests C-321, 2, 3, and 171-H.
The carbon monoxide NDIR analyzer is a triple
range instrument, zero to 0.35, 1.0 and 2.7%. C0_
and H_O give negligible interferences.
Carbon, hydrogen, and sulfur analyses of solid
materials were performed using equipment listed in
Appendix D. Calcium was analyzed by digestion followed
by EDTA titration. Some sulfur analyses (bed material*)
were performed gravimetrically by barium chloride pre-
cipitation (Eschka) .
The Leco procedure is time consuming with many of
the bed material samples due to porcelain formation.
-------
5-33
5.6 Materials
Coal used in FBC tests was Powhatan 7.1% C, 4.5% S,
6.1% ash, 13,123 Btu/lb higher heating value.
Three coals were used in FBM testing; these
are the Powhatan containing 4.5% S and 71% C; another
also Powhatan, containing 3.3% S and 73.5% C; the third
Rivesville coal containing 12.4% ash, 71% C and 3.85% S.
FBC and FBM coals were crushed to V top size.
Limestone used was commercially available No. 1359
double screened -8+20 mesh material normally sold as
poultry grit.
Salt used was mostly table salt. Its size consist
was: on 25 U.S. mesh, 0.1%; on 30, 1.0%; on 50, 85.2%;
on 70, 11.9%; on 80, 0.9%; on 120, 0.2%. Some FBC
testing was done with road -salt having k" top size.
-------
6-1
6. RESULTS OF BENCH AND PILOT SCALE TESTS
6.1 Reduction of Emissions of SO,,
The major purpose of the bench scale experimental
program has been to evaluate various modes of operation
of a fluidized-bed combustor, incorporating limestone
materials to control SO- emission. FBC operation is a
batch process in which a charge of limestone is calcined,
reacted with SO2, and finally as an option, can be
discarded or can be regenerated by increasing coal rate
so that hotter, less oxidizing conditions are created.
A regenerated charge of lime could then be run through
additional sorption-regeneration cycles, if desired.
Such batch cyclic operation resembles, but is not con-
gruent with, tandem FBM/CBC sorption-regeneration
operation.
During the period covered by this report, sintered
ash bed operation with fine limestone addition has been
de-emphasized as a S0_ control means, due to the poor
limestone utilization experience at several installations
using this technique.
The FBC and FBM units contain provisions for fly
ash reinjection; no reinjection studies were conducted
during the period covered by this report. Some minor
improvement in limestone utilization might have been
achieved via reinjection since the fly ash contained
some unreacted lime.
6.2 FBC Tests
Of the extensive FBC test series originally planned,
only four were conducted. Available funds were then
transferred to the FBM long duration testing program.
-------
6-2
FBC testing was conducted with -8+20 mesh, 1359
limestone beds and V x 0 Powhatan 4.5% sulfur coal,
with and without activity enhancement additives. (See
Appendix E for background information.) Run No. C-321
was initiated with low coal and air rates and a low bed
level to expedite calcining. After establishment of
coal combustion and calcining, further bed material
was added. Following completion of most of the calcining,
as evidenced by the flue gas CO- level strip chart
record, data acquisition began. (See Table 1.) Bed
temperature was 1580 F. Since the FBC hood insulating
liner used in previous fly ash combustion tests was
still in place, heat transfer was limited and high
excess air was needed to limit bed temperature. Beginning
at 200 ppm flue gas SO2 built up gradually to 2100
ppm* after 5.5 hours of combustion (including calcining
time). Analyses of bed material samples show a corres-
ponding buildup of bed sulfur content (See Figure 15)**.
At the relatively low ingredients rates and bed level,
apparent combustion efficiency is 90% or more, although
fly ash collection efficiency appears low, especially
in Conditions 1-4. A sulfur balance (see Appendix F)
was compiled for this run showing the amounts accumulated
in fly ash, flue gas, and bed material. The principal
sulfur inventory at short run times (<7 hours) in the
* If no acceptor were present, the coal and air
rates correspond to a maximum of 2750 ppm SO_.
** Bed calcining period was about 2 hours. More precise
definition is not possible since a C02 emission tail-
off of about 30 min. usually follows stone feed
cutoff.
-------
6-3
TABLE 1. FBC TEST DATA SUMMARY
Test No. C-321
1.
2.
3.
4.
5.
6.
7.
8.
9.
10.
11.
lla.
12.
13.
14.
15.
16.
17.
18.
19.
20.
21.
22.
22a.
22b.
Flue
23.
24.
25.
26.
27.
Time (hours) _
Air rate, lb/hr/ft
Bed temperature, °F
Bed depth, in.
(1359 Limestone)
Bed particle size
-8+20
Coal input rate
Ib/hr/ft2
Carbon input, Ib/hr
Fly ash output, Ib/hr
Output C content,
%wt.
Output S content, %
Average 2 . 26
Output Ca content, %
Ratio. Ca/S in output
Carbon output, Ib/hr
Carbon burned
Combustion eff.,%
Superficial velocity.
ft/sec
Cooling probe position,
4 probes 100% inserted
Fuel heat input,
KBtu/hr
Sulfur input, Ib/hr
NaCl rate, Ib/hr
Sulfur output, Ib/hr
CaSO./CaO ratio in
fly ash, wt.
Bed sulfur content,
%wt.
Bed Ca content,
%wt.
Limestone feed rate
Gas Composition:
CO,, %vol.
of %
CO, %
SO , ppm
HC , ppm
-1
2
710
1580
16
51
31.2
N/A
27.6
2.71
20.5
7
N/A
N/A
N/A
9.7
-2
2%
710
1610
16
51
31.2
2.8
47.8
2.52
12.0
4.7
1.33
29.9
96
9.8
throughout the
575
1.98
0
.076
0.48
2.58
51.7
0
10.6
6.8
0
200
3
575
1.98
0
.070
0.86
N/A
N/A
0
10.9
6.4
0
450
15
-3
3.8
710
1580
16
48
29.1
2.2
48.8
2.53
6.1
2.4
1.07
28.
96.3
9.7
test
535
1.85
0
.056
2.6
5.86
51.8
0
11.1
6.8
0
900
0
-4
N/A
710
1560
16
48
29.1
2.5
46.
1.88
5.1
2.7
1.15
28.
96
9.6
535
1.85
0
.047
2.0
N/A
N/A
0
10.7
7.
0
1500
0
-5
5.5
710
1570
16
48
29.1
4.2
53.9
2.08
5.0
2.4
2.26
28.6
92
9.6
535
1.85
0
.087
6.
7.57
48.7
0
11.1
6.7
0
2100
0
-6
6.6
710
1560
16
48
29.1
5.1
56.7
1.95
5.3
2.7
2.9
26.2
90
9.6
535
1.85
4.0
.099
2.0
8.96
49.6
0
11.
7.
0
500
0
-7
7
710
1540
16
48
29.1
7.5
40.
2.15
9.76
4.5
3.
26.1
90
9.5
535
1.85
4.0
.161
0.93
9.75
40.6
0
10.9
•.6.0
0
650
0
-------
6-4
TABLE 1. (continued)
Heat Balance, KBtu/hr
28.
29.
30.
31.
32.
33.
34.
Wall loss
Hood loss
(insulated)
Probe loss
Flue gas loss
(at T bed)
Carbon loss
Ash heat content
Total loss
-1
13.0
28.0
209
263
•v-11
0.5
524
-2
10.0
26.7
209
269
18.7
0.5
. 534
-3
13.4
35.7
201
261
15.1
0.4
527
-4
11.2
30.0
194
257
16.2
0.4
509
-5
12.6
33.6
176
257
32.
0.7
512
-6
10.7
28.5
179
255
41.
0.8
515
-7
10.7
28.5
164
251
42.
1.2
498
Notes: N/A = not available. "Output" material is that collected by the
dust collector.
Bed weight post test not available. 16 in. of used bed usually
weighs 64 Ib.
-------
6-5
IU
9
8
7
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z
0
o
fy ^
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a
UJ
00
2
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y^ONDSl
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'ION -.
10. CD
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NoCI
JECTION
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) 1 234567
CALCINING RUN TIME, HOURS
PERIOD ,
FIGURE 15. BED SULFUR ACCUMULATION VS. TIME
FBC RUN- C-321
-------
6-6
absence of either regeneration or the enhancement addi-
tive is laid down in bed material (sulfated lime).
The variable flue gas SO levels (often 1000-2000 ppm)
during the heavy calcining period account for most of
the deviation between input (fed in coal) and total
outputs, since the percent recovery increases with
time as the flue gas SO2 increases.
Starting at 5.7 hours, coarse salt,k" top size,
addition was initiated, with a screw feeder dropping
salt into the coal feed air stream. Salt was %" road
de-icing type material. The pressure drop across the
feed screw, plus backflow of fly ash, caused intermittent
loss of NaCl flow. Nevertheless, SO_ levels with the
aged lime bed are reduced dramatically by salt addition.
At the bed temperature and air rate used, the calculated
NaCl vapor rate in flue gas is about 1 Ib/hr, based on
published vapor pressure data, and the flue gas isokinetic
sampler displayed a high salt content. Probably the
chloride content of salt is tied up as calcium chloride
in bed material or fly ash, and sodium is tied up as
sodium calcium sulfate in bed material and fly ash.
The apparent effects of the salt addition to an
aged, unregenerated bed containing CaO and 32% CaSO.
are:
a) The SO_ emission rate from coal combustion is
cut back to rate approximately that with freshly calcined
CaO. (See Table 1, Condition 6.)
b) The rate of fly ash collection is increased.
It is hypothesized that the fly ash contains a dust of
CaNa2 (SO.) or similar double salt whose crystal lattice
parameters are sufficiently different from calcined
limestone that particle outer shell strength is reduced
-------
6-7
and surface removal occurs naturally by attrition in
the fluidized bed, leaving exposed active calcium oxide.
c) The rate of sulfur laydown in the bed material
is restored to a value approximating that with freshly
calcined CaO (See Figure 15); i.e.,CaO activity has been
enhanced.
d) The rate of darkening of fresh bed material
(the brown discoloration that may be due to iron pickup
as CaFe-0.) is arrested.
Several names have been proposed for this process:
1) SO- Acceptor Process with Chemical Scrubbing
2) SO Acceptor Process with Chemical Attrition
Following completion of Run C-321, the FBC hood
liner insulation was removed and a vibrator added to
the salt feeder to improve feeder reliability. Run
C-322 was then initiated with a fresh bed to test lime-
stone bed operation at higher coal rate, lower flue gas
O-, and deeper bed conditions. Run temperature was 1600 F.
Flue gas SO- built up at about the same rate as previously.
(See Table 2) At the higher coal and air rates,
fly ash amounts collected were considerably greater and
combustion efficiencies somewhat lower. The significantly
higher hydrocarbon levels in this test (120 ppm) are attri-
buted to both lower O_ and reduced freeboard height, due
to the deeper bed in a constant height apparatus.
Coarse salt addition was attempted following growth
of SO- to 1000 ppm*, a lower level than in the previous
test. Due to the higher low-bed pressure in this test,
fly ash blowback occurred rapidly and only intermittent
If no acceptor were present, the coal and air rates
correspond to a maximum of 3800 ppm SO-.
-------
6-8
TABLE 2. FBC TEST DATA SUMMARY
Test No. C-322
1.
2.
3.
4.
5.
6.
7.
8.
9.
10.
11.
lla.
12.
13.
14.
15.
16.
17.
18.
19.
20.
20a.
20b.
20c.
Flue
21.
22.
23.
24.
25.
Time (hours) _
Air rate, Ib/hr/ft^
Bed temperature, °F
Bed depth, in.
(Limestone)
Bed particle size,
-8+20
Coal input, Ib/hr/ft
Carbon input, Ib/hr
Fly ash output, Ib/hr
Output C content, %
Output S content, %
Output Ca content, %
Ratio Ca/S in fly ash
Carbon output, Ib/hr
Carbon burned
Combustion eff., of
carbon, %
Superficial velocity,
ft/sec
Cooling probe position
Fuel heat input,
KBtu/hr
Sulfur input, Ib/hr
Sulfur output
Bed sulfur content,
wt.%
Bed calcium content
Ratio Ca/S in bed
Salt feed rate
Gas Composition:
CO , %vol.
of %
CO, %
SO-, ppm
HC,ppm
-1
1.9
800
1590
22
N/A
74
45
20,7
27.6
2.26
17.0
7.5
5.7
39.3
87.3
10.8
3 in
830
2.86
0.47
2.36
49.1
21
0
14.8
3
0
200
150
-2
2.9
780
1630
N/A
N/A
74
45
20.7
22.4
2.52
18.5
7.4
4.6
40.4
90
10.6
3 in
830
2. 86
0.52
4.39
51.3
11.7
0
14.8
3.1
0
1000
120
-3
3.9
760
1595
N/A
N/A
65
39.8
25.5
*29.4
2.42
15.7
6.5
*7.5
32.3
N/A
10.4
3 in
731
2.52
0.61
6.26
46.5
7.4
0
15.
3.2
0
1800
180
-4
4.9
770
1620
N/A
N/A
81
49.6
22.8
26.6
2.57
16.3
6.4
6.1
43.5
88
10.5
4 in
915
3.15
0.59
7.10
47.4
6.7
0
15.
2.9
0
**1800
180
-5
5.9
770
1610
N/A
N/A
77
46.9
31.7
32.4
2.18
19.2
8. 8
10.2
36.7
78.3
10.5
4 in
862
2.97
0.69
8.70
46..?
5.3
0
14.7
3.
0
2750
170
-6 Avg
6.9
780
1620
N/A
N/A
76
46.5
24.9
29.7
2.23 2.3
19.7
8.8
7.4
39.1
84.1 85.
10.6
4 in
855
2.95
0.56
9.71
43.8
4.5
0
14.6
3.2
0
2900
120
-------
6-9
TABLE 2. (Continued)
Heat
26.
27.
28.
29.
30.
31.
32.
33.
Balance, KBtu/hr:
Wall loss
Hood loss
(uninsulated)
Probe loss
Flue gas loss
(at T bed)
Carbon loss
Hydrocarbon loss
Ash heat content
Total loss
-1
2
181
183
306.2
80.4
2.6
6.4
762
-2
3
191
208
308.7
65
2
6.4
784
-3
2
153
169
285.2
106
2.9
7.7
725
-4
3
176
187
310.2
86
3
7
772
-5
3
166
197
298.7
144
2.8
9.7
821
-6
3
166
197
307.6
104.3
2
7.6
788
*
**
FBC banked 10 min.
7 . 6 Ibs makuo bed <
between -2 and
added between -
-3.
3 and
-4 (Raw
basis)
Initial bed mass 187.9 Ibs (raw); 105 Ibs (calcined)
Final bed mass 88 Ibs after test (as result of elutriation, sampling)
Note: No salt addition (see text).
by the dust collector.
"Output" material is that collected
-------
6-10
feed of small amounts of salt could be achieved, with
only small perturbing effects on SO2. The FBC unit was
then banked, the salt hopper cleaned out, and a fresh
batch of fine salt introduced.. The unit was restarted
and fine salt addition attempted. Although no fly ash
blowback occurred, the 1/80 HP variable speed motor
was unable to operate with this much frictional drag.
and either stalled or stripped its set screws. Salt
feed was therefore abandoned, and the balance of tb^
test conducted as a straight limestone bed te^t.
Since there was evidence of some bed material being
carried over and collected with the fly ash, at the high
combustion rate and bed level conditions, a small amounu
of fresh limestone bed material (7.6 Ib) was ted very
slowly over a 40 min. period between Conditions 3 and 4.
The effect on'flue gas CO-, due to calcining at this low
makeup rate was minimal. In a full scale boiler using
this once-through type of process, withdrawals of sulfur
would be made in the forms of partial!'/ reacted Ibcd
material as well as fly ash containing CabO , ^ad mai;oup
limestone would be required.
The bed sulfur content and sulfur balance data tor
this run are shown in Figure 16 and Appt/.n.J • JL p 7 . m';c
length of the calcining period is interpreted from (ho
CO- level strip chart record.
Run No. C-323 was initiated with low coal and a.i r
rates, and a low bed level to expedite ca3.c''riirvi . A?tr r
establishment of coal combustion and calcining, further
bed material was added. Following completion of mo^t of
the calcining, as evidenced by the flue gas CO- level
strip chart record, data acquisition began (See Table 3).
Bed depth was 14 in. (static) . Due to the limited ho«*t
-------
6-11
IU
9
8
7
38 6
H 5
z
UJ
z
8 4
(E
Z)
u.
3 3
a
UJ
CO
2
1
n
/
/
/ c
/ PI
f
/
'
COINING
ERIOD
?
/
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LIMES'
" ADDITI
PERIOI
/
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ON
[}
X
s
1234
RUN TIME, HOURS
FIGURE 16- BED SULFUR CONTENT VS. TIME
FBC RUN C-322 (LIMESTONE BED)
-------
6-12
TABLE 3. FBC TEST DATA SUMMARY
Test No. C-323
1.
2.
3.
4.
5.
6.
7.
8.
9.
10.
11.
12.
13.
14.
15.
16.
17.
18.
19.
20.
21.
22.
23.
23a.
Flue
24.
25.
26.
27.
28.
Air rate, Ib/hr/ft
Bed temperature ,°F
Bed depth, in.
Bed particle size,
-8+20 2
Coal input, Ib/hr/ft
Input carbon content,
%wt. ' .
Carbon input, Ib/hr
Fly ash output, Ib/hr
Output C content,
%wt. ":
Output S content,
%wt.
Output Ca content,
%wt.
Ratio Ca/S
Carbon output, Ib/hr
Carbon burned
Combustion eff.,%
Superficial velocity,
ft/sec
Cooling probes
position
Fuel heat input,
KBtu/hr
Sulfur input, Ib/hr
NaCl rate, gal/hr
Sulfur .output, Ib/hr
Bed sulfur content,
%wt.
Bed Ca content,
%wt.
CaCO, Feed Rate
Gas Composition:
CO , %vol.
0-7 %
CO, %
SO , ppm
HCf ppm
-1
790
1650
14
58
71
35.5
2.9
43.2
2.46
13.2
5.3
1.2
34.3
-
11.3
4 in
662
2.25
0
.071
3.98
53.6 •'
0
13
4.8
0
600
10
-2
800
1600
14
58
35.5
6.
56.0
2.08
7.24
3.5
3.3
32.2
90.7
11.0
4 in
662
2.25
0
.125
5.70
49.4
0
11.8
6.
0
950
30
-3
790
1610
14
55
33.4
4.6
46.2
2.05
15.13
7.3
2.1
31.3
94.
11.1-
4 in
662
2.11
0
.094
6.85
49.2
0
12.
5.8
0
1850
0
-4
800
1530
14
79
48.3
7.7
51.8
2.21
7.62
3.4
4.0
44.3
92.
10.7
3 in
900
3.06
0
.17
8.61
41.6
0
12.5
4.8
0
660
0
-5 Avg
800
1525
14
79
48.3
7.7
47.
2.43 2.25
10.85
4.4
3.6
44.7
92.5
10.7
3 in
900
3.06
1.4**
.18
8.26*
37.3
0
12.3
5.
0
660
0
* Aged sample, hydrated
**26 wt% aqueous solution
-------
6-13
TABLE 3 (Continued)
Heat
29.
30.
31.
32.
33.
34.
35.
36.
Balance, KBtu/hr:
Wall loss
Hood loss
Probe loss
Flue gas loss at
14000F
Carbon loss
Ash heat content
H_O AH
TOtal loss
-1
8.3
111
214
256.4
17
0.9
0
608
-2
5.9
100
201
259.1
47
1.8
0
616
-3
6.5
110
191
254.3
30
1.4
0
594
-4
3.2
119
143
272.1
57
2.3
21
618
-5
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
-------
6-14
transfer at this low bed level, temperature could not
be maintained below 1610°F (above the optimum value for
SO- absorption), at O_ levels of about 5.8%. Beginning
at 580 ppm, flue gas SO_ built up gradually to 1850
ppm* after 3.7 hours of combustion (including calcining
time). Analyses of bed material samples show a corres-
ponding buildup of bed sulfur content (See Figure 17).
At the moderate ingredients rates and low bed level,
apparent combustion efficiency is 91% or more, although
fly ash collection efficiency appears low, especially
/
in conditions 1 and 3. A sulfur balance (See Appendix
Figure F-3) was compiled for this run, showing amounts
accumulated in fly ash, flue gas, and bed material. As
in previous non-regeneration tests;, the principal sulfur
inventory at short run times (<6 hours), in the absence
of the enhancement additive, is laid down in bed material
(sulfated lime).
Starting at 4 hours, salt solution addition was
initiated. Pressurized saturated salt solution (35
grams NaCl in 100 ml HO) was fed from a pressurized
vessel through a valve and flowmeter and dribbled into
the coal-feed air stream, at the coal introduction point.
This method is less malfunction-prone, and smoother
than the screw feeding of coarse or fine salt, but imposes
a heat loss on the system in the form of water vaporization.
Visual inspection of the coal-feed air mixed stream
during operation disclosed no buildup of evaporated salt
deposits in the coal-feed system. As previously, the
SO_ levels in the flue gas with the aged lime bed are
reduced dramatically by salt addition. The additional
If no acceptor were present, the coal and air rates
correspond to a maximum of 3680 ppm SO_.
-------
6-15
8
• 6
UJ
»-
2
O
O
Id
CD
>t-SALT ADDITION
0 UGED SAMPLE)
01234
TIME RUN, HOURS
FIGURE 17. BED SULFUR CONTENT
VS. TIME FBC RUN C-323
-------
6-16
effects of salt noted previously are also prominent --
increased fly ash production and increased rate of sulfur
laydown in bed material. As already stated on p. 6-5,
vaporized salt condenses in the cooled flue gas as an
aerosol.
The test was abruptly terminated prematurely by
a city-wide power failure. The condition 5 bed sample
was inadvertently aged about 1 month before analysis,
leading to hydration and carbbnation, and hence low indicated
S and Ca contents, but the ratio Ca/S follows the orderly
progression of the other conditions. Compared to Test
No. C-322, the hood heat loss is lower in this test due
to the lower bed level and resulting in increased freeboard
height with less bed particle impingement on the cooled
surfaces — simultaneously, there is a lower carbon loss
and higher combustion efficiency, with reduced flue gas
hydrocarbon levels.
Following completion of Run C-323, the FBC solid
salt feeder" was reinstalled in the system with a new
low friction feed screw. Run C-324 was then initiated
with a fresh bed to test limestone bed operation at lower
air rates, lower flue gas 0_, and deeper bed conditions.
Run temperature was 1540°F. Beginning at 150 ppm, flue
gas SO_ built up gradually to 850 ppm* after 4.4 hours
of running (See Table 4). Combustion efficiencies averaged
about 89.5%. Fine salt addition (screw fed into the coal
feed air stream) was initiated following the condition
No. 3, at an SO_ level lower than in the previous test.
Even at the low air rate (low feed air manifold pressure)
salt feed was erratic. Several periods of down time
* If no acceptor were present, the coal and air rates
correspond to a maximum of 3100 ppm SO_.
-------
TABLE 4. FBC TEST DATA SUMMARY
Test No. C-324
1.
2.
3.
4.
5.
6.
7.
8.
9.
10.
11.
12.
13.
14.
15.
16.
17.
18.
19.
20.
21.
22.
23.
24.
Air rate, Ib/hr/ft
Bed temper at ure,°F
Bed depth, in.
Bed particle size, -8+20
Coal input, Ib/hr/ft2
Input carbon content, %wt.
Carbon input, Ib/hr
Fly ash output, Ib/hr
Output C content, %wt.
Output S content, %wt.
Output Ca content, %wt.
Ratio Ca/S in output
Carbon output, Ib/hr
Carbon burned
Combustion efficiency, %
Superficial velocity, ft/sec
Cooling probes position
Fuel heat input, KBtu/hr
Sulfur input, Ib/hr
NaCl rate, Ib/hr
Sulfur output, Ib/hr
Bed sulfur con tent , %wt.
Bed Ca content, %wt.
Ratio Ca/S in. bed
-1
720
1590
17
N/A
53
71
32.6
N/A
39.3
2.31
22.9
9.9
3.3
29.3
90
10.4
3 in
609
2.07
0
0.19
3.39
53.2
15.7
-2
730
1560
17
N/A
58
N/A
35.5
8.3
39.8
2.60
16.22
6.2
3.3
32.2
91
10.3
4 in
662
2.25
0
0.22
4.85
56.8
11.7
-3
720
1550
17
N/A
60
N/A
36.7
8.8
45.5
2.54
11.80
4.7
4.0
32.7
89
10.1
N/A
688
2.34
0
0.22
6.45
54.7
8.5
-4
740
1530
17
N/A
58
N/A
35.5
8.
50.6
2.42
15.18
6.2
4.1
31.4
88.5
10.6
N/A
662
2.25
3*
0.19
7.75
51.4
6.6
-5
700
1470
17
N/A
53
N/A
32.6
16.2
35.3
2.89
18.95
6.5
5.7
26.9
83
9.5
N/A
609
2.07
'3
0.47
9.86
46.1
4.7
-6
710
1520
17
N/A
59
N/A
35.8
16.6
44.3
3.89
12.70
3.3
7.3**
28.5
N/A
9.9
N/A
666
2.27
2
0.65
7.76
45.8
5.9
-7
720
1520
17
N/A
59
N/A
35.8
8.5
26.4
3.49
22.50
6.4
2.3
33.5
94
9.0
N/A
666
2.27
2
0.30
11.0
45.4
N/A
- 8 Avg .
710
1540
17
N/A
47
N/A
28.4
20****
33.7
3.35 2.47*****
3_ 4******
22.40
6.6
6.7**** T
21.7 ;-;
77**** 89.3
10.0
N/A
530
1.8
0
O.G7
8.42
41.4
N/A
See Footnotes on page following.
-------
TABLE 4. (Continued)
Test No. C-324
Flue Gas Composition:
25. C0_, Vol. %
26. 0 , %
27. CO, %
28. NO, ppm ***
29. S0_, ppm
30. HC7 ppm
Heat Balance,' KBtu/hr
31. Wall loss
32. Hood loss'
33. Probe loss
34. Flue gas loss at 1400°F
35. Carbon loss
36. Ash heat content
37. Total loss
NOTES:
*
**
***
****
*****
******
-1
-2
-3
-4
-5
-6
-7
Salt feed erratic.
Down time: 5 min. 2:25 to 2:30; 3.5 min. 2:48 to 2:52.
NO recorder erratic.
18 minute condition.
Without salt.
With salt.
Small sorbent addition made intermittently; see text.
-8
13
4.4
0
390
280
5
2
124
170
234
47
2.5
580
12.9
4.
N/A
300
430
40
3..1
115
195
239
47
2.5
602
13.3
3.7
N/A
250
660
90
2
116'
189
237
57
2.6
604
13.2
2.8
N/A
270
430
2
2
127
195
243
58
2.3
628
13.
3.
N/A
180
430
10
2
93
139
226
81
4.5
546
13.5
3.4
N/A
240
460
2
2
108
171
231
103
4.8
620
13.
4.8
N/A
200
490
2
2
108
184
237
33
2.5
567
14.4
2.2
N/A
140
1300
170
N/A
N/A
N/A
227
95
5.8
•^622
Average
-------
6-19
occurred in attempts to feed salt uniformly. With salt
addition, the SO- level was reduced to a fairly uniform
430 ppm. Simultaneously there was a dramatic reduction
in flue gas hydrocarbon content, to essentially zero
ppm, attributable to the catalytic effect of salt on
vapor phase combustion. Fly ash rate is increased by
salt addition, as previously. Note that the chemical
attrition theory implies that 58 grams of salt can yield
as much as 278 grams of calcium sodium sulfate (CaNa_ (SO.) ~) ,
Some salt is lost at 1500°F as vapor. The observed ratio,
8 Ib additional fly ash to 3 Ib salt fed, is consistent.
with the theory. There was a temporary nitric oxide
(flue gas) reduction at the start of salt feeding, but
no lasting effect. In later FBM tests with salt addition,
salt reduced the NO output.
At six points during the test, limestone bed material
increments were fed over time intervals of 1, 2, or 3
minutes, and at the rate of 1 Ib/min, to reduce bed tem-
perature and make up bed level after attrition losses.
In each pulse, typical bed temperature drop was 50°F;
C0_ increases 2.5 to 3 percent, both effects due to cal-
cining. In addition, hydrocarbon readings abruptly increase
by 200 to 300 ppm; O level temporarily drops by about 2.5
percent in conformity with the hydrocarbon rise; NO
levels drop about 100 ppm in accord with the O- change;
SO- temporarily increases 50 ppm in accord with the 0~
shortage, but quickly recovers and lines out at about 30
ppm below the value preceding limestone pulsing. The
source of hydrocarbons during limestone addition may be
the result of an organic impurity in the raw stone, an
effect on the coal feed entering via the same port, or
an effect on bed hydrodynamics.
-------
6-20
Fly ash sulfur contents in this run average 2.5%
without salt, and 3.4% when under the influence of salt.
Bed sulfur content, rises smoothly to 9.86% at 6 hours
of running; beyond this time the sulfur values become
erratic, probably due to aging of samples and hydration
prior to analysis. The apparent ratios Ca/S in condi-
tions 5, 7, and 8 in the fly ash are higher than those
in the bed material; this is an interesting effect, and
was noted in conditions 5 and 6 of Run C-322 where no
salt had been fed; it implies that the bed particle core
is enriched in sulfur and that sulfur (as SO2, SO.,,
SO~ ,SO^ or other form) under coal combustion conditions
is mobile in the particles and that the surface, which
erodes, becomes depleted in sulfur as it is enriched
in iron, and possibly other contaminants. This mobility
has been named D.E.S., or Dynamic Exchange of Sulfur.
Sxperiments required to understand D.E.S. were described
in an unpublisned progress report (11). Electron micro-
probe studies at Argonne National Laboratory (12) appear
to confirm the existence of D.E.S.
Sulfur balance data for this test are shown in
Table F-l. The sulfur balances in conditions 1 through
5 are good; thpse. in conditions 6 through 8 are unsatis-
tory, in line with the erratic bed sulfur content values
mentioned previously. As previous, most sulfur is
retained in the bed.
Following condition 7, salt was cut off and excess
air reduced to drive-up SO_ emission. This was a short
condition, and SO_ rose by 800 ppm to 1300 ppm, at which
point the test was terminated. Hydrocarbons rose to
170 ppm.
-------
£-21
A discussion of the bed particle size vs time data
recorded in these tests is presented in Section 6.6.
Following completion of test No. C-324, the program
financial and manpower resources were devoted solely
to larger scale FBM testing (see next section).
6.3 FBM/CBC Tests
As discussed in Section 4, Introduction, the demon-
stration goals of the FBM testing were:
a) 90% or better S0_ removal from the primary
cell flue gas;
b) 3 to 4% (vol.) or more SO_ in regeneration
section flue gas;
c) 98% or better carbon burnup efficiency in the
overall system.
6.3.1 Test with the Short CBC/Regenerator
The integrated FBM/CBC boiler system was operated
and data recorded in a preliminary SO- Acceptor Process
test. A crushed limestone bed was used for SO_ capture
in the FBM. Simultaneous fly ash and coal feed to the
CBC were used. During the final portions of the test,
the CBC stack gas contained between 1.5 and 2.4 percent
S0?, the concentration depending inversely on 0_ level,
indicating regeneration of circulating bed material at
about 1870°F. The concentration factor, defined as the
ratio of the SO- concentration in the regenerator region
off-gas to that in the sorbent region off-gas, was on
the order of 20. S0_ output from the FBM was above
900 ppm.
-------
6-22
The purpose of Test No. B-18 was to determine
tbe physical limitations 'of the current FBM/CBC system
when operating in the S0_ Acceptor Process MK I (two
cell), mode. In this, mode., one region (the primary cell)
of a f luidized-bed boiler is operated under conditions
favorable to sulfur acceptance by lime (T...,,^ <1550°F,
Excess O_ >2%) while a smaller, adjacent region is
operated under conditions favorable to sulfur rejection
(T >1800°F, Excess 0_ <4%). This high temperature
O£jU £.
zone would also burn the fly ash from the primary zone.
Bed particles are made to circulate between the two regions,
for example, by means of diffusion* or under the influence
of a pressure differential. This test was made with E.
Ohio, Pittsburgh No. 8 seam "Powhatan" coal with 12.1%
ash and 4.5% sulfur contents. A crushed (-8+20) 1359
limestone bed was used. Run duration was about ten
hours. The intent of the test was to capture sulfur
and generate a partially sulfated lime bed in the boiler;
and then, demonstrate simultaneous regeneration of aged
bed material in a hotter, less oxidizing zone (the CBC
acting as a regenerator) . To. expedite circulation of
sulfated boiler lime to the CBC and regenerated lime
back to the boiler, additional intercommunicating slots
were cut in the baffle which separates the two regions.
A total of six square inches 'was provided compared to
the 4 inches used in Ref. 1. Due to the increased inter-
communication area, the heat sink effect of the FBM
*"• • :
made it necessary during the test to drive up the CBC
temperature by feeding coal to supplement the fuel value
of the carbon bearing fly ash. Additional air required
In portions of later tests, diffusional circulation
was supplemented by mechanical circulation.
-------
6-23
to burn the coal brought the CBC air rate to 1190 Ib/hr,
close to 20% of the total for the system. This contrasts
with the 10% design value. A low amount of air to the
regeneration zone is very important in the regenerative
mode of operation, since a concentrated SO- stream is
desired.
The combination of the high volatile coal being
used in the FBM and the less erosive nature of lime bed
particles compared to ash bed material appeared to lead
to somewhat thicker than normal carbon deposits on the
400 F boiler tubes. This reduces the heat absorption
from the bed and raises the bed temperature*. Normally
this is counteracted by raising the bed level until the
desired equilibrium is restored. Attempts were made to
provide adequate boiler waterwall heat transfer by increasing
bed mass. However, due to the inadequate freeboard and
high air velocity in the CBC (short configuration),
much loss of bed material was experienced, with associated
plugging of the CBC gas sample line. The overbed baffle
screen, which was effective in the short CBC with an
ash bed, did not seem as effective in knocking down
regenerated lime particles.
Test data are summarized in Table 5, corresponding
to about the ninth hour of running.
Due to the coal being fed at various rates to the
CBC, and to the large fraction of lime particles in the
cyclone product, the heat balance around the unit is
incomplete and carbon-burnup data are not presented.
* This problem did not recur in subsequent
tests.
-------
6-24
TABLE 5. FBM-CBC TEST SUMMARY
Test No. B-18
FBM Heat Balance
Fuel heat input, KBtu/hr
Boiler steam gain
Circulating H2O absorption
Flue gas loss (includes CBC)
Thermal efficiency,%
5886
3248
1295
911
77.2
B
N/A
3400
N/A
N/A
N/A
FBM
Bed material
Air rate, Ib/hr
Superficial velocity, ft/sec
Coal rate, Ib/hr
Bed temperature, °F
Fly ash output, Ib/hr
Fly ash carbon content, %
Fly ash sulfur content, %
Fly ash calcium content
Carbon combustion eff. ,%
Flue Gas Concentration
O2 , vol. %
SO2 ppm
NO ppm
CBC
Air rate, Ib/hr
Superficial velocity, ft/sec
Coal feed
Bed temperature, °F
Flue Gas Concentration
O2/ vol. %
S02, %
NO ppm
calcined limestone
4774 N/A
7.6 N/A
450 N/A
1610 1630
125 N/A
56.7 N/A
2.57 N/A
7.15 N/A
78. N/A
5.
970
360
1102
14.5
Yes
1850
1.3
1.55
330
4,
1040
380
N/A
N/A
Yes
1880
2.44
260
-------
•6-25
Due to inability to maintain the desired high
bed level, the bed was usually above 1550°F, the temperature
at which better S0_ capture would occur (see Test Log,
Appendix H).
During most of the test, the hydrocarbons analyzer
gave erratic output and results are not reported. The
failure of the hydrocarbon analyzer is not related to
the operation in the SO2 acceptor mode, and did not recur
in subsequent tests.
As shown in Appendix H, flue gas S0_ levels of
up to 1.5 to 2.4 percent were achieved from the CBC/
regenerator. At the CBC air rate used, the 2.4 percent
value exceeds the steady state value (sulfur input to
the combined system); i.e., the total bed inventory was
experiencing a depletion in sulfur and the 2.4 percent
level would not have been maintained indefinitely. Bed
sulfur content data are shown in Figure 18.
Based on the results of this test, a new CBC/
regenerator as discussed in Section 5.3, was designed
having sufficient freeboard to eliminate regenerated
lime particle entrainment, as well as sufficient refractory
insulation, giving improved heat economy to avoid the
necessity of coal feed to maintain the desired temperature.
It appears likely that the bed intercommunication area
2
provided in this test (6 in ) was excessive, inhibiting
achievement of the desired temperature differential.
2
Two slots, each 2 in , are probably sufficient; if at
different levels, with bed materials of different sulfur
content and temperature a thermosiphon effect will occur.
Manometer data indicated expanded bed heights of 16" (FBM)
and 22" (CBC). Since the overbed pressure differential
was 2.9" H-O, there was probably some admixture of high
S0_ CBC gas with the FBM flue gas, partially accounting
for the 970-1040 ppm S0_ readings off the FBM.
-------
BED SULFUR CONTENT, WT. %
3 rooj-feui o->
/
X
X
x
X
/
/
X
yX
X
I INITIATE
CBC
OPERATION
^^
t
CONDI
(TAB
"\
I
TION
LE 5)
K)
CM
3456
RUN TIME, HOURS
8
10
FIGURE 18. BED SULFUR CONTENT VS. TIME, FBM RUN B-18
ILLUSTRATING REGENERATION
-------
6-27
The principal reaction occurring in the regenerator
process is:
C + CaSO4 + JsO2 -* CO2 + CaO + S02
which is endothermic by 750 cal/gram of sulfur. The sulfur
release is about 9 Ib/hr. This endotherm is more than
offset by the carbon-burnup exotherm (7800 cal/gram).
A number of factors can lead to an energy deficit
within the fluidized-bed for operation in the 2000 F
range. Energy can be lost through the walls, in heating
bed particles which enter the system at a lower temperature,
in heating inert dust and the carbon which does not burn
within the bed, and in heating nitrogen and excess oxygen.
An energy balance in a high temperature bed can be achieved
by preheating the reactants, by heat exchange against
products and/or by minimizing energy losses.
6.3.2 FBM/CBC Extended Run (Run 168H): Test with the Tall
CBC/Regenerator*
6.3.2.1 Test Operation and Burnup Results
The purpose of this test with the new CBC (See
Figure 8) was to continuously regenerate lime bed material
and demonstrate sulfur capture in the FBM of 90% or better
based on coal input, and to study CBC carbon burnup and
simultaneous flue gas SO_ output as a function of CBC
conditions (temperature, oxygen level, additive coal
feed, carbon content of input fly ash). The test duration
was 72 hours plus initial time (about 8 hours) establishing
burnup/regeneration conditions. Test data are summarized
in Table 6. Coal was the "Powhatan" 71% C, 4.5%S, 12.1%
ash material. The 90% capture goal was easily met and
exceeded. Fifty-nine capture measurements were logged,
This test was preceded by a shakedown run of 28
hours' duration. Coal feed problems were experienced
and no data are reported.
-------
TABLC o. FDli/CBC SYSTEM LOG - 168H RUN
Day
3 8
9
NR
NR
NR
NR
10
FBM
Time
1808
1900
2015
2110
2200
2300
0007
0137
0200
0255
0400
0430
0610
0720
0747
0840
0949
1051
1152
1312
1404
1524
1644
1745
1845
1945
2100
2130
2300
2400
0100
0200
0300
0400
0500
0635
0730
0821
0930
1029
1128
Bed Coal Rate Air Rate
T, «P
1550
1550
1540
1570
1555
1570
1550
1490
0147
1500
1540
1540
1490
1470
1450
1450
1480
1480
1475
1460
1485
N/A
1500
1505
1495
1550
1500
1540
1590
1540
1520
1520
1500
1480
1510
1530
1530
1510
1525
1520
1525
Level
LOW
13.
N/A
N/A
16.6
N/A
N/A
16.5
16.
N/A
15.2
15.2
16.
N/A
N/A
17.
17.
17.
17.
18.
18.
N/A
17.
17.
17.
16.
16.
16.
17.5
N/A
17.1
N/A
N/A
17.5
16.6
17.3
N/A
N/A
17.5
17.5
18.
Ib/hr
500
500
466
522
522
514
N/A
N/A
540
540
656
557
557
N/A
N/A
N/A
525
500
560
417
N/A
H/A
545
N/A
N/A
N/A
548
548
548
548
600
610
H/A
N/A
N/A
600
592
N/A
639
591
576
Ib/hr
5910
6090
6060
H/A
5995
N/A
N/A
6468
N/A
N/A
N/A
6545
6588
N/A
N/A
6494
6160
6365
6895
6538
N/A
Banked
6882
7152
7142
7346
N/A
7207
7365
N/A
7250
N/A
Banked
7310
7562
7411
N/A
7130
7115
7025
6930
Fly Ash
Carbon, %
N/A
46.5
46.6
N/A
54.5
N/A
55.
N/A
N/A
N/A
N/A
40.3
52.7
53.9
N/A
58.6
55.5
N/A
52.5
33.5
57.4
N/A
55.
43.6
57.5
N/A
N/A
51.9
51.5
N/A
N/A
N/A
52.8
N/A
47.8
56.
41.2
54.6
35.4
57.
59.7
Flue Gas
S02- pjim
350
150
250
230
N/A
220
H/A
N/A
N/A
H/A
150
250
300
N/A
200
150
200
250
310
3BO
200
N/A
300
250
200
200
230
250
280
230
200
N/A
150
170
250
260
250
250
240
260
Flue Gas
Ib/hr
N/A
N/A
6331
H/A
6358
H/A
6329
N/A
H/A
N/A
H/A
K/A
6959
6954
N/A
N/A
6500
N/A
7291
6880
6882
N/A
7237
7536
7488
N/A
H/A
7618
7732
N/A
N/A
N/A
N/A
N/A
7973
7800
^7833
7516
7587
N/A
N/A
S Release
Rate, Ib/hr
2.28
0.9
1.75
1.53
1.61
1.46
1.54
N/A
N/A
K/A
1.09
1.79
2.3
N/A
1.46
1.08
1.43
1.75
2.50
2.885
1.52
H/A
^2.79
2.49
2.06
1.63
1.63
1.93
2.13
2.28
1.85
1.60
N/A
1.22
1.49
2.15
2.24
2.07
2.09
1.86
1.99
4 S
Bed S
Capture fc
89.9
96.
91.6
53.4
93.1
93.6
93. 3
N/A
N/A
H/A
95.4
92.8
90.8
N/A
93.7
95.3
94.0
92.2
90.1
84.6
93.8
N/A
^88.6
89.5
91.6
93.3
93.3
92.1
91.4
90.7
93.1
94.1
H/A
95.4
94.5
92.
~91.6
~92.2
92.7
93.0
92.3
H/A
N/A
7.84
7.84
7.84
7.34
7.84
7.84
7.84
4 . 'S
N/A
2.50
3.68
4.95
N/A
5.60
6.20
6.88
8.15
Salt in
N/A
H/A
H/A
N/A
H/A
N/A
H/A
H/A
2.82
N/A
N/A
N/A
N/A
N/A
1.40
1.41
N/A
1.24
1.26
1.16
1.11
Fly Ash-
4
S
N/A
1.50
1.71
N/A
1.78
N/A
N/A
N/A
N/A
1.89
H/A
1.78
1.70
1.74
H/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
K/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
1.84
N/A
1.88
1.68
1.54
1.65
1.76
1.45
1.45
Ca
H/A
N/A
N/A
N/A
N/A
N/A
H/A
K/A
N/A
N/A
K/A
N/A
N/A
N/A
N/A
N/A
5.30
N/A
6.63
5.42
N/A
N/A
H/A
N/A
N/A
N/A
H/A
8.28
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
H/A
N/A
H/A
°2
t
3.8
4.
4 .7
3.5
2.8
2.8
3.
J . 5
4.8
3.3
3.
3.5
2.5
3.4
3.5
3.2
2.7
2.8
2.7
3.0
3.
N/A
3.4
4. 5
5.7
5.4
6.7
b.4
3.2
3.5
3.
2.6
N/A
4.
4 .
3.7
3.8
3 . u
4.1
4.5
4.2
CO
\
0.3
0.7
H/A
K/A
K/A
N/A
N/A
H/A
H/A
N/A
N/A
N/A
1.8
N/A
N/A
1.2
1.5
1.1
1.2
0.9
2.8
N/A
0.85
0.8
0.44
0.31
0.3
N/A
2.2
N/A
N/A
N/A
N/A
N/A
N/A
0.102
N/A
0.105
0.17
0.14
0. 15
HC
pprc
550
3-0
400
800
900
1000
H/A
N/A
100
1000
1000
500
500
350
450
95
240
250
240
280
400
N/A
200
180
80
55
20
30
1000
300
500
300
50
300
16-
400
400
350
700
440
400
NO
ppm
260
220
290
260
260
260
N/A
N/A
250
190
210
210
290
N/A
270
260
260
260
280
260
200
N/A
280
290
350
370
280
270
260
320
310
370
N/A
310
320
320
320
310
310
320
320
C02
11.9
11.1
11.5
12.5
13.
13.
N/A
N/A
N/A
N/A
15.
11.8
11.
11.
10.5
10.8
10.7
10.9
11.
10.7
11.
N/A
11.
11.5
11.3
11.5
8.
8.2
15.
13.8
13. B
N/A
13.
12.
11.6
13.5
13.8
13.6
13.7
13.3
13.4
Steam
Ib/hr
3600
3500
3500
3400
3550
3700
3750
3750
3800
3700
4100
4100
4100
3800
3800
3800
3900
3900
4100
4000
3900
N/A
3800
3850
3600
4000
4000
4000
4000
4000
4000
4000
N/A
3900
4200
4200
4200
4100
4200
4100
4050
CBC
4
6.8
7.8
3.5
2.5
2.0
1.6
1.1
0.
0.
0.
0.
5.
2.5
2.
2.6
1.8
3.
2.7
2.4
2.6
4.5
N/A
0.
0.
1.2
1.2
1.6
1.3
1.2
1.3
1.5
1.5
2.0
2.4
2.
3.2
3.
3.
3.5
N/A
N/A
CBC
Ash
C, 4
N/A
N/A
12.9
N/A
39.
N/A
31.3
N/A
N/A
N/A
N/A
N/A
30.3
26.2
N/A
N/A
.24,.
18.5
17.6
15.8
7.75
N/A
38.7
35.8
36.1
N/A
N/A
10.4
26.0
N/A
N/A
N/A
N/A
N/A
46.6
44.6
47.7
62.5
66.
56.5
59.6
CBC
S02
N/A
N/A
1.0
2.4
2.4
3.4
0.72
0.4
N/A
N/A
6.
0.5
N/A
2.4
1.6
1.5
1.0
1.5
0.4
1.25
>0.5
N/A
4.
4.
t.
3.3
0.75
1.6
3.
2.8
4.4
4.1
H/A
0.96
0.84
1.2
1.1
1.2
1.05
1.5
1.35
CBC
Air Rate
Ib/hr
N/A
577
N/A
N/A
580
N/A
777
N/A
N/A
N/A
N/A
914
912
N/A
N/A
929
949
920
867
730
^900
N/A
617
897
924
924
N/A
622
640
N/A
609
N/A
N/A
850
605
750
N/A
727
727
727
722
NR * nonregenerating portion of test
N/A - Not Available
*Anlysis performed on material collecteti by dust collectors only.
-------
TABLE 6. (Continued) Page 2.
Day Time
3 8 1823
1908
2015
2110
2218
2315
9 0001
0137
0200
0300
0400
0450
0610
0720
0800
0840
0949
1051
1152
1312
1409
1524
1644
1745
1845
1945
2100
2130
2300
2400
10 0100
0200
0250
0400
0300
0630
0730
0821
0930
1029
1128
FBM
S Input
in Coal
Ib/hr
22.5
22.5
21.
N/A
23.5
N/A
23.1
N/A
N/A
N/A
N/A
N/A
25.
•v.25.0
N/A
N/A
23.6
22.5
25.2
18.76
N/A
Banked
24.5
• N/A
N/A
N/A
N/A
24.6
24.7
N/A
27.
27.4
N/A
N/A
27.
27.
26.6
N/A
28.8
26.6
25.9
CBC
Temp.
Of.
1990
2000
2000
1980
1955
1950
1840
1920
1900
1890
1940
1980
1910
1840
1920
1940
2000
2000
2020
1995
1935
N/A
1930
1920
1980
2110
2010
2020
1980
1900
1850
1910
1930
1840
2010
1980
1990
2005
2010
2040
2030
CBC
NO
pprn
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
790
860
1080
1000
1200
N/A
N/A
4VO
480
500
140
780
N/A
465
N/A
290
N/A
N/A
N/A
N/A
N/A
N/A
90
70
110
100
CBC
(C02& S02)
%
N/A
N/A
19.4
22.
22.5
23.
19.7
N/A
25.
25.
25.
17.
N/A
23.
23.
22.5
17.5
20.5
17.3
21.
N/A
N/A
>25.
>25.
>25.
23.3
21.
21.
25.
25.
25.
25.
23.
21.5
20.5
20.5
19.8
19.
17.8
19.2
19.7
System
C Burnup
% Overall
N/A
N/A
97.7
N/A
92.3
N/A
96.7
N/A
N/A
N/A
N/A
N/A
•v.93.3
•v.94.5
N/A
N/A
97.3
N/A
97.5
97.1
98.7
N/A
93.6
91.4
90.5
N/A
N/A
98.2
95.6
N/A
N/A
N/A
N/A
N/A
91.8
88.
'96.6
•^89.
••.96. 6
N/A
N/A
CBC
CO
%
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
0.7
1.4
0.28
0.07
0.07
N/A
N/A
0.55
0.46
2.2
2.7
0.37
N/A
0.4
N/A
1.7
N/A
N/A
N/S
NX*
N//
N/f.
0.71
0.90
0.63
0.84
CBC
HC
Ppm
N/A
N/A
100
300
350
400
300
N/A
N/A
N/A
100
0
N/A
20
10
15
16
16
26
28
N/A
N/A
180
90
30
22
30
50
120
130
70
80
N/A
50
70
110
120
100
130
80
58
FBM
Limestone
Add. Rate
Ib/hr
14
40
28
85
85
71
70
73
84
91
91
80
77
77
77
84
0.
0 to 22
17
27
22
N/A
36
60
N/A
3
9
e
N/A
N/A
42
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
3.0
CBC
Flue Gas
Ib/hr
N/A
N/A
•\-768
N/A
612
N/A
820
N/A
N/A
N/A
N/A
N/A
953
962
N/A
N/A
1192
N/A
922
745
976
N/A
653
969
992
N/A
N/A
828
683
N/A
N/A
N/A
N/A
N/A
631
795
•v774
•v774
773
N/A
N/A
CBC
Flue Gas
Sulfur
Ib/hr
N/A
N/A
8.5
16.2
16.2
30.7
6.5
N/A
N/A
N/A
63.
N/A
N/A
25.4
17.
19.7
13.1
19.7
3.7
10.3
3.6
N/A
28.8
42.7
43.7
36.
N/A
14.6
22.6
21.1
31.5
29.4
N/A
9.5
6.
10.5
9.4
10.3
9.0
12.7
11.4
CBC
Coal Rate
Ib/hr
8
e
e
e
e
e
e
16
16
16
16
16
16
16
16
16
e ,,
e
e
N/A
e
N/A
e
•1.66
•v27
N/A
N/A
e
e
15.8
15.8
15.8
N/A
N/A
29.
23.3
23.3
25.
33.4
33.4
33.4
CBC
CBC Bed
% S
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
4.75
N/A
3.18
6.50
5.31
N/A
N/A
5.44
6.62
6.84
N/A
7.91
N/A
N/A
5.01
N/A
N/A
N/A
N/A
2.47
N/A
N/A
N/A
N/A
N/A
0.80
1.48
N/A
0.65
0.48
N/A
N/A
Fly
% S
N/A
N/A
1.10
1.41
(2.45)
N/A
1.89
N/A
N/A
2.04
N/A
1.84
1.88
1.76
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
3119
2.18
1.90
2.36
0.84
2.05
1.91
Ash
% Ca
N/A
N/A
16.22
18.21
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
17.1
19.63
19.55
N/A
17.95
N/A
N/A
N/A
N/A
N/A
N//>
23.4
N/a
N/A
N/J>.
N/A
N/A
N/f.
N/A
N/A
11.39
N/A
N/A
N/A
N/A
Concentratioi
S02 Factor
CBC/FBM
N/A
N/A
40
104
104
154
33
18
N/A
N/A
400
20
N/A
80
80
100
50
67'.
131 "
33
25 a
N/A i
114 ^
133
160
165
38
70
120
100
191
205
N/A
64
50
48
43
48
42
63
52
-------
TABLE
Day
3 10
11
NR
NR
NR
NR
NR
NR
6. (Continued)
Tilflo
1243
1342
1441
1542
1646
1758
1057
1953
2045
2130
2245
2345
0030
0100
0200
0315
0400
0510
0600
0700
0830
0902
1003
1051
1200
1300
1414
1520
1630
1724
1B17
FBM
1535
1530
1530
1510
1535
14BO
1460
1520
1470
1470
1500
1540
1510
1515
1500
1515
1520
1540
1535
1500
1540
1540
1540
1550
1550
1520
1510
1470
1510
1490
Off
Bed
IS.
18.
18.
19.
18.
18.
17.
17.
17.
17.2
17.
N/A
15.4
H/A
15.4
N/A
16. «
N/A
16.5
17.7
N/A
N/A
16.
N/A
N/A
N/A
N/A
18.
18.
19.
19.
Cool Rate
see
568
558
N/A
N/A
510
540
540
N/A
N/A
594
594
590
N/A
590
N/A
N/A
N/A
N/A
595
595
650
656
635
N/A
N/A
467
525
550
534
Plug
Air Rate
7010
6840
6960
7040
£740
6916
7010
7066
7160
7324
7260
1.7350
7496
N/A
7270
N/A
7098
N/A
7320
7013
N/A
6953
6818
6818
6883
N/A
N/A
5983
7070
6890
N/A
Ply Ash
52.8
50.
40.1
N/A
41.6
47.3
39.2
N/A
N/A
N/A
46.9
43.7
N/A
N/A
N/A
33.
N/A
N/A
53.9
N/A
52.1
N/A
50.6
H/A
55.
N/A
41.7
56.6
25.7
54.4
N/A
Flue Gas
350
250
260
200
310
260
250
225
N/.\
100
100
N/A
150
250
160
270
160
230
230
140
220
200
N/A
295
200
200
150
360
1000
1600
N/A
Flue Gas
Ib/hr
-------
TABLE 6. (Continued) Page 4.
Day Time
3 10 1243
1342
1441
1542
1646
1758
1857
1953
2045
2130
2245
2345
11 0030
0100
0200
0315
0400
0510 .
0600
ofoo
0830
0902
1003
1051
1200
NR 1300
NR 1414
NR 1520
NR 1630
NR 1724
NR 1817
FBM
S input
in Coal
Ib/hr
26.4
25.5
25.1
N/A
22.5
23.
24.3
N/A
N/A
N/A
26.7
26.7
N/A
N/A
N/A
•v26.6
N/A
. N/A
N/A
26.7
26.8
29.2
29.6
28.6
•v.28.6
N/A
21.0
23.6
24.8
24.
N/A
CBC
Temp.
op
2080
2120
2080
2060
2020
2020
2050
2060
2000
1960
1930
1970
1940
1970
1930
1900
1880
2000
1890
1940
1970
1960
1880
1970
1975
2110
1960
2030
2060
2070
1950
CBC
NO
Ppm
95
550
1000
1000
920
660
820
840
N/A
610
460
N/A
N/A
250
45
N/A
70
60
190
310
N/A
270
0
90
N/A
N/A
880
N/A
1000
1080
N/A
CBC
(C02S S02)
%
20.5
23.
21.7
22.5
>25.
>25.
24.5
25.0
N/A
24.
24.
N/A
23.
N/A
22.
N/A
18.
N/A
N/A
25.
N/A
>25.
N/A
24.5
N/A
N/A
16.5
N/A
13.6
15.
N/A
System
C Burnup
% Overall
N/A
98.5
97.3
N/A
96.3
97.7
97.9
N/A
N/A
N/A
•^97. 8
•v.97.4
N/A
N/A
N/A
•v96.4
N/A
N/A
•^94. 3
N/A
96.3
N/A
•v.97.9
N/A
94.7
N/A
98.0
98.7
N/A
99.1
N/A
CBC
CO
«
0.97
0.035
0.0035
0.007
0.08
0.18
0.03
0.0105
N/A
0.35
0.26
N/A
0.35
N/A
O.B2
N/A
0.5
N/A
N/A
0.26
N/A
0.37
N/A
1.07
N/A
N/A
0.006
N/A
0.04
0.04
N/A
CBC
HC
pprn
40
23
e
N/A
5
28
62
42
N/A
12
3
3
0.5
4
12
150
250
100
40
1.5
H/A
300
8
8
e
20
47
150
100
170
10
FBM
Limestone
Add. Rate
Ib/hr
N/A
N/A
0
0
30
29
100
0
26
26
26
39
39
43
43
94
94
94
94
19
19
19
45
15
8
55
e
a
5
45
N/A
CBC
Flue Gas
Ib/hr
N/A
850
1010
N/A
734
1207
1120
N/A
N/A
N/A
800
•^828 10
N/A 11
N/A
N/A
681
N/A
N/A
N/A
774
777
N/A
887
N/A
1017
N/A
^930
1060
N/A
1176
N/A
CBC
Flue Gas
Sulfur
Ib/hr
10.5
12.3
12.3
6.9
11.8
31.9
29.
24.7
20.4
14.8
16.1
18.3
15.2
13.7
10.2
9.8
9.8
7.8
16.8
17.1
15.4
15.0
17.6
15.2
6.3
1.6
0.7
0.33
N/A
N/A
N/A
CBC
Coal Rate
Ib/hr
e
0
e
N/A
33.
33.
N/A
yes
8
11.8
11.8
11.8
11.8
16.7
38.
20.
40.
e
9
e
e
N/A
T
N^A
N/A
N/A
N/A
N/A
CBC Bed
% 5
N/A
N/A
N/A
N/A
3.35
N/A
2.01
N/A
H/A
J/A
H/A
0.57
N/A
N/A
N/A
N/A
N/A
11/A
N/A
N/A
0.53
N/A
0.32
H/A
0.36
H/A
N/A
N/A
N/A
H/A
M/A
CBC
Fly Ash
% S
N/A
2.53
0.96
N/A
0.77
1.08
1.22
N/A
N/A
N/A
1.19 .
1.42
N/A
N/A
N/A
1.2
N/A
N/A
N/A
1.85
1.62
N/A
1.89
2.16
N/A
N/A
0.70
0.64
0.75
0.84
N/A
% Ca.
N/A
N/A
N/A
N/A
N/A
22.1
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
10.73
N/A
9.39
N/A
N/A
N/A
19.95
12.22
17.41
13.5
N/A
Concentration
SO2 Factor
CBC/FBM
36
52
43
38
47
93
94
89
N/A
200
190
~130
133
72
88
49
82
46
87
150
82
88
•v61
53
28
7
5
N/A
N/A
N/A
N/A
-------
6-32
yielding an average of 92.7%. Continuous capture rate
recording was not possible, since the single SO- analyzer
must intermittently be switched to CBC gas analysis.
The 98% carbon burnup goal (based on ash carbon analysis)
was met or exceeded in 6 of the 34 burnup determinations.*
97% burnup was exceeded in 15 of the 34. In order to
exceed 98%, it is necessary to avoid coal feed to the
CBC, since the CBC combustion efficiency is rarely over
96%. In 10 of the 34, FBM fly ash carbon content**
data appear to be in error based on ingredients rates
and flue gas analysis. The carbon burnup vs. steam rate
data are given below in Figure 19. The SO- concentration
factor = 100 goal was met furing less than 50% of the test
duration. During 98% carbon burnup conditions in the cur-
rent apparatus, it is difficult to achieve as high as
3% SO_ in the CBC regenerator off-gas (See Figure 20).
The CBC air rate requirements for effective carbon burnup
are the major factor in the problem. When coal was not
being fed to the CBC, sulfur balances tended to be poor.
This indicates that our measurement of a 3% SO_ gas, using
an 0.5% SO-instrument with N_ dilution, was unsatisfactory
(the unit was later modified to 5.0% range).
The best compromise operating conditions in the appa-
ratus used in this test appear to be exemplified by
2.4% SO_, 97.9% burnup; and 1.6% SO_, 98.2% burnup. These
data are summarized in Table 7.
*Note that carbon burnups reported are probably con-
servative since no ash weighing system was provided.
Final fly ash analytical results are listed in
Appendix K.
** Fluctuations in FBM ash carbon content may be due to
the current procedure of screwing the contents of the
first (coarse) dust collector into the second (fine
fraction) dust collector without a riffle box or
other systematic mixing device.
-------
6-33
4500
. 4000
UJ
cr
5
LJ
to
3500
200 25O 300 350
FBM CARBON BURNUP, Ib/hr
Q COAL RATE MEASUREMENT IS QUESTIONED
A BAD ASH % C
400
450
FIGURE 19. FBM STEAM RATE VS. CARBON BURNUP, RUN I68H
-------
6-34
100
U>
a>
t-
z
UJ
o
cr
UJ
a.
y 95
UJ
o
UJ
Q.
o
tr
^
CD
I
Z
O
CD
tr
<
o
2
UJ
o
Z
<
D
LOCUS OF DATA WITH ACCEPTABLE
CARBON BALANCES
Q
a
O.I
0.5
1.5
2.5
VOL. % IN C8C GAS
FIGURE 2O. FBM-CBC SYSTEM CARBON BURNUP VS.
CBC S02 OUTPUT, RUN I68H
Q BAD ASH C ANALYSES
A COAL FEED TO CBC
-------
TABLE 7.
6-35
SUMMARY OF BURNUP DATA, RUN 168H
Point
No. **
30B
28B
27B
19A
21A
25B
7A
35A
3A
39A
37A
3B
20A
16B
2B
SB
28A
23B
2 IB
11B
12B
19B
7B
5A
14A
6B
29A
23A
24A
38A
13A
17A
25A
36A
CBC
02
%
5.
3.
3.
2.4
4.5
N/A
1.1
2
3.5
3.5
3.
1.5
2.6
2.5
1.8
0.5
1.3
0.3
0.6
1.0
0.8
1.3
0.4
2.0
2.0
0.1
1.2
0.
0.
3.
2.5
3.
1.2
3.2
CBC
SO 2
%
0.03
0.03
0.07
0.4
0.4
0.56
0.72
0.84
1.0
1.05
1.1
1.1
1.25
1.3
1.3
1.45
1.6
1.8
1.8
1.9
2.0
2.0
2.35
2.4
2.4
2.4
3.0
4.
4.
1.2
N/A
1.0
4.
1.2
CB
Overall
%
99.1
98.7
98.0
97.5
98.7
94.7
96.7
91.8
97.7
96.6
96.6
97.3
97.1
96.4
98.5*
96.3
98.2
97.9
96.3
97.8*
97.4
94.3
97.9
92.3
94.5
97.7
95.6
93.6
91.4
-v89.
^93.3
97.3
90.5
88.
CBC
T, °F
2070
2030
1960
2020
1930
1970
1840
2010
2000
2010
1990
2080
1990
1900
2120
2020
2020
1980
1970
1930
1970
1890
2050
1955
1840
2020
1980
1930
1920
2000
1910
2000
1980
1980
S Capture
in FBM
%
N/A
89.3
94.5
90.1
93.8
94.3
93.3
94.5
91.6
92.7
91.6
91.5
84.6
91.4
92.1
89.3
92.1
92.
93.3
96.8
95.
92.8
91.6
93.1
N/A
90.8
91.4
88.6
89.9
92.2
90.8
94.
91.6
92.
Remarks
Non-regenerating
Non- regenerating
Non-regenerating
C imbalance***
C imbalance
Coal to CBC
Coal to CBC
Coal to CBC
Coal to CBC
C imbalance
C imbalance
C imbalance
C imbalance
Coal to CBC
Coal to CBC
C, O imbalances
Coal to CBC
C imbalance
Coal to CBC
Coal to CBC
C imbalance
C imbalance
Coal to CBC
C, O imbalances
Coal to CBC
Coal to CBC
Coal to CBC
These carbon burnup values based on flue gas analysis were not
corroborated by heat balance calculation (See Section 6.3.2.2).
**See Table 6
***Flue gas rate and analysis not in agreement with ash-based
burnup calculation.
-------
£--36
Based on the test results, process equipment
redesign was undertaken, based on a 3-reactor SO Acceptor
Process concept: (1) FBM as currently set up; (2) an
auxiliary fluidized regenerator of 0.6 ft cross section,
burning coal and regenerating bed material pumped from
the FBM; (3) a conventional CBC, essentially as currently
set up, burning fly ash from (1) or from both (1) and
(2). The bed temperatures of (2) and (3) are essentially
equal. Regenerated bed material from (2) returns by fluidized
flow or gravity to either (1) or (3).
One possible way to promote carbon burnup under
regenerating conditions is to operate the CBC at higher
temperatures*. For example, at 2120°F, 98.5% burnup
was achieved at 1.3% SO_ and 1.8% O_ (see point 2B on
Table 7). At 2020°F, 98.2% burnup was achieved at 1.3%
O_ and 1.6% SO_ (see point 28A). CBC temperatures above
o
2120 F were avoided in this test (with a refractory
lined CBC) from fear of. slagging. However, the ash
fusion behavior in a large excess of lime environment,
when no coal is fed has not been studied, and this fear
may be unfounded. The 3-reactor concept, of course,
obviates excessively high CBC temperatures. The boiler
system was operated for over three days during run 168H,
with one interruption of about 2 hours' duration. There
were two minor upsets of 2 and 15 minutes when steam
production rate was reduced but did not cease. In each
instance of SO capture below 90%, temporary limestone
feed disorders were the sole cause. Bed material for
analysis is removed at about 10 Ib/hr; this, plus attri-
tion, must be made up intermittently as needed to trim
* Another way is to reduce gas velocity, by building
a CBC of larger grid area.
-------
6-37
FBM bed level, bed temperature, and capture percentage.
Based on the test results (see limestone rate data in
Table 6) a limestone requirement for this process of
about 5% of coal rate appears reasonable. The Powhatan
coal did not appear to add oversize inert matter (rocks)
to the bed. Calculations based on fly ash calcium analyses
tended to confirm the 5% value. The coal particle size
analysis for this test is shown in Figure 21.
6.3.2.2 Gaseous Emissions (Run 168H)
The available FBM SO emission vs. bed composition
data are collected in Figure 22. There is considerable
data scatter due to the combined effects of temperature,
O_ level, and previous history of the lime bed. While
there is an incentive to keep bed sulfur below 3.5%
to guarantee good capture, it is interesting to note
that good capture can still be obtained at sulfur levels
above 7% with continuous passage of bed material through
a regenerator. This capability is not observed in batch
experiments where bed sulfur continually rises, and is
probably due to surface or "superficial" regeneration of
particles. The small particles having higher surface
area/volume ratio, are more completely regenerated and
are more fully activated for subsequent SO_ capture.
The larger particles retain a core of higher CaSO.
content* after passing through the regenerator, and bias
the overall analytical result, so that a 7% sulfur bed,
after undergoing heavy regeneration, can behave as a
3% sulfur bed would after undergoing a lower level of
regeneration. Evidence was presented in Section 6.2,
page 6-20 for lower S/Ca ratio on the surface of particles
See discussion of the D.E.S. concept ir. Section 6-2,
page 6-20.
-------
6-38
10 20 3040 50 60 70 80
90
95 97.
4
8
10
UJ
S l6
18
z 20
£ 30
o
(o
e 50
en
3 100
200
\
s
v
s
s
N
K
1
V
>
t
1
s
^ ,
1
s
\
\
\
\
^
§ i
o
PARTICLE SIZE, MICRONS
.1 I 5 15 25 35 45 55 65 75 85 90 95
OVERSIZE % BY WEIGHT
FIGURE 21. COAL PARTICLE SIZE ANALYSIS, RUN I68H
-------
300
FBM 200
S02
PPM
100
X
X
*
X
*
4% 02
>T 151
X
*
fl
1535° F
I450°F
3.2 02
41475° F
2.7 02
4.7 02
34567
FBM BED SULFUR, WT %
8
FIGURE 22. FBM FLUE GAS SOe CONCENTRATION VS BED SULFUR
CONTENT - RUN I68H
en
I
OJ
-------
6-40
in an aged bed. A few selected bed samples were screened
and the screen fractions analyzed for S, but results were
inconclusive.
At one point in the test, salt feed to the FBM
was performed (see Table 6). The immediate effect was
an increase in coal combustion efficiency in the FBM
causing the CBC temperature to drop since it was suddenly
starved for fuel (and its 0_ output to increase) and the
S0_ output of the FBM was immediately cut by 47%, confirming
the earlier FBC test results with salt feed.
The available FBM NO emission vs. bed composition
data are collected in Figure 23. By operating at 6%
sulfur, rather than 1% sulfur, a 10% reduction in FBM
NO emission is predicted. Note from Table 6, however,
that a typical split of NO emissions is 2/3 FBM, 1/3
CBC due to the higher temperature and O_ level of the
CBC. Several possible explanations for the slight but
noticeable effect of bed sulfur level and NO production
exist:
a. The effect is real, and possibly due to oxi-
dation of nitrogenous matter in coal by CaSO. in a mild
fashion yielding N_ rather than NO, in relatively oxygen-
deficient regions of the bed near the coal feed points.
b. The effect is real and possibly due to stickiness
of CaSO. laden particles to coke balls thereby covering
the balls and affecting coal activity. Low coal activity
would reduce "hot spot" NO generation. Such coke balls
are frequently observed with lime particle coats when
large lump size coal (3/4") is fed. The formation of a
low melting CaS - CaSO., - CaSO. transient eutectic appears
possible.
-------
FBM
NO,
PPM
200
100
34567
FBM BED SULFUR, WT %
8
FIGURE 23. FBM FLUE GAS NO CONCENTRATION VS BED SULFUR CONTENT
RUN -I68H
I
.fc.
-------
6-42
c. The apparent effect is an artifact due to
lower bed mean particle size at higher CaSO. content
and particle density. A bed of smaller particles could
lead to greater temperature uniformity and reduction of
hot spots believed to generate excess NO. A set of experi-
ments in the FBC using coarse and fine limestone fractions
could be expected to elucidate this possibility. Experi-
ments on NO reduction had been scheduled as part of this
contract but were cancelled when the long duration run
became the priority item.
Condensed data on CO and HC emissions in this
run are in Appendix I.
6.3.2.3 Boiler System Heat Balance
Heat balance data for several selected conditions
during this run are collected in Table 8. Only those
conditions in which carbon-burnup efficiency as calculated
from ash carbon analyses and/or flue gas analyses is
above 96.5% are treated here. When coal is added to the
CBC, heat balances tend to be poorer. The reaction heats
of calcining fresh limestone, and capture and desorption
of SO2 are accounted for under a term, "net CaSO. formation.
The unaccounted heat is correlated with apparent thermal
efficiency of the system in Figure 24. This indicates
that the major measurement errors are occurring in the
ingredients rates and steam output, rather than in the
individual loss terms.. Recorders for coal rates and air
rates would have provided a more accurate record than data
logged by operators.
6.3.2.4 Bed Specific Gravity
Since bed sulfur content seemed to have a noticeable
effect on pollutants release rate, it was hoped that bed
specific gravity could be used to provide a quick, on-line
-------
TABLE fl. MEAT BALANCE DATA RUN 1B8H
Daj
FBM bed temperature. °F
SO] Capture. «
Cartoon Bumup, %
CBC SO,. »
Total preheated air, Ib/hr
Total preheated air. KBlu/hr
CBC Air. lb/br
Coal Heating value. KBtu/hr
Net CaSO. Formation. KBW/hr'
Total Input. KBtu/hr
Steam abaorptton. KBlu/hr
Circ. H2O abaorpllon KBtu/hr
Total (FTjOtSl. )
7t of coal value
Flue Gae loee. KBtu/hr
Slowdown allowance KBtu/hr
Hvdrocarbon loao. KBtu/hr
CO loes. 'KBtu/hr
CBC Carbon Iocs KBtu/hr
% or coal value
Hot Flj A«h loai KBtu/hr •
FBM collector duet cerbcn loea
Nitric oride loaa KBtu/hr
Lift air loaa KBtu/hr
Radiation ft unaccounted
Total Loeete
a Rate Queetloned
U
1
1490
40
99.1
0.03
6000
473
1110
8963
26
7483
4161
160]
5764
S3. 1
041
11
117
9
31
0. 73
21
102
N/A
N/A
N/A
1263
11 U
141
1470 1310
88. 3 94. 3
98.7 96
0.03 0.07
6600 6600 •
191 391
990 090
6887 6109*
N/A 116
7233 6623
3939 3939
1538 1336
5497 3497
80. 1 80«
695 626
20 20
115 110
11 79
72 99
1.09 1.6
19 15
116 3:
1 H/A
39 H/A
176 H/A
1366 1001
t> 9 »
U51 1404 0007
1475 1483 1550
61 83.8 93.3
93 96.7 98.7
.4 0.4 0.72
•760 7000 6400
45 302 444
104 730 730
723 7129" 672S««
10 112 37
781 7741 7221
415 3959 '3806
167 1339 1841
363 5318 3447
83 77,4 81
90 815 740
1 20 12
2 18 64
37 132 188
14 74 174
.0 1.04 1. J»
1 11 14
9 116 101
N/A 3
3 36 37
H/ 386 H/A
161 1711 ISM
& 9
1340 1480
91.6 94
97.7 97.3
1 0 1.0
6400 1000
459 <3t
730 1110
6099 1076
30 60
«5BB 1»0
35« 3851
1417 1464
4963 5415
81.7 76.5
740 834
21 23
31 "
188 361
111 288
1.1 >•'
18 23
SS 103
3 4
" Jo
N/A «•
1310 1116
Coal to CBC-v
10 10
0930 1441
131 1330
97 91. A
86 n.l
U 1.1
770 7530
47 410
89 080
816 72J9
N/A H/A
884 rfil
425 4103
195 1531
580 3633
71 71. 5
93 681
1 84
6 14
84 57
23 157
6 1.1
1 20
98 98
4 4
34 11
902 409
1359 1640
*
10 .
0730
1J30
91.6
98.8
7800*
90S
680
8048
S/A
6563
4195
1636
• 6911
71.4
838
19
J9
31
118
1.7
21
71
4
31
106
1113
Coal to CBC
9
1311
1480
86.0
97.1
1.15
69(0.
301
880
6948
87
7900
4031
1318
5970
80.1
761
14
13
183
124
1.8
14
U
4
36
130
1429
1342 2130
1530 1340
91.1 91.1
96.4 98.1
1.3 1.6
7500 1300
461 „ 376
950 770
7430 1166
N/A N/A
7609 1141
4161 4060
1532 1940
3714 3800
76.9 18.6
881 1063
31 11
11 3
43 97
MS 101
3.36 1.4
14 14
D7 67
H/A I/A
40 44
in 185
1712 1388
1003 3143
1340 1500
81 96,6
91.9 95.4
1.6 1.9
7400 7830
436 454
91C- 710
8381 a 7914
»/A B/A
9019 8314
4869 4051
19U 1611
6131 5673
71.9 71.6
846 883
14 14
17! 30
105 139
143 361
1.69 4.3
31 23
106 83
3 4
18 32
631 M3
1JJJ 3347
1345
1540
93
97.4
1
7700*a
437
760
7914
17
6394
4235
1583
5840
73.7
900"
U
46
35
160
1
10
73
4
31
781
1091
^
1637 1738 AVG
1480 1480
91.6 90.8
97.9 97.6
1.15 3.4
7700 7800
438 431
1060 1090
7495 7103
-103 -93
7826 1435
4052 4003'
1612 1624
5664 5626
71.8 79.1
914 937
14 . 21
34 48 »
» " •.!. •
124 132 *6B
1.64 1.7
20 20
36 7S
N/A 4
38 37
476 61
17J4 1378
v ^_
-------
6-44
1000
900
70 75 80
CALCULATED THERMAL EFFICIENCY,
Q COAL RATE DATA
QUESTIONABLE
85
90
FIGURE 24. HEAT BALANCE-RUN 168 H
-------
15-45
estimate of sulfur content without waiting for chemical
analysis. Specific gravities (poured without tapping)
were measured in a 100 cm graduate*. Data are collected
in Appendix J. Although the highest bed densities occur
at the higher sulfur contents, data scatter is excessive
and other factors (particle size, shape, coal ash ingredient
contaminants) must also affect bed density. In most cases,
CBC beds are about 2% less dense than FBM beds, in line
with the effect of regeneration.
Bed particle size distributions for this run are
described in Section 6.4.
6.3.2.5 Sulfur Analyses
Fly ash and bed material sulfur analyses were
performed on many of the FBM and CBC samples (see Table
6). Most fly ash samples are below 2% S. The highest
FBM fly ash S% is 1.89 (average 1.69). When coal is fed
to the CBC, the CBC fly ash tends to be high in both C
and S, implying that some S is being eliminated from the
system, tied up in pyrolized coal material rather than
as CaSO.. Under regenerating conditions, without coal
feed to the CBC, its ash contains less S than the FBM
fly ash input; i.e., the calcium sulfate content of
the ash is being "regenerated" along with the bed material.
The average CBC ash is 1.53 %S, 10% less than the FBM
• -f
average. For comparison, the fly ash S contents of the FBC
coal-burning nonregeneration tests previously discussed
were all over 2.2% tending to indicate that fly ash from
the regenerative FBM operation contains surface material
abraded from the low-sulfur surface layer of the partly
regenerated particles.
* Ash particle densities can be measured using water
displacement, but lime particles react with water.
-------
6-46
The sulfur balance data for this run are presented
in Appendix F, Table 2. The imbalance is 15%. The
principal error is believed to occur in the measurement
of CBC S02 level, using the 0.5% instrument with N_ dilu-
tion (see Section 5.5). A 17% error in measuring a 3%
SO- stream by this technique appears reasonable; thus
it is assumed that the measured 92.7% S capture in the
FBM is accurate. This instrument has since been converted
to triple range operation: 0-900, 0-5000 ppm and 0-5%
SO_. A separate continuous 0 to 10% SO_ regenerator flue
gas recording analyzer is required. This would have
permitted on-line FBM and REG/CBC SO- recording without
interruptions.
6.3.3 FBM/CBC Run with Coal Feed to CBC (Run 169H)
Following completion of the FBM/CBC long duration
run, another test was made in which FBM fly ash was
discarded and the CBC run on coal only. CBC temperatures
were in the 1950-1990°F range. Six levels of CBC O2
were studied between 0.4 and 3.0 percent. Used lime bed
material from the previous run (168H) continued to give
good SO_ capture in the FBM.
The carbon combustion efficiency of the CBC burning
coal only during test 169H is plotted versus the oxygen content
of the CBC off gas in Figure 25. These points are shown
as circles. It may be seen that there is a poor correlation
between O- and carbon combustion-efficiency. In experi-
ments described in Reference 1 a similar burnup efficiency
was achieved at a far higher O_ level (Test C-319-5,
95.5% efficient, 9.8% 02, 1980°F).
Data from Test 168H is also potted on Figure 25.
For this test the carbon burnup efficiency values are
for the tandem system, i.e. CE = (Carbon burned/Carbon fed
-------
6-47
100
m
A
I 2 3 3.5
or more
02 IN CBC FLUE GAS , VOL.%
KEY
CBC - CARBON COMBUSTION EFFICIENCY -
RUN I69H (COAL FEED ONLY, NO FLYASH)
TANDEM CYCLE- CARBON COMBUSTION EFFICIENCY -
RUN I68H (FLYASH ONLY FED TO CBC)
TANDEM CYCLE COMBUSTION EFFICIENCY -
RUN I68H (COAL 8 FLYASH FED TO CBC)
1970°= CBC BED TEMPERATURE
FIG. 25.- CBC CARBON BURNUP AND TANDEM CYCLE
BURNUP VS. OXYGEN LEVEL
-------
6-48
to FBM and CBC) x 100. When fly ash was the sole fuel
to the CBC (points shown as squares in Figure 25) the
combustion efficiency was relatively high. When coal was
added to the CBC (points shown as triangles) the systems
carbon combustion efficiency was lower.
The band shown on Figure 25 shows the data scatter
which probably results from differences in factors such
as fly ash feed particle size, carbon content and feed
rate.
Sulfur balances with coal feed appear good (see
Table 9). Unlike the situation with fly ash feed for fuel,
good regeneration with coal feed can be achieved at any
CBC flue gas O_ level between 0% and 3%.
The goal of CBC flue gas being 1/11 of the total
FBM plus CBC flow (or less) was not achieved, limiting
CBC SO_ level to less than 3% in several cases. As
previously, the CaCO, makeup needs are modest (20 Ib/hr
including bed sample withdrawals). The CO analayzer
was inoperative for this test. The desired design data for
the 3-reactor concept (see below) were obtained from this
test.
6. 4 Three-Zone SO,, Acceptor Process MK II Concept
(FBM/CBC/REG)
6.4.1 Background
The difficulties encountered in attempts to con-
tinuously regenerate sulfated lime beds in the existing
CBC have been described. The CBC flue gas SO- levels
were usually too low. The existing CBC was designed to
achieve high carbon burnup efficiency in a high-velocity
apparatus using high excess air levels. Due to the labora-
tory structural layout, it was not feasible to rebuild
the CBC to utilize low-velocity combustion at lower
excess air levels to achieve the same burnup efficiency.
Another alternative which could not be pursued was
-------
6-49
TABLE 9. FBM-CBC RUN WITH COAL FEED TO CBC (Run 169H)
CONDITION NO.
FBM: 02, %
HC , ppm
SO2, ppm
NO , ppm
C02, %
Steam, Ib/hr
Fly Ash C,
Air Rate,
CBC: T, °F
02, %
*CO2 , meter
HC , ppm
SO2, %
NO , ppm
Air Rate ,
S Emission
Fly Ash C,
C Burnup,
CBC only
%
Ib/hr
(%)
Ib/hr
, Ib/hr
%
%
Coal Input Rates FBM
Ib/hr:
Bed Circulation
Ib/hr
CBC
Rate,
1
8.
120
25
330
10.
3700
45.
9080
1990
0.
>25
60
3.
390
955
38
24.
94.
620
82
380
5
5
5
4
6
3
5
2
8.5
120
25
330
10.5
3700
48.8
9080
1990
0.6
>25
60
3.6
390
955
38
19.8
95.8
620
82
380
3
8.5
120
25
330
10.5
3700
47.5
9080
1990
0.8
>25
60
3.6
390
955
38
17.9
96.3
620
82
380
4
7.8
130
150
290
9
3700
51.3
8950
1970
1.2-1.8
21.5
0
2.49
540
928
25.4
30.4
92.6
620
82
153
5
4.
800
80
250
11.
3450
62.
6114
1950
2.
19.
8
2.
580
897
19.
20.
95.
540
82
153
2
2
4
75
5
0
8
7
6
6
3.
800
80
250
11.
3450
N/A
6314
1950
3.
19.
10
2.
580
907
29.
18.
96.
540
82
880
8
2
0
5
95
5
2
2
Sulfur Input Rates
Ib/hr:
FBM
CBC
28
3.
7
28
3.7
28
3.7
28
3.7
23.
3.
4
7
23.
3.
4
7
Bed Material Bulk Density 73.5 lb/ft3
Limestone Makeup 164 lb/8 hrs.
*Includes SO2 response
(thermal conductimetric determination of CO2 in N2)
Note: N/A = not available
-------
6-50
increasing the bed level of the CBC. It is unfortunate
that this course could not have been pursued. The high
excess air principle was found to be incompatible with
bed regeneration when fly ash was the sole fuel. The
use of CBC coal feed to achieve rapid regeneration was
found to lower carbon burnup efficiency. PER reported
to EPA that EPA's goals (high carbon burnup, high regenera-
tor flue gas S0_ concentration, high FBM flue gas de-
sulfurization) might be met most expeditiously in a
3-zone apparatus in which the high-temperature functions
(burnup and regeneration) were physically separated.
6.4.2 Extended Run Test Operation (Run 171H)
Following construction and shakedown testing of
the three-reactor system, a long duration test of the
integrated FBM/CBC/REG pilot plant scale boiler system
in the SO., Acceptor Process MK II (three-cell) mode,
was conducted over an 8-day period. One hundred fifty-
six hours of boiler operation were logged. Coal fed was
37 tons. Average regenerator flue gas SO_ concentration
was 4.1% (volume) or 13.5 Ib/hour, (25 determinations),
as determined by IR and wet assay techniques. The FBM
SO2 capture and CBC burnup efficiency data are discussed
below. Some down time of the system resulted from (a)
the developmental nature of the regenerator (it was an
appendage to the FBM) and (b) lack of maintenance on the
FBM/CBC components due to budget constraints. See Test
Log in Table 10.
A 6.1% ash, 3.3% S coal (Powhatan) was used during
most of the test and its average carbon burnup exceeded
99%; the highest CBC output carbon level was 11.9%*
C level in CBC dust collector collected fly ash.
-------
TABLE 10
FBM/CBC/REG. SYSTEM IOC - Run 171H
Da_y
7/12
7/12
7/13
081
7/13
7/14
7/14
7/15
04
06
08
7/15
1000
1200
1400
1600
1800
2000
2200
2400
0200
O400
0600
30-oass
1000
1200
1400
1600
1800
2000
2200
2400
0050
0400
1800
2010
2200
2400
0200
00-0545
00-0745
00-0845
1000
1200
1400
1435
1000
1500
1560
1550
1630
1525
1480
1480
1500
1540
1480
1520
1540
1480
1460
1420
1500
1460
1460
1480
1540
N/A
N/A
N/A
1440
1520
1530
1450
1440
1470
1500
1460
1500
1*10
N/A
N/A
Bed
Level
H/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
16
16
N/A
N/A
N/A
19
N/A
N/A
iT.
N/A
H/A
N/A
N/A
N/A
(J/A
N/A
N/A
N/A
N/A
N/A
16
N/A
N/A
N/A
N/A
Coal
Rate .
N/A
445
550
510
425
405
421
476
480
465
481
451
409
445
475
460
497
474
515
499
N/A
N/A
N/A
500
505
518
488
600
508
500
560
600
N/A
594
N/A
Air
Rate
Ib/hr
N/A
N/A
N/A
N/A
4700
4540
4560
4590
4650
N/A
4770
5258
N/A
N/A
N/A
5650
N/A
N/A
6220
N/A
N/A
N/A
N/A
N/A
N/A
N/A
H/A
N/A
N/A
67JO
H/A
tl/A
II/A
N/A
N/A
Fly A»h
Carbon
N/A
N/A
N/A
N/A
49.6
29.7
N/A
36.5
53.5
N/A
N/A
49.9 (845)
36.2
27.5
N/A
N/A
N/A
21.8
27.3
N/A
N/A
N/A
N/A
N/A
N/A
24.6
N/A
N/A
N/A
50.4
50.6
46.8
N/A
N/A
N/A
F 1 uc Ga i
SO, ppa
^ rr^
820
600
320
410
370
310
200
S/A
N/A
220
220
550
460
280
200
N/A
200
N/A
360
350
240
N/A
N/A
160
190
225
270
N/A
240
270
320
600
500
600
N/A
CBC '
Flue Gas
Rate
Ib/hr
N/A
N/A
N/A
N/A
5090
4880
4920
5030
5020
N/A
5150
5594
N/A
N/A
N/A
6010
N/A
N/A
6600
S/A
N/A
N/A
N/A
~ 6600
~ 6GOO
~6600
,.6600
-6600
,-,6600
7120
~7120
~7120
~7120
~7120
N/A
Coal
•• Syste
Bed S
_ — £_
N/A
N/A
N/A
N/A
3.6
3.61
4.51
N/A
3.48
4.86
N/A
3.56
3.57
3.92
N/A
N/A
N/A
4.74
3.74
N/A
N/A
N/A
N/A
N/A
1.32
N/A
3.48
2.98
3.4
3.7
4.96
3.22
N/A
N/A
N/A
Feed Plug.
m Burnup
Bed Ca
!
N/A
N/A
N/A
N/A
N/A
13.3
N/A
N/A
N/A
N/A
N/A
N/A
23.22
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
22.58
N/A
N/A
N/A
N/A
N/A
17.37
N/A
N/A
N/A
N/A
n
°2
2L
3.5
7.
4.
6.5
3.5-
4.6
4.3
4.2
2.6
3.0
3.0
3.0
2.6
3.5
4
3.3
3.3
3.4
3.4
2.9
5
N/A
4.2
6.5
6
5.5
6.3
6.5
8
4.5
3.6
5.7
a
N/A
N/A
CO
2<
0.22
0.06
N/A
0.16
0.12
N/A
0.22
N/A
N/A
N/A
0.2
0.35
0.45
0.4
0.23
N/A
0.45
N/A
0.42
0.5
0.15
N/A
N/A
0.12
0.11
0.13
0.1
N/A
N/A
0.06
0.2
0.4
0.3
N/A
N/A
HC
PP"1-.
1600
1400
1500
1400
1000
1000
1000
N/A
N/A
N/A
1900
1500
1700
1500
1000
N/A
1600
N/A
1550
1800
N/A
N/A
N/A
270
300
450
250
N/A
N/A
750
600
2500-
1600
•
N/A
CO
co2
f
17.
15.
N/A
9.
8.5
8.3
8.2
N/A
N/A
9.5
9.3
9.0
9
9.1
9
N/A
9.3
N/A
10.1
11
10.7
N/A
N/A
12.
12
12.2
12.
N/A
8
9.
9.5
6
6
8
N/A
Steam
Ib/hr
N/A
3200
3700
3200
3000
3100
3200
3100
3200
3100
3100
3100
3100
3100
3200
3250
3250
3300
3300
3300
3500
N/A
N/A
3050
2950
3350
3350
3350
3350
3150
3400
3400
3400
3500
n
°2
S_
N/A
N/A
N/A
N/A
5.
3.
N/A
2.
5.
6.
N/A
N/A
N/A
7.5
N/A
5.
N/A
7.
7.
N/A
N/A
N/A
N/A
N/A
5
N/A
7
N/A
N/A
N/A
4
N/A
N/A
N/A
N/A
Ash
C_,_Jt
N/A
N/A
N/A
N/A
N/A
0.
1.
N/A
1.
1.
N/A
1.
1.8
0.7
N/A
N/A
N/A
0.5
1 .
N/A
N/A
N/A
N/A
N/A
N/A
3.9
N/A
N/A
N/A
11.95
2.6
3.9
N/A
N/A
N/A
ppm
N/A
N/A
820
N/A
N/A
600
2500
1400
1000
240
270
410
370
1200
5000
500
100
1000
190
N/A
N/A
N/A
N/A
700
110
130
2SO
N/A
N/A
Jft
N/A
N/A
Air
Rate
Ib/hr
N/A
N/A
N/A
N/A
1176
1167
1264
1035
1114
N/A
990
975
N/A
N/A
N/A
1245
N/A
N/A
1315
1315
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
1300
N/A
N/A
N/A
N/A
N/A
C "
Burnup
X
Overall
N/A
N/A
N/A
N/A
N/A
99.
99.
"a.
99.
N/A
S/A
99.
99.
N/A
N/A
N/A
99.
99.
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
99.0
99.
99.
N/A
N/A
N/A
N/A
1920
1950
2010
1970
1960
1940
1920
2010
1920
1980
1950
1960
I960
1970
1970
1960
1980
1980
N/A
N/A
N/A
N/A
1960
2010
1920
1930
1800
1870
2020
2030
2000
N/A
N/A
CO
EES
N/A
1300
1200
1100
400
360
280
900
80
60
50
30
100
50
50
100
55
60
50
N/A
N/A
N/A
N/A
90
150
60
50
N/A
N/A
N/A
100
N/A
N/A
N/A
UU2
_2 —
N/A
N/A
N/A
N/A
N/A
N/A
N/A
15.
N/A
N/A
14.
S/A
12.5
N/A
N/A
13.
N/A
13.
N/A
N/A
N/A
N/A
N/A
N/A
12.
N/A
13.
N/A
N/A
N/A
11.5
N/A
N/A
N/A
N/A
HC
ppm
N/A
N/A
N/A
N/A
N/A
N/A
N/A
100
N/A
N/A
0
S/A
90
N/A
H/A
50
N/A
48
N/A
N/A
N/A
N/A
N/A
N/A
150
N/A
85
N/A
N/A
N/A
160
N/A
N/A
N/A
N/A
-------
Day
7/12
7/12
7/13
7/13
7/14
7/15
7/15
Time
1000
', 1200
. 1400
1600
1800
2000
2200
0000
0200
0400
0600
0800
1000
12
14
16
1800
2000
2200
2400
0050
0400
1600
2010
2200
2400
0200
0400
0600
0800
1000
1200
1400
1435
1600
System
S Input
In Coal.
Ib/hr .
N/A
15.7
19.1
17.8
15.1
14.3
14.8
16.7
16.6
16.3
16.8
15.7
14.4
15.9
16.8
16.3
17.5
16.8
18.2
14.9
N/A
ri/A
H/A
17.1
17.6
16.3
17.3
21.0
18.0
17.7
19.7
21.0
N/A
20.6
N/A
CBC Bod
y s
N/A
N/A
N/A
N/A
2.75
3.64
3.6
N/A
2.99
2.66
N/A
3'o
2.86
4.16
N/A
N/A
N/A
3.78
4.05
H/A
. N/A
, N/A
N/A
N/A
3.56
N/A
N/A
N/A
N/A
2.39
3.62
2.88
N/A
N/A
N/A
ran
Limestone
Ib/hr
Feed Rate
N/A
153
100
6
64
60
120
84
16
52
50
1
68
91
104
65
44
0
0
95
0
N/A
34
34
64
63
68
53
53
53
98
60
N/f
N/A
N/A
CBC
Flue Cas,
Ib/hr
.N/A
.N/A
.N/A
N/A
1216
1207
1304
1075
1154
N/A
1030
1015
.N/A
N/A
N/A
1285
N/A
N/A
1355
1355
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
1300
N/A
N/A
N/A
N/A
N/A
System*
S
Release
Rate.
Ib/hr .
N/A
N/A
N/A
N/A
2.8
2.45
4.65
2.6
2.3
N/A
1.56
3.85
2.85 FBK
H/A
N/A
2.03
N/A
H/A
2.89
N/A
N/A
H/A
N/A
H/A
N/A
N/A
N/A
N/A
2N/57 <2'12
N/A """
N/A
N/A
N/A
N/A
Reg .
« «: Coal Rate,
Captur<
N/A
N/A
N/A
(System)
61.5
82.9
66.6
64.5
66.3
N/A
90.7
75.5
H/A
N/A
H/A
87.5
H/A
H/A
84.1
N/A
N/A
N/A
N/A
N/A
N/A
H/A
N/A
N/A
N/A
85.5
N/A
N/A
H/A
N/A
N/A
;
(FBM)
86.3
88.4
92.7
93.6
93.4
N/A
92.5
76.4
79
N/A
N/A
91.9
N/A
N/A
65.7
N/A
N/A
N/A
H/A
N/A
H/A
H/A
N/A
N/A
H/A
68.0
M/A
H/A
H/A
H/A
H/A
Ib/hr
H/A
N/A
N/A
tl/A
31
26
29
32
30
30
30
27
N/A
36
N/A
36
36
36
38
IVA
N/A
H/A
18
18
29
36
36
37
37
37
37
37
N/A
H/A
N/A
Reg .
Air Rate
N/A
H/A
N/A
N/A
262
257
257
261
262
N/A
261
241
N/A
H/A
N/A
336
N/A
H/A
336
H/A
H/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
Reg.
Fly Ash
C %
H/A
M/A
H/A
N/A
N/A
N/A
N/A
N/A
N/A
H/A
N/A
N/A
N/A
H/A
N/A
N/A
H/A
H/A
H/A
N/A
N/A
N/A
H/A
N/A
H/A
N/A
N/A
N/A
M/A
N/A
N/A
N/A
H/A
N/A
N/A
Hog.
T. JF
N/A
1900
1900
1900
1950
1950
1980
2040
1950
2040
2000
2000
1920
1800
1900
1940
1940
I860
1930
1900
N/A
H/A
N/A
1970
1940
2040
1970
2010
1980
I960
2010
2030
2020
;:/.-
r;/A
Flue Gas
SOj %
N/A
H/A
H/A
6.3
N/A
H/A
5.8
7.16
8.5
3.9
3.9
7.24
7.1
N/A
N/A
N/A
6.6
6.42
N/A
N/A
N/A
H/A
N/A
2.3
5.
5.5
N/A
N/A
N/A
1.45
1.8
2.9
N/A
N/A
H/A
Flue Gas
Ib/hr
Rate
N/A
N/A
N/A
N/A
282
276
277
281
282
N/A
281
259
N/A
H/A
N/A
360
H/A
N/A
361
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
H/A
N/A
H/A
N/A
H/A
N/A
N/A
Output
s.
Ib/Tlr
N/A
H/A
N/A
VA
19.5
.VA
17.6
22.1
26.3
12.
12.
20.6
20.2
N/A
N/A
N/A
26.1
N/A
25.4
N/A
N/A
N/A
N/A
N/A
N/A
H/A
H/A
H/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
Reg.
Bed S
N/A
N/A
N/A
N/A
2.56
2.44
0.60
N/A
1.02
0.91
H/A
0.44
3.35
2.64
H/A .
N/A
H/A
2.08
1.77
N/A
N/A
N/A
N/A
N/A
1.36
N/A
N/A
N/A
3.04
1.66
1.01
0.24
N/A
N/A
N/A
ReS.
°2
N/A
H/A
N/A
N/A
N/A
0.8
0.3
0.3
N/A
0.5
0.4
0.7
• 0.4
N/A
N/A
N/A
1.2
0.7
0.7
0.5
H/A
N/A
N/A
5.
N/A
0.5
H/A
N/A
N/A
1.0
1.0
0.5
0.5
N/A
N/A
CO
A
N/A
H/A
N/A
0.01
H/A
H/A
0.27
N/A
0.27
N/A
0.23
0.19
H/A
H/A
N/A
N/A
0.19
0.31
N/A
N/A
N/A
H/A
H/A
N/A
N/A
0.36
N/A
0.04
0.08
1.04
0.36
N/A
H/A
N/A
N/A
HC
Epm
N/A
N/A
H/A
3000
N/A
N/A .
200
N/A
200
N/A
150
100
K/A
N/A
N/A
100
180
380
N/A
N/A
N/A
N/A
N/A
460
300
600
350
120
60
200
100
N/A
N/A
N/A
H/A
Reg.
C°2
y
N/A
N/A
N/A
21
N/A
N/A
25
N/A
25
N/A
25
25
N/A
N/A
N/A
24
N/A
25
N/A
N/A
N/A
H/A
23
25
25
15
15
N/A
12
15
N/A
N/A
N/A
N/A
Valuo is for sum of sulfur in flue gas from FBM and CBC unless
noted "FBM"
-------
TABLE 10
FBMi
Bed
Day
7/15
7/16
7/17
7/17
7/17
7/18
7/1 B
7/19
Time
2000
2200
0000
0200
0400
0600
0800
1000
1200
1400
1600
1800
2000
2200
2400
0200
0400
0600
0800
1000
1200
1400
1600
1800
2000
2200
2400
0200
0400
0600
0800
1000
1200
1400
1600
1800
2000
2203
2400
0200
040U
T. °f
N/A
N/A
H/A
1750
1450
1640
1470
1480
1470
1470
1510
1470
1470
1450
1460
1470
1700
'1670
1510 •
1510
1470
1490
1510
1520
1440
1525
1530
1600
1470
1500
1450
1440
1440
1390
1500
1500
1540
1520
1520
1470
1460
In
N/A
N/A
N/A
N/A
N/A
16
N/A
N/A
N/A
N/A
N/A
N/A
N/A
17.5
N/A
H/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
17
17.5
N/A
N/A
N/A
N/A
N/A
16
N/A
N/A
N/A
r./i\
N/A
N/A
14
N/A
N/A
(CONTINUED)
Coal Air
Rate Rate
lb/hr
H/A
N/A
N/A
N/A
536
536
509
493
N/A
N/A
502
450
N/A
484
465
520
N/A
525
506
446
N/A
567
520
N/A
477
472
473
490
445
3B9
366
387
392
392
368
N/A
450
375
460
447
N/A
Ib/Tir
N/A
N/A
N/A
N/A
N/A
7605
N/A
N/A
N/A
N/A
N/A
N/A
N/A
7600
N/A
N/A
N/A
N/A
N/A
H/A
N/A
N/A
H/A
H/A
7605
7150
N/A
5795
N/A
6045
N/A
5600
N/A
H/A
H/A
H/A
72*5
7295
6945
N/A
N/A
Fly Ash
H/A
N/A
N/A
N/A
H/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
38.8
N/A
26.2
28.4
31.2
N/A
38.2
H/A v/,
"is""
H/A
N/A
46.1
N/A
N/A
N/A
N/A
36. >
34.9
N/A
33...
Pluc Gas
S02 ppa
N/A
N/A
H/A
210
250
N/A
430
6BO
560
N/A
960
710
710
720
N/A
310
N/A
H/A
500
340
5 SO
• 580
300
310
N/A
490
• 490
N/A
. 500
' 200
• 230
210
• 540
• 660
N/A
N/A
480
460
280
2SO
210
Flue Ga
IVhr
N/A
N/A
N/A
N/A
N/A
8008
N/A
N/A
N/A
N/A
B/A
N/A
N/A
7963
N/A
H/A
H/A
H/A
H/A
N/A
N/A
N/A
N/A
N/A
7962
7504
N/fc
6162
N/A
6336
N/A
5900
N/A
N/A
N/A
N/A
7662
7604
1314
N/A
N/A
Bed
a S
N/A
N/A
N/A
N/A
N/A
N/A
N/A
2.66
N/A
N/A
N/A
N/A
N/A
2.22
2.63C
N/A
N/A
N/A
1.12
H/A
1.68
Salt In
2.2
2.51
1.39
N/A
1.9
H/A
N/A
2.20
N/A
N/A
1.82
N/A
Now Coal
N/A
3.92
3.92
N/A
N/A
3.36
Bed
Ca
1
N/A
N/A
N/A
N/A
N/A
N/A
N/A
36. «1
N/A
N/A
N/A
N/A
N/A
N/A
!230)
N/A
N/A
H/A
H/A
H/A
N/A
N/A
11.25
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
25.39
N/A
N/A
N/A
N/A
H/A
N/A
N/A
N/A
Fly Ash
X
S Cl
N/A
N/A
N/A
N/A
H/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
1.56
N/A
2.02
1.85
1.61 15.
N/A
1.61 17,
N/A
N/A
2.34 23.
N/A
N/A
1.97
N/A
N/A
N/A
N/A
1.64
2.74 23
N/A
2.64
°2
[• 3L
N/A
N/A
N/A
1.9
N/A
3.
5.B
1.2
N/A
N/A
N/A
N/A
H/A
H/A
7.2
N/A
H/A
3.2
9.4
5.6
N/A
4.6
6.4
7.4
2 N/A
8.5
.5 B
1.6
4
88 3
2.2
6
6
N/A
N/A
N/A
6.5
6.3
.9 N/A
5.4
4.
m
JL
. N/A
N/A
N/A
0.18
N/A
N/A
0.11
N/A
0.06
0.13
0.36
0.09
0.08
0.19
N/A
0.19
0.18
N/A
0.05
0.09
O.OB
0.09
0.27
0.23
N/A
0.08
0.13
H/A
H/A
0.35
0.10
0.09
0.12
0.09
0.09
0.09
0.23
0.09
0.09
0.19
0.13
Hr
ppro
N/A
H/A
N/A
500
N/A
N/A
200
100
N/A
N/A
130
110
ISO
160
N/A
160
130
900
50
310
100
160
100
55
N/A
50
N/A
N/A
N/A
N/A
10
BO
10
30
20
50
120
150
320
150
C02
JL.
N/A
N/A
N/A
8.
N/A
7.5
6.5
1.5
1.5
H/A
6.5
B.6
8.6
8.
B.I
B
8
1
8.6
8.5
8.5
9.
9.5
9.6
N/A
10.5
11.5
N/A
N/A
N/A
7.8
7.6
B
7
9
9
8.6
9
9.5
9.5
9.5
Steam
IbAr
U/A
U/A
U/A
N/A
H/A
H/A
U/A
4000
3100
3t>50
3450
3500
3150
3200
3400
3400
4200
3800
3000
3400
3250
3500
3500
3500
3300
3300
3400
3400
3150
3000
2950
2950
2950
2750
K(/A
2950
3100
2950
3100
3200
3250
CBCi
°2
£
N/A
6
9.7
b
B
b
3
b
4.6
4
2
2
3
6
10
10
10
10
10
9.B
10
9
10
10
1
10
10
N/A
N/A
N/A
N/A
N/A
4.2
3
4.0
4.4
10.2
N/A
N/A
N/A
N/A
Ash
C
£_
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
H/A
22.2
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
8.5
N/A
2.1
6.5
2.6
N/A
3.0
N/A
N/A
9.9
N/A
N/A
7.9
N/A
N/A
N/A
N/A
7.6
20.4
N/A
4.9
S02
ppa
N/A
N/A
N/A
N/A
N/A
N/A
N/A
9000
1000
N/A
850
650
240
580
320
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
H/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
420
630
630
N/A
N/A
Air
Rate
IbAr
N/A
N/A
N/A
N/A
N/A
1480
N/A
N/A
N/A
N/A
N/A
N/A
H/A
H/A
H/A
H/A
H/A
H/A
H/A
N/A
N/A
N/A
N/A
N/A
N/A
960
N/A
1095
N/A
1295
N/A
N/A
N/A
N/A
N/A
N/A
890
1165
863
N/A
N/A
System
C
Burnup
X
Overall
N/A
N/A
N/A
N/A
N/A
N/A
N/A
H/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
H/A
H/A
H/A
N/A
N/A
99.
N/A
4-3.
99.
99.
N/A
99.
N/A
N/A
99.
N/A
N/A
99.
H/A
H/A
H/A
N/A
99.
95.1
N/A
99.
T. °F
N/A
N/A
N/A
1940
1930
2030
1900
2010
1B80
1970
2000
1980
2000
1800
1.980
W50
2000
.1960
1660
1990
2010
2040
1840
I960
1950
1960
2030
2010
1970
1950
1960
1910
1920
1500
1780
1960
2010
1885
1910
1960
1960
CO
ppro
N/A
N/A
N/A
N/A
N/A
N/A
N/A
1300
22000
N/A
N/A
N/A
H/A
H/A
H/A
H/A
100
100
H/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
H/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
300
300
300
N/A
H/A
C°2
_x_
N/A
N/A
N/A
H/A
N/A
H/A
H/A
13.
12.5
H/A
1.1
14
11.5
H/A
N/A
N/A
N/A
H/A
H/A
H/A
N/A
N/A
N/A
N/A
N/A
N/A
H/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
10.6
13
7
9
N/A
N/A
HC
£E»
N/A
N/A
N/A
N/A
N/A
N/A
N/A
40
100
H/A
40
40
40
40
N/A
40
15
30
H/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
H/A
JO
<0
30
N/A
N/A
•Carbon Free Basis
-------
TABLE 10 (CONTINUED)
Dai
7/15
7/16
7/17
7/18
7/16
7/19
Tint
2000
2200
0000
0200
0400
0600
OSOO
1000
1200
1400
1600
ieoo
2000
2200
2400
0200
0400
0600
0300
1000
1200
1400
1600
1300
2000
2200
2400
0200
0400
0600
0800
1000
1200
1400
1600
ieoo
2000
2200
2400
0200
0400
Syateti
S Input
In Coal.
lb/Hr
N/A
N/A
N/A
N/A
N/A
N/A
16.3
16.3
N/A
N/A
17.9
16.1
N/A
17.1
16.6
18.5
N/A
18.6
18.0
16.1
N/A
18.0
16.4
N/A
16.8
17.0
16.9
17.4
15.9
14.1
13.4
13.8
14.3
N/A
N/A
N/A
20.8
17.4
21.2
20.8
N/A
CBC
Bed
X
_5_
N/A
N/A
N/A
N/A
N/A
N/A
N/A
0.93
N/A
N/A
N/A
N/A
N/A
2.52
N/A
N/A
N/A
N/A
1.19
N/A
1.49
N/A
2.0
1.82
1.51
N/A
1.66
.N/A
N/A
2.35
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
2.97
N/A
3.3
Lijneitone
Lb/Hr
F««d Rats
N/A
N/A
88
08
ea
88
86
N/A
N/A
N/A
64
56
71
71
64
0
0
0
302
N/A
149
0
Salt
0
63
61
61
0
0
24
0
67
67
63
63
N/A
N/A
6*
64
55
152
103
CBC
Flue Gas.
Lb/Mr
N/A
N/A
N/A
N/A
N/A
1520
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
H/A
N/A
N/A
N/A
on
N/A
N/A
N/A
1000
N/A
1135
N/A
1335
N/A
N/A
N/A
N/A
N/A
N/A
930
1205
900
N/A
N/A
S
Rel.
Lb/Mr
N/A
N/A
N/A
N/A
N/A
2.2
N/A
N/A
N/A
N/A
N/A
N/A
N/A
6.2
N/A
N/A
S/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
2.7
4.0
N/A
3.3
N/A
1.39
N/A
1.39
N/A
N/A
H/A
N/A
4.0
4.0
2.24
H/A
N/A
(FBM)
(FBM)
(FBK)
(FBX)
(FBM)
(FBM)
(FBM)
(FBM)
(FBM)
(FBM)
Capture
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
64
N/A
II/A
N/A
N/A
N/A
N/A
N/A
N/A
(FBM)
N/A
N/A
84 (FBM)
76.5 (FBM)
N/A
61.1 (FBM)
N/A
90.1 (FBM)
N/A
90.2 (FBM)
N/A
N/A
N/R
N/A
80.8 (FBM)
77 (FBM)
89.5 (FBM)
N/A
N/A
H/A
N/A
N/A
N/A
N/A
N/A
46
0
N/A
37
40
37
41
33
37
40
41
39.1
39.'
41
35
39
38
34
31.!
43
40
37
38
37
40
37
41
N/A
N/A
29
33
29
32
35
35
Reg.
Air
Ib/Hr.
N/A
N/A
N/A
N/A
N/A
280
N/A
N/A
N/A
N/A
N/A
N/A
:i/A
230
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
l.VA
N/A
295
384
N/A
265
N/A
265
N/A
N/A
N/A
N/A
N/A
N/A
305
305
325
N/A
N/A
Reg.
X C
In Aah
N/A
N/A
N/A
H/A
N/A
N/A
N/A
N/A
H/A
H/A
N/A
N/A
N/A
N/A
N/A
N/A
H/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
16.8
12.5
N/A
N/A
N/A
N/A
N/A
12.3
13.6
N/A
N/A
N/A
N/A
27.7
36.4
N/A
11.7
Peg.
Jemp.
N/A
N/A
H/A
1620
1660
2040
1980
1650
2090
1940
2030
1980
1940
1940
1940
2000
1900
1H90
1900
2000
1960
1970
2050
1960
1930
2050
I860
1900
i860
1980
1880
I860
1900
H/A
H/A
1900
2070
2040
2000
1940
1900
N/A
N/A
N/A
N/A
0.8
0.7'
1.0-
0.85
0.92-
1.45
1.3
1.7
i.e
1.65
1.65
1.45
N/A
N/A
2.1
3.5
3.
N/A
2.6
2.4
2.4
3.
3.8
2.9
N/A
N/A
1.4
1.4
1.8
N/A
N/A
2.6
2.7
2.7
H/A
H/A
4.5
Rate
N/A
N/A
N/A
N/A
300
N/A
N/A
N/A
N/A
N/A
N/A
N/A
310
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
315
412
N/A
289
N/A
289
N/A
N/A
N/A
0
N/A
N/A
J26
324
345
N/A
N/A
Reg.
N/A
N/A
N/A
H/A
N/A
N/A
H/A
N/A
N/A
H/A
H/A
6.2
5.7
5.7
N/A
K/A
H/A
N/A
N/A
N/A
N/A
8.4
8.4
13.6
12.1
9.2
N/A
N/A
4.5
N/A
N/A
H/A
N/A
9.4
9.7
9.6
10.2
N/A
14.5
Reg.
Bed
S
_¥.
N/A
N/A
N/A
N/A
N/A
N/A
N/A
0.58
N/A
N/A
H/A
N/A
N/A
N/A
N/A
H/A
H/A
H/A
1.32
N/A
0.31
N/A
0.27
1.39
0.35
N/A
N/A
N/A
N/A
N/A
N/A
N/A
0.22
N/A
N/A
N/A
1.64
0.56
N/A
N/A
1.43
RAO
Keg •
°2
-*_
N/A
N/A
N/A
1.2
N/A
1.6
N/A
2
0.8
2.5
0.4
0.2
0.9
0.4
0.2
0.6
0.1
0.6
N/A
1.1
0.8
0.4
1.4
0.8
0.3
0.4
N/A
H/A
N/A
N/A
N/A
2
1.2
N/A
N/A
2
2
1.6
1.4
1.6
N/A
Flue
CO
2S_
N/A
N/A
H/A
N/A
0.42
0.87
0.13
N/A
N/A
0.07
N/A
0.-33
0.73
0.52
0.19
0.51
N/A
0.27
N/A
N/A
1.13
N/A
0.52
0.32
1.13
0.48
0.56
0.73
0.52
0.50
N/A
0.53
0.67
N/A
N/A
N/A
0.35
0.56
N/A
N/A
0.42
Gag:
HC
EE»
N/A
N/A
N/A
N/A
350
600
300
N/A
450
80
140
N/A
450
370
no
120
N/A
N/A
N/A
H/A
200
N/A
100
250
650
375
730
1600
150
1400
1200
340
600
N/A
N/A
N/A
100
300
N/A
N/A
1400
CO
UU2
_*_
N/A
N/A
N/A
N/A
11
7.5
10
11
8
13
12.5
11
10.5
12.7
13
11
N/A
N/A
N/A
15
23
N/A
21
16
15
22
23
19.5
18
N/A
N/A
12
14.5
N/A
N/A
14.3
16
17
N/A
N/A
20
Possible Mr Infiltration
-------
TABLE 10 (CONTINUED)
Day
7/19
Tine
0600
0800
1000
1200
1400
1600
1753
FBM
T. °F
1460
1490
14CO
1410
1460
1400
N/A
Bed
Level
In,
N/A
N/A
19.5
N/A
16.5
N/A
15
Coal
Rate
N/A
450
465
354
563
483
483
Air
N/A
6700
N/A
N/A
N/A
N/A
N/A
Fly Ash
Carbon X
N/A
N/A
35.4
32.6
22.2
N/A
N/A
2
ppm
220
380
320
280
N/A
420
N/A
Flue Gas
Ib/hr
N/A
7040
N/A
N/A
N/A
N/A
N/A
Bed
S
%
N/A
N/A
3.58
4.03
2.37
N/A
5.02
Bed
Ca
%
N/A
N/A
N/A
29.4
N/A
N/A
N/A
Fly Ash
X
S Ca
N/A N/A
N/A N/A
N/A N/A
1.62 N/A
1.48 N/A
1.09 10.6
N/A N/A
°2
£_
N/A
5.5
6.
6.
7
8
6
CO
JL
N/A
0.19
0.25
0.23
N/A
0.22
0.16
HC °°2
ppm X
N/A N/A
500 9.7
550 9.6
520 9.6
500 10.5
N/A 10
200 12
Steam
Ib/hr
3250
3300
3200
3050
3150
2650
2900
CBC
o Ash
°2 C X
N/A N/A
4. N/A
N/A 18.6
5.5 13.07
6.6 10.7
7.6 N/A
N/A N/A
2
BE"
N/A
N/A
N/A
N/A
N/A
500
N/A
Systejn
Burnup
Overall
X
N/A
N/A
96.0
97.2
98.1
*
N/A
1920
1975
1900
ieeo
1960
1430
H/A
CO
ppn
N/A
N/A
N/A
N/A
N/A
1600
N/A
N/A
N/A
N/A
N/A
N/A
11.
Il/A
HC
pptti
N/A
N/A
N/A
N/A
N/A
2/0
N/A
*Hlth 12* alh coal
10(CONCLUDED)
Dav.
7/19
Tine
0600
OBOO
1000
1200
1400
1600
1753
CBC
System Bed
S X
Input S
N/A N/A
20. B N/A
21.4 3.33
16.6 2,60
25.7 2.35
22.2 N/A
N/A t!/A
FBM
Limestone
Ib/lir.
Food Rate
112
62
62
61
0
0
0
CBS
Flue Gas
Ibi/hr
N/A
H/A
H/A
N/A
IJ/A
tl/A
;J/A
FBM
S Release
Ib/hr
N/A
2.94 (FBM)
Il/A
H/A
?..2 (PBM)
N/A
N/A
S Capture
t
N/A
66.3 (FBM)
N/A
N/A
91.3 (FBM)
N/A
N/A
Reg.
Coal
UVMr
30.4
31
30
30.5
33
31
N/A
Reg Air Flov
lh*/hr
N/A
• N/A
N/A
N/A
N/A
N/A
N/A
Reg.
riy Ash
N/A
N/A
26.5
25.6
N/A
37.3
N/A
1890
1895
1880
2050
1940
1670
N/A
Keg .
3.6
3.6
N/A
2.5
3.69
1
N/A
Reg.
Bed
5
Reg.
N/A
N/A
0.36 (1100)
0.83 (1300)
N/A
N/A
O.B6
-------
6-56
(18 determinations). The other coal (used later in the
test) was "Rivesville" (see Section 5.6). With the
Powhatan coal, average regenerator fly ash carbon content
was 13.4% (5 determinations). Based on this result a
CBC treatment of the Regenerator output ash may not be
necessary. During the first 38 hours of the test, the
CBC ash collector was faulty. The result of this on the
burnup measurements is that they are questionable.
Average REG bed sulfur content was 1.11% (26
determinations). Average FBM bed sulfur content was
3.05% (33 determinations). The REG SO- output and bed
S contents imply a bed circulation rate of about 670
Ib/hr (about 0.6 FBM bed changes per hour), which is
consistent with the transfer screw system operating rate
plus natural circulation between fluidized beds. In a
boiler system initially designed to incorporate both a
CBC and a regenerator, the REG-FBM slot lengths would be
less than in the present system, and natural circulation
would be enhanced. The present Regenerator is an add-
on device. The need for screw system operation (an
expedient) would be reduced.
Median SO- emission from the FBM was 320 ppm
with the 3.3% S coal, equivalent to a capture of 89%
in the FBM, based on coal input to the system. Since
this capture is less than that experienced in Run 168H,
it appeared possible that a system operating change
had acted to reduce capture. One possibility was
that the REG fly ash (which contains an average
of 4% S and up to 5.4% S and was fed to the CBC),
-------
6-57
also contained a poison which was recaptured in the CBC
under oxidizing conditions. To flush possible poisons
from the system, an unnaturally high limestone feed
rate was imposed on the system (averaging 55 Ib/hr) and
bed material periodically withdrawn and discarded.
To facilitate carbon balance calculations for
various portions of this test, fly ash streams were analyzed
periodically for carbon, calcium, and sulfur. These
results are collected in Table 11. These materials were
not analyzed for arsenic. For coal No. 1 (the 6.1%
ash,3.2% S Powhatan coal) the average S/Ca ratio in the
FBM fly ash was 0.15 (7 determinations). The median Ca
content of these ashes was 10.5% (22 determinations).
For coal No. 2 (the 12.4% ash, 3.85% S Rivesville coal)
the FBM fly ash Ca contents were all below 8.4% in line
with the heavier inert loading, and the average S/Ca
ratio in these ashes was 0.24 (6 determinations). The
fly ash from the CBC tended to have a S/Ca ratio of 0.08
regardless of coal origin (12 determinations).
The dilution of the CaSO. in the fly ash by elutriated
CaO in the CBC is probably due to the relatively high
gas velocities used. Dissociation of the CaSO. content
of the fly ash or bed in the CBC would be undesirable,
from an air pollution standpoint.
At one point in the test, salt addition to the
FBM was performed. Salt was screw fed from a weighing
hopper and mixed with the coal feed. The result was
an immediate 48% reduction in SO_ emissions and a 32%
reduction in fly ash carbon content (see Table 10, entry
"7/17-1400").
-------
6-58
TABLE 11
FLY ASH CHARACTERISTICS - RUN 171H
Vessel
FBM
FBM
FBM
FBM
FBM
FBM
FBM
FBM
FBM
FBM
FBM
FBM
FBM
FBM
FBM
FBM
FBM
FBM
FBM
FBM
FBM
FBM
Time
***
12-1850
12-1905
12-2050
12-2200
12-2300
13-0210
13-0456
13-0835
13-1045
13-1250
13-1800
13-2030
13-2230
14-0712
14-2330
14-2400
15-0845
15-1045
15-1245
16-1700
17-1130
17-1535
17-1830
17-1830
17-2030
17-2330
* ..
Input
%C
49.6
N/A
29.7
N/A
36.5
53.5
N/A
49.9
36.2
27.5
N/A
21.8
27.3
41.9
24.55
24.6
50.4
50.6
46.8
N/A
38.8
26.2
28.4
N/A
3.12
38.2
%Ca
(as rec'd)
15.6
N/A
29.47
N/A
16.2
10.1
N/A
6.9
15.1
10.05
11.4
10.05
10.05
9.68
11.7
N/A
10.6
9.1
9.4
N/A
9.7
16.95
15.5
11.4
10.45
10.85
%S
1.71
N/A
1.97
N/A
2.18
1.64
N/A
1.4
1.37
2.29
N/A
1.23
1.09
1.68
1.0
N/A
N/A
1.76
3.14
N/A
1.56
2.02
1.85
N/A
1.61
1.61
Ve s s e 1
CBC
CBC
CBC
CBC
N/A
CBC
CBC
CBC
CBC
CBC
CBC
CBC
CBC
N/A
CBC
CBC
CBC
CBC
CBC
CBC
CBC
CBC
CBC
CBC
CBC
N/A
**
Output
%C
N/A
N/A
0
1.
N/A
1.
1.
1.
1.8
0.7
N/A
0.51
1.
N/A
N/A
3.9
(2300)
11.95
2.6
3.9
22.2
8.5
2.7
6.5
N/A
2.6
N/A
%C
37.65
N/A
N/A
42.2
N/A
32.2
N/A
33.5
25.75
39.31
20.3
28.8
35.6
N/A
17.3
27.7
21.6
36.9
35.6
20.
22.5
23.5
22.1
20.3
26.5
N/A
%S
(1705)
2.4
N/A
2.41
N/A
2.68
2.63
2.32
1.72
1.8
N/A
(2045)
2.56
N/A
1.82
2.2
1.55
3.43
3.20
2.28
2.30
0.83
N/A
N/A
2.2
N/A
1.24
N/A
-------
6-59
TABLE 11 (continued)
Vessel
FBM
FBM
FBM
FBM
FBM
FBM
FBM
FBM
FBM
FBM
FBM
FBM
FBM
Time
17-2345
18-04'00
18-0515
18-1145
18-2115
18-2300
18-2400
19-0435
19-1100
19-1000
19-1300
19-1200
19-1500
Input %Ca
%C (as rec'd)
N/A
38.9
36.8
40.1
38.6
new
N/A
34.9
33.4
N/A
35.4
N/A
32.6
22.2
N/A
NLA
15.1
w/salt
NLA
6.97
coal
15.6
N/A
14.5
6.01
N/A
6.18
N/A
8.3
%S
N/A
N/A
2.34
1.97
1.84
2.74
N/A
2.69
1.62
N/A
1.48
N/A
1.09
Vessel
CBC
N/A
CBC
CBC
CBC
CBC
N/A
CBC
CBC
CBC
CBC
CBC
CBC
Output
%C
3.
N/A
9.9
7.9
7.6
20.4
N/A
4.9
N/A
18.6
N/A
13.07
10.07
%C
19.8
N/A
29.1
NLA
17.8
9.
N/A
21.31
26.92
N/A
13.05
N/A
12.64
%S
1.42
N/A
2.36
2.72
2.0
1.54
N/A
1.99
1.93
N/A
1.20
N/A
1.55
* Input to CBC ** Output of CBC
*** "12-1850" means July 12, 1972, @ 1850 hours
Note: NLA = Sample expended, no longer available
-------
6-60
Regenerator fly ashes were also analyzed for calcium
(average 23%, 10 determinations and carbon, median 14%
11 determinations). Fly ash calcium values were converted
to CaCO., equivalent output from each reactor (see Table 12) .
There is a slight dependence of FBM fly ash calcium
output on limestone feed rate to the FBM. This may be
attributed to a certain fraction of the makeup limestone
particles being inherently weak. Dead slow on the lime-
stone feeder represented 60 Ib/hr, with the particular
drive gearing in use. Rates below this value reflect
intermittent operation of the feeder. When limestone
feed rate is less than 25 Ib/hr, median calcium fly ash
output is calculated to be 18 Ib/hr CaCO, equivalent.
When feed.rate is above 52 Ib/hr, median calcium fly ash
output is calculated to be 26 Ib/hr CaCO., equivalent.
The CBC calcium output tends to be higher than the FBM,
due to its higher superficial velocity* as set in this test.
The FBM fly ash median S content was about 1.7%.
This ash could not be fed to a CBC operating with an
agglomerating ash bed** at temperatures above 2,000 F
without a substantial excess of lime present, as SO_
emissions up to 2,400 ppm might result. However, it is
possible that the functions of carbon burnup, ash agglomera-
tion and bed regeneration can be recombined in one vessel,
and development of this concept is needed.
The system heat balance data are collected in
Table 13. With the particular configuration of add-on
regenerator in use in this test, radiation losses are high.
* See Table 12 for typical velocity values.
** Reference Goldberger Patent, U.S. #3,171,369.
-------
6-61
. TABLE 12
CALCIUM BALANCE, FBM/CBC/REG TEST NO. 171-H
Day
12
12
12
13
13
13
13
13
15
15
15
17
17
17
18
18
18
19
19
Med-
ian
Aver-
age
Time
1800
2000
2300
0200
0800
1000
2000
2200
0800
1000
1200
1800
2000
2330
0600
2000
2400
1000
1400
CaCO, Fed Output in FBM Fly
lb/nr CaCO, equiv. , Ib/hr
64
80
84
16
1
68
0
0
53
98
60
63
61
0
12+12
69
55
62
0
60
44
41
75
33
26
13
23
12
14
25
24
27
27
15
18
24
19
60
16
22
24
27
Ash Output in CBC Fly Ash
CaCO-, equiv. , Ib/hr
55
70
46
46
48
26
41
64
28
75
73
30
32
37
35
19
92
67
46
48
Typical superficial vel., ft/sec 9 15
-------
TABLE 13
Hsat Balance Data
Run 171-H
DAY
TIME
FBM Bed temp. ,°F
SO, capture %
Carbon Burnup %
Reg S02 %
Total preheated air, Ib/hr
KBtu/hr
CBC Air, Ib/hr
FBM coal heating value, KBtu/hr
Total input, KBtu/hr
Steam Absorption, KBtu/hr
Circ. HjO Abs. KBtu/hr
Total (RjO + St. )
% of coal value
Flue gas loss, KBtu/hr
Slowdown allowance
Hydrocarbon loss
Limestone calcining
CO loss, KBtu/hr
CBC carbon loss
% of coal value
(tot fly ash loss
FBM dust coll. carbon loss
Lift air loss
Radiation and unaccounted
Total losses
12
1800
1525
86
99+
6.3
5545
. 411
1130
5577
6345
3045
1437
4482
80
707
15
60
50
25
48*
0.8
16
71
16
494
1452
12
2000
1480
88
99+
5.8
5300
367
1120
5315
5987'
3147
14K*
4563
85
664
16
60
62
45
45*
0.8
16
41
19
151
1057
12
2400
1500
94
99+
7.1
5340
360
985
6246
6970
3147
1422
456)
73
652
16
60
66
47
54*
0.8
14
41
N/A
1147
2031
13 13
0200 0845
1540 1540
93 78
99+ 99+
8.5 3.9
5420 5880
376 452
1065 925
.6299 5918
7056 6724
3248 3248
1451 1422
469) 4670
74 78
662 672
17 17
60 100
H N/A
46 83
54* 51*
0.8 0.8
14 13
74 51
VA N/A
1054 615
1981 1602
13
1000
1480
79
99+
7.1
5880
452
925
5367
6120
3147
1422*
4569
85
672*
16
112
53
107
45*
0.8
13
30
N/A
104
1099
13
2000
1460
91
99+ ,
6.4
6675
463
1195
6220
7155
3350
1440
4790
76 ..
852
17
122
N/A
118
54*
0.7
12
15
27
685
1902
13
2200
1480.
86
99 +
6.4
7295
523
1265
6758
7780
3350
1501
4851
71
866
17
122
N/A
118
58*
0.8
16
22
21
1166
2406
!15
0800
1500
88
98.7
1.5
ilOOO
574
1300
ii561
. 7579
J198
1517
4715
71
882
17
64
42
61
92
1.3
15
64
21
1074
2290
15
1000
1460
88
99 +
1.8
8000
574
1300
7349
8331
3451
1517
4X8
«7
889
17
«4
77
41
62*
0.8
20
73
21
1582
2789
15
1200
1500
88
99 +
2.9
8000
574
1300
7873
l»2<
3451
1517*
4)41
43
191
17
213t
47
122
67*
O.I
21
73
21
1959
3389
17
1800
1520
84
99 +
N/A
N/A
N/A
N/A
N/A
•N/A
3552
''•/A
N/A
»/A
N/A
18 <*
i
N/A Ki
N/A
N/A
58*
K/A
N/A
28
N/A
N/A
N/A
t Coal Feeder
50% plugged
-------
6-63
6.4.3 Arsenic Analyses
In test 171H, an apparent loss in sorbent activity
was observed. When fly ash from a regenerator is fed
to a Carbon Burnup Cell, along with the primary boiler
cell fly ash, an interference with SO_ capture in the
primary boiler cell occurs. This interference may be
due to adding the REG to the system. We hypothesize the
REG fly ash may contain a poison. In a test to be described
in the next section of this report, this hypothesis was
tested by discarding the REG fly ash and using an unpoisoned
bed. The result was gratifying, up to 97% S0_ capture.
A possible poison is arsenic, which is present
in coal to the extent of 0.01% or less (mostly as
arsenopyrites , FeAsS) and would be captured in the FBM
as calcium arsenate. The REG flue gas when cobled would
then contain a smoke of As_O,, free or absorbed on carbon.
A series of arsenic analyses was therefore performed.
See Appendix Q. Condensed results are:
Coal (Powhatan) 7
Limestone 0
FBM bed : (average of 12 8.4
determinations increasing
with time)
REG output fly ash 10.9
(median of 7 determinations)
REG bed (median of 13 7.6
determinations )
CBC bed (2 determinations) 4.4
The colorimetric analytical procedure used is given in
Reference 6. Good results were obtained with all materials
except coal, where excessive foaming usually occurred.
A reasonable correlation existed between REG bed As and S
contents .
-------
6-64
The results are consistent with a bed poisoning
mechanism in which As in coal is captured by lime in the
FBM as calcium arsenate which hypothetically interferes
with SO_ capture (as a shell?); desorbs in the regenerator
under high temperature - low O. conditions and is carried
in REG fly ash as arsenous oxide (as a smoke or absorbed
on carbon) and is recaptured in the CBC under oxidizing
conditions. Apparently, once arsenated, the bed is
"regenerated" (dearsenated) only with difficulty*. It
is significant that the FBM bed As content increases with
time. The definitive experiment which needs to be per-
formed is to perturb the capture of SO2 by lime in a
batch experiment using the existing equipment for aqueous
additive solutions. Sodium arsenate could be added,
to see if limestone activity is reduced. The effect of
salt addition on As removal from the beds should also
be researched.
6.5 FBM/REG Test (Test No. 172H)
Following completion of the long duration FBM/CBC
REG Test (No. 171-H), it was desired to test system
operation with the REG fly ash, together with whatever
poisons it contained, discarded. An 8-hour test was
therefore performed, in which the FBM and REG only were
used. A fresh -8+16 mesh limestone bed was used. Fly
ash was discarded. Four FBM fly ash samples analyzed
25.6, 23.2, 15.9 and 23.4% carbon. Regenerator flue
gas SO_ content was 10% volume. FBM flue gas SO- varied
between 70 and 110 ppm. FBM CO_, 0_, NO, CO and tempera-
ture ran in the usual range. These results are consistent
It is reasonable that at a given temperature, "DBA"
(As-O- mobility) would proceed less rapidly than
2.DES" (SO™ mobility, see page 6-15) from molecular
diffusivity and molecular weight considerations.
-------
6-65
with facile capture of S0_ by regenerated lime in the
absence of a bed poison. The 70 ppm S0_ value corresponds
to 97% capture.
6.6 Bed Particle Size Determinations
Bed material samples were withdrawn at regular
intervals during the FBC and FBM/CBC tests for calcium
and sulfur analyses as well as sieve analysis. Carbon
analysis showed zero carbon in the bed, as would be expected.
The six sieve fractions from the Run C-321 samples are
tabulated in Appendix Table G-l and shown graphically
in Figure G-l. No significant size change is indicated
by the raw data or the probability plot. When treated
statistically, the weight mean particle size defined by
d = I d.x.
(where x. = weight fraction of total in i sieve fraction.)
shows a small size reduction as run duration increases.
(See Appendix Figure G-2). On a trial basis, the data
were also treated according to
d' =
but the data scatter is greater by this method due to the
anomalous -25 mesh fraction at 3.7 hours sampling time.
A reliable determination of the effect of salt on attri-
tion rate is not possible from this test.
The Run C-322 samples were analyzed similarly (see
Appendix Table G-2 and Figures G-3 and G-4). At the
higher combustion intensity and coal and air rates an
apparent attrition rate occurs, where d is the time rate
of diameter decrease:
-------
6-66
Test
No.
321
322
Air
Rate
Ib/hr
610
670
Coal
Rate
Ibs/hr
41
64
Bed
Depth
Inches
16
22
d,
|jm/hr
2.2
9
Fly Ash
Rate
Ib/hr
2.9*
25.
* exclusive of salt adding
Characteristics of the five sieve fractions for the
Run C-323 samples are tabulated in Appendix Table G-3.
There is no significant size change between the 2.7,
3.7 and 4.7 hour samples. "5.1 hr." is an aged sample.
The 1.7 hr. sample has a much smaller mean diameter,
reason not known.
The Run C-324 samples were analyzed similarly (see
Appendix Table G-4 and Figure G-5). In the 2.1, 3.1
and 4.1 hour samples, particle mean diameter increases
at the rate of 28 pm/hr, presumably due to coal ash
absorption. These are periods of low fly ash generation.
Following salt addition, the particle mean diameter
decreases at the rate of 6 ym/hr, and the average fly
ash generation rate is higher. The 6.96 hour diameter
is anomalous.
The FBM Run B-18 samples were analyzed similarly
(see Appendix Table G-5). Particle mean diameter decreases
at an apparent average rate of 14 pm/hr. The data scatter
is unsatisfactory in the 9.4 and 10 hour samples.
Selected bed particle size data from Run 168H are
collected in Appendix G, Table G-6, and Figures G-6
and G-7. Since these beds are all at least 36 hours
old*, mean particle sizes are considerably less than
the batch FBC beds (C series) (nonregeneration operation)
and FBM B-18 beds described earlier. Elutriation rate
i.e., the test has been in progress for at least
this time before the sample in question.
-------
6-67
of fines, depends on air rate and particle density (sulfur
and contaminants content) among other factors. Mean
particle size depends upon makeup limestone rate (-8+20
mesh) and elutriation rate, as well as rate of fines
production which may be related to severity of regeneration
as well as coal particle size.
Selected Run 171H beds were particle size analyzed
and the results are listed in Appendix G, Table G-7.
The mean particle size fell during the first six days.
The diameter percentage loss average per day was 1.7
over the first five days. During the 7th and 8th days
Rivesville coal was fed; it contains 1/4" top size rocks
and these affect the bed size consist.
-------
7-1
7. PRELIMINARY FLUIDIZED-BED BOILER .DESIGNS, HEAT
BALANCES, AND COST ESTIMATES.
7.1 Preliminary 30 Megawatt Boiler Designs, Heat
Balances and Cost Estimates
The following design outlines are based upon (1)
the incremental boiler cost, including dust collector with
S0_ abatement system and (2) the overall power plant cost,
for a multiple module fluidized-bed boiler installation
vs. a conventional pulverized coal installation.
7.1.1 30 Megawatt Packaged Boiler Concepts (Figure
26, 27 and 28)
The 30 megawatt designs, for 300,000 Ib .steam/hr,
1,270 psig, 925°F have been developed in three successive
stages, based essentially on proven pilot boiler (FBM)
operation.
(a) RV-I, Figure 26; A non-regenerative limestone
(once thru) unit with four primary boiler cells and a
Carbon Burnup Cell, utilizing primary superheater elements
in the boiler section with secondary superheater elements
in the carbon burnup section. Coal and limestone feed
are mixed via overhead pressurized "run-around" mass flow
conveyor to metering screws which, in turn, deliver the
coal/limestone mixture to vertical feeders extending through
the boiler convection section to inbed split 180 coal
feed sections.
(b) RV-II, Figure 27; A 3-combustion zone unit
(boiler, limestone regenerator, Carbon Burnup Cell) with
primary superheater split between freeboard and inbed
immersed sections* of primary cells. An attemperator
* Freeboard is also commonly termed "slop zone" in
fluidized-bed boilers.
-------
N
v
FIGURE 26.
COAL-FIRED FLUIDIZED-BED UTILITY BOILER, FACTORY ASSEMBLED, 300,000 LB/HR, 1270 PSIG, 925°FTT, RV-I,
-------
LOW SO* GAS TO PREHEATER
MD PREOPtTATOR
CONCENTRATED S02 GAS TO RECOVERY
PLANT OR SCRUBBER
BOILER OUST COLLECTOR
NOTE: For schematic,
see Figure 30.
aULPHATED KB TO 'ROT
—SECONDARY SUPERHCATf* (SSH)
CARBON BUMNUP OCU. (CBC)
YASH FfED
BED REOtNDMTHM CELL (REG)
FORCED DRAFT
AIR HEADER
FIGURE 27. 300,000 LB/HR PACKAGED FLUIDIZED-BED BOILER (1270 PSIG, 925°F) FOR HIGH-SULFUR COAL, RV-II.
-------
7-4
is provided between the slop zone and inbed portion of
the primary superheater for superheater control. Final
superheat is provided within the carbon burnup and regenera-
tor sections. Economizer surface is provided in a separate
package installed over the boiler drum. Express gas flow
is provided for even above bed draft and for optimum jetted
bed particle return.
Coal feed is of the through-grid mushroom distrib-
utor type. Coal is supplied by a single-point-supplied
run-around pressurized mass flow conveyor, from which
two (one on each side) metering screws receive and deliver
a controlled quantity of coal to each mushroom feeder,
as shown in the detail of Figure 27. Limestone feed
may be via separate injection point, as shown, or may be
mixed in with the coal entering the mass flow conveyor.
Air distribution is by means of 15° downthrow nozzle
buttons, as used in the experimental equipment, which
preclude sifting of bed material back into the plenum
upon shutdown.
The limestone and flue gas cycle is described
in the schematic detail of Figure 27, as demonstrated
in the FBM pilot boiler. Low SO. flue gas from the
1500-1600 F boiler beds pass through the slop zone,
convection banks and economizer section into the rein-
jection collector and thence to a regenerative air pre-
heater and fly ash precipitator. Flue gas from the Carbon-
Burnup Cell passes through the discard fly ash collector
to the same air preheater and precipitator. Concentrated
SO_ flue gas, in typically about 4% of total flue gas
stream, is directed, via a dust collector, to the sulfuric
acid recovery plant or to a small lime scrubber.
-------
7-5
The low temperature boiler beds (absorption beds)
pick up sulfur primarily in the form of CaSO. in the
calcined lime particles. Boiler (absorption bed) material
is continuously removed from the front of the unit and
transported to the regenerator section operating at
1900-2000°F and about 0.5% 0,,. Regenerated limestone (as
CaO) with some residual sulfur content passes by gravity
back into the low temperature boiler bed via slots in
the boiler/regenerator tube wall, as demonstrated in
pilot FBM boiler. The Carbon-Burnup Cell, operating at
2000°F, 3% O_, "floats" on the regenerator with minimal
amount of bed material exchange (as necessary to equalize
level).
Fuel supply to the regenerator is by separately
controlled coal feed supply. Fuel for the Carbon-Burn-
up-Cell is the fly ash from the reinjection collector.
Lightoff is by single burner, "burrowing" into the
bed of any coal containing combustion zone, followed by
automatic propagation as demonstrated in the FBM pilot
boiler. Lightoff for this concept will be in the regener-
ator section, since the regenerator communicates with all
four boiler cells and the Carbon-Burnup Cell, thus ob-
viating the need for more than one burner.
A predicted heat balance for this concept is
presented by Table 14.
(d) RV-III, Figure 28: A modification of the
RV-II design with the following modifications:
(1) Superheater tubes are in the boiler
section, consisting of primary convection surface
followed, via an attemperator, by the secondary
superheater immersed in the fluid bed;
-------
7-6
TABLE 14
RV-II HEAT BALANCE AND SURFACE SUMMARY
Circuit
Secondary SH
Primary SH
Inbed Eva p.
Convection
Evap.
Economizer
Duty Temperature
MBtu/hr °F
22 2000
792
61 1550.
581
162 1550
581
37 1550
581
48 (Over 1108
boiler) 525
(Over 2000
CBC/REG ) 525
<--> 2000
-* 925
<--» 1550
-> 792
+-+1550
H- 581
H. 1108
<-->. 581
-* 730
^ 385
•+ 730
H- 385
Surface
425
2650
3570
4250
5850
1720
A 78.5% effective Ljungstrom preheater lowers 341,000 Ib/hr
of boiler and CBC flue gas from 730°F to 275°F to raise
328,000 Ib/hr of air from 70°F to 470°F.
-------
U» K>t OAS TO >«ti«AT»
AND NKQPITATQH
FIGURE 28. 300,000 LB/HR PACKAGED FLUIDIZED-BED BOILER, 1270 PSIG, 925°F
FO3 HI3H-SULFUR COAL, RV-III.
-------
7-8
(2) The closely spaced boiler convection
section is eliminated (drum ligament problems);
(3) Combustion zone heat pickup is augmented
by two additional open spaced tube rows which do
not interfere with maintenance accessibility,
plus additional superheater surface;
(4) Extended economizer banks provide for
cooling REG and CBC off gases;
(5) Elimination of expensive alloys in
secondary superheater.
Predicted heat balance for the RV-III unit is
outlined in Table 15.
7.1.2 30 Megawatt Boiler Cost Estimates, Fluidized-
Bed (RV-III vs. Pulverized Coal)
The incremental cost* of the packaged RV-III
fluidized-bed boiler is summari^ed by Table 16. Item
breakdowns are presented by Appendix L.
The incremental cost of a field erected pulverized
coal unit is estimated, per Appendix M, at $2,538,000
plus $1,000,000 for a wet limestone S0_ abatement system.
7.1.3 Overall 30 Megawatt Power Plant Cost Estimates,
Fluid Bed vs. Pulverized Coal
A comparison of the estimated overall 30 Megawatt
Power Plant cost, is provided in Table 17.
7.2 Preliminary 300 Megawatt Fluidized Boiler Concept,
Heat Balances and Cost Estimates
The unit presented herein as Figure 29 is an 8-
boiler module unit with integral limestone regeneration
and carbon burnup section. Capacity is 1,900,000 Ibs/hr G
The incremental cost excludes items common to any
boiler system of equal capacity„ Feed pumps are
an example.
-------
7-9
TABLE 15
RV-III HEAT BALANCE AND SURFACE SUMMARY
Section
Boiler
Secondary
Superheater
Primary
Superheater
Heat Traps
"REG "
Economizer
CBC
Economizer
Boiler
Economizer
Ljungstrom
Duty Temperature
MBtu °F
190 1550 «-+
587
57 1550 «--»
658 -*
35 1550 ->
694 +
TOTAL STEAM GENERATION
2 2000 H-
393 +-
12 2000 ->
421
41 1177 -»-
525 «-
43 715 -+
604 *•
Diag
1550
587
1550
925
1177
587
730
385
715
393
715
429
2502
70
Surface
Sq . Ft .
6,550
2,070
8,300
16,900 ft
330
2,320
16,700
66,100 ft
Actual_total economizer duty is 46 MBtu versus 55 MBtu as listed,
The 9 MBtu difference is a "safety factor" which may be applied
to manufacturer's margin and unaccounted. Theoretical surface
is on the order of 16,900 ft2 vs. the 19,300 ftz surface as
tabulated.
Allowing for air leakage in the Ljungstrom, corrected exit gas
temperature becomes 220 F.
-------
7-10
CONCfNTMTID SO, «« W
HKOVtUT M.AH
C«C COLLCCTOH
FIGURE 29. 300-JW UTILITY BOILER CONCEPT, 1,900,000 LB/HR
2400 PSIG, 10000F, REHEAT 1,600,000 LB/HR, 650 PSIG, 1000°F.
-------
7-11.
TABLE 16
INCREMENTAL COST, 300,000 LB/HR
FLUIDIZED-BED BOILER INSTALLATION
COST WEIGHT
BASIC BOILER (Dollars) (Tons)
1. Main boiler package (including REG,
CBC furnaces and secondary super-
heater, with insulation, casing,
and grid 310,830 165
2. Primary superheater 44,000 43
3. Economizer (all sections) 109,950 110
4. Air plenum, casing enclosure,
ducts, etc. 16,930 24
5. Structural supports, platforms, etc. 13,930 50
6. Boiler trim 46,703 3
7. F.D. fan and 700 HP motor drive 21,200 6
Basic boiler subtotal 563,000 401
Erection 170,000
TOTAL BASIC BOILER 733,000
SUPPLEMENTAL BOILER PLANT ITEMS, INSTALLED
8. Dust collector and precipitator 201,000
9. Bed moving system 4,000
10. Coal and limestone supply (from bunkers) 65,700
11. Lightoff system 3,000
12. Ash moving 4,000
13. Controls and instrumentation 70,000
-------
7-12
TABLE 16
(Continued)
14. Air preheater 139,000
15. Incremental misc. boiler connections,
piping 20,000
16. Incremental boiler connections,
electrical 10,000
Auxiliary boiler plant subtotal 517,000
Total Boiler Plant
(excluding building) $1,250,000
Boiler Plant Building 105,000
TOTAL BOILER PLANT $1,355,000
Note: This figure excludes constant cost items; e.g., the
electric generation plant, common-to-any system
auxiliary equipment, land acquisition, etc.
-------
7-13
TABLE 17
30 MEGAWATT* POWER PLANT COST ESTIMATES
FPC
cate-
gory
310
311
312
314
315
316
Items
Land and land rights
Structures and improvements
Boiler Plant equipment:
120 Boiler and accessories
121 Draft equipment
122 Feedwater equipment
123 Fuel handling and storage
124 Fuel burning equipment
125 Ash handling equipment
126 Water supply and treating
128 Boiler Instruments and Cont
129 Boiler plant piping
Subtotal of 312:
(Boiler Plant Equipment)
Turbine Generator
Accessories and electrical
equipment
Miscellaneous plant equipment
External SO- Scrubbing***
Other expenses
TOTAL PLANT COST
Cost per KW Plant Capacity
Costs in
Packaged
F.B. with
Integral
Abatement
50
850
1,044
30
220
300
70
120
100
rols 100
300
2,284
2,500
700
100
350
800
7,634
$ 254
Thousand Dollars
Field Erected
P.C. with
SO- External 90-
Abatement
70
1100
) 2538**
)
220
300
Included above**
120
100
Included above**
300
3,578
2,500
700
100
1,000
800
9,848
$ 328
* Net plant capacity after subtracting auxiliary loads.
**FPC items 124 and 128 are included in total items 120
and 121 figures.
***Recent quote from Chemical Construction Co.
-------
7-14
AIR
•H4
COMB.,. 19
PLENUM
r T^T
22
BED CYCLE
BOILER LIMESTONE *ED (I) (CoO) ABSORBS SULFUR AT LOW
(I50O-I60O*F) TEMPERATURE UNDER OXIDIZING CONDITIONS.
ACQUIRING A SULFATE (CoSCU) SHELL. PARTIALLY SULFATED
BED IS MOVED FROM AREA (2) TO REGENERATOR BED (3) AT
HIGH (200O°F) TEMPERATURE, AND LOW OXYGEN CONCENTRATION
WHERE CoSO4 REVERTS TO CoO. REGENERATED BED RETURNS
TO BOILER BED (I) VIA SLOTS (4). MAKEUP LIMESTONE IS
FED AT (5). CBC BED (6) OPERATES INDEPENDENTLY OF OTHER
BEDS, EXCEPT FOR RESTRICTED OPENING (7) FOR BED LEVEL
EQUALIZATION.
GAS AND FLY ASH CYCLE
BOILER FLUE GAS (8), PASSES TO MAIN COLLECTOR (9), MIXES
WITH CBC FLUE GAS (16), THEN TO PREHEATER AND PRECIPITATOR
AT (10). CARBON BEARING FLY ASH IS DISCHARGED AT (II) FOR
TRANSPORT TO CBC (6). REGENERATOR GAS (12) PASSES TO REG
COLLECTOR (13) AND TO SOz RECOVERY (A SCRUBBER) AT (14).
FLY ASH IS DISCHARGED AT (15) FOR INJECTION TO CBC ALONG
WITH MAIN FLY ASH STREAM (II).
CARBON BURNUP CELL GAS (16) PASSES TO CBC COLLECTOR (17)
TO MAIN BOILER FLUE GAS DISCHARGE (10). FLY ASH (18) IS
REMOVED TO STORAGE SILO.
PREHEATED COMBUSTION AIR DUCT (19) FEEDS PLENUM SECTIONS
AT (20), (21) AND (22).
FIGURE 30. GAS, FLY ASH, AND BED REGENERATION SCHEMATIC.
-------
7-15
1,000°F followed by 1,600,000 Ib/hr of reheat from 650°
to 1,000°F. The boiler is approximately 32 ft. wide by
80 ft. long and 30 ft. high (top of economizer).
The gas, fly ash and bed regeneration cycle is
schematically indicated and described by Figure 30.
Coal is supplied to two pressurized Redler dis-
tributors (see Figure 29) which delivers the coal to
four metering screw arrangements which are periodically
supplied and discharged to mushroom feeders (see detail
in Figure 27. Limestone feed may be via separate inputs
into each cell or may be mixed with coal before the
Redler distributor. The Redler distributor ("Runaround
Redler Conveyor") is designed to move about double the
amount of coal required in order to assure positive uniform
feed to the metering screws.
Fly ash from the boiler and regenerator dust collectors
is pneumatically introduced to the CBC section through
the rear wall of the boiler. Experience has indicated
that fly ash feed ports may not be located less than 2 ft.
apart for uniform distribution. Fly ash from the CBC
collector is discharged to silo*
The intent is to discharge the sulfated bed material
from the front of each boiler cell and pneumatically
transport the same to the regenerator in similar manner
as practiced on the pilot boiler (FBM). The regenerated
limestone gravitates back to the 8 boiler beds. The
CBC merely "floats" on the system; interconnection be-
tween CBC and REG is solely for bed level equalization
and lightoff purposes.
-------
7-16
Evaporation is handled by the natural circulation
vertical wall tubes and inclined boiler tubes connecting
to the steam drum. At the 2,400 psig conditions, with
the superheat and reheat indicated, evaporation (steam
generation) is a minor part of the total heat absorption.
Superheat is handled with primary and secondary
sections with attemperation between the two sections.
Reheat is shown as a single imbedded section; detailed
design might indicate a need for a split reheater with
attemperature.
Economizer surface is distributed above the steam
drum and lowers the gas temperature entering the air
heater into the 800 F range.
The heat balance and surface requirements are sum-
marized by Table 18.
Cost estimates for the unit are the subject of
Table 19 as broken down in Appendix N. These are the
incremental costs for the boiler, controls, coal supply
and boiler auxiliaries, only, i.e., the estimated cost
for installation in an existing plant. The overall power
plant cost estimate is the subject of Table 20. The
incremental boiler installation (based on 1971 costs)
is estimated at $24 per KW capacity with the overall
plant estimated at $124 per KW capacity.
Based upon an annual consumption of 800,000 tons
of 4.5% sulfur coal, approximately 64,000 tons of concen-
trated SO- should be delivered to the acid plant for a
by-product of 106,000 tons of 93% H SO..
The concept as shown is based upon an extension
of the 30 Megawatt concept described, heat balanced and
estimated in Table 14, which, in turn, was based upon
-------
7-17
TABLE 18
300 MEGAWATT FLUIDIZED BED BOILER SINGLE BED LEVEL,
NATURAL CIRCULATION HEAT BALANCE AND .SURFACE REQUIREMENTS
Circuit
INBED AND SLOP ZONE
Main Boiler Evap
Reg Evap
CBC Evap.
Primary Superheater
Secondary Superheater
Reheat
CONVECTION
Evaporator
Overbad Primary SH
Main Economizer
Reg Economizer
CBC Economizer
Preheater (Ljungstrom)
Duty,
MBtu
532
29
136
114
538
315
40
195
273
12
75
343
Heat
<
1550
674
2000
674
2000
674
1550
694
1550
760
1550
650
1550
674
1550
760
1178
620
2000
523
2000
517
794
693
Transfer,
'F
«--»• 1550
«-+ 674
I--*- 2000
+•+ 674
t-v 2000
«-+ 674
++ 1500
674
«--*• 1550
+ 1000
++ 1550
-> 1000
•» 1230
«--> 674
-*• 1242
*~* 694
AV
+ 794
--•* 523
AV
-> 794
•*- 517
AV
H. 794
•••-+ 480
AV
->• 234
•e 80
Surface
Sq. Ft.
16,000
1,200
1,410
3,480
21,400
11,650
9,365
14,750
91,490
2,430
14,150
454,400
-------
7-18
TABLE 19
INCREMENTAL COST OF 1,900,000 LB/HR, 1000°F SINGLE LEVEL,
NATURAL CIRCULATION FLUIDIZED-BED BOILER INSTALLATION
WITH 1,600,000 LB/HR JEHEAT FROM 650°F TO 1,000:°F . :
Installed Cost
1. Boiler, complete, including
superheater, atemperator,
reheater, economizer, grid, nozzles,
insulation, trim, and boiler supports $4,385,000
2. Instruments and controls 400,000
3. F.D. fan and drives 86,000
4. Dust collector and precipitator 825,000
5. Bed transport 40,000
6. Coal supply 274,000
7. Lightoff system 12,000
8. Fly ash transport 38,000
9. Air preheater 800,000
10. Duct work 150,000
11. Miscellaneous direct boiler piping
and electrical 200,000
TOTAL $7,210,000
Note: This Table excludes constant cost items, e.g., the
electric generation plant, common-to-any system
auxiliary equipment, land acquisition, etc.
-------
7-19
TABLE 20
OVERALL 300 MW FLUIDIZED BED POWER PLANT CAPITAL
ESTIMATE
COST
FPC
CATEGORY
310
311
312
314
315
316
120
121
122
123
124
125
126
128
129
ITEM
Land and land rights
Structures and improvements
Boiler plant equipment:
Boiler and accessories $6,288
Draft equipment 236
Feedwater equipment 1,050
Fuel handling and storage 2,350
Fuel burning equipment 286
Ash handling equipment 350
Water supply and treatment 13f>
Boiler instrunentation/control 400
Boiler plant piping 1,800
TCTAL ITEM 312
Turbine generator
Accessary electrical equipment
Miscellaneous plant equipment
Other expenses
TOTAL PLANT COST
COST IN
THOUSAND
DOLLARS
290
4,000
12,896
10,750
2,430
530
6,150
37,046
-------
7-20
pilot boiler (FBM) experience. As such, it is two
"generations" away from an actual operating unit. Exper-
ience with a smaller commercial prototype will undoubtedly
change many of the concepts (Figure 29) and design features.
-------
8-1
8. SULFUR RECOVERY
Many processes have been suggested for removal or
recovery of sulfur oxides from gas streams. For example,
about 40 recovery processes were reviewed recently by
Arthur G. McKee and Co. (Reference 8). Considering
economics and the current state of technology, the use
of limestone wet scrubbing appears to offer one of the
most dependable emission control systems at this point.
The concept is under continuing study and earlier work
has been comprehensively reported by Tennessee Valley
Authority (Reference 9). The process is thus suitable
either for consideration for installation, or as a base
case for comparison with other methods, to be applied
to the So_ Acceptor Process Regenerator output gas stream.
TVA has reported (Reference 10) investment estimates
for two sizes of generating plants (200 and 1000 MW), for
both limestone injection in the boiler followed by wet
scrubbing, and for limestone addition at the scrubber.
Annual operating cost estimates are also given. These
data make possible a rational application of the TVA
figures to other sizes. The present study used the TVA
relations for the most part. For the fluidized-bed cases,
a straight prorating is not suitable, for while the SO_
load is nearly as large (90%) as the SO_ load in the
corresponding pulverized fuel case, the volume of gas
to be treated is relatively very small (4%). Therefore,
the investment items were classified as to whether they
would be primarily influenced by SO_ load or by gas load,
and special consideration was given to those items governed
by gas load.
The chemicals and utility costs were recalculated
for the specific conditions of the cases in the present
study. Investment costs were escalated to 1971 values
-------
8-2
using the Machinery and Equipment Index of the Bureau
of Labor Statistics.
The estimates of capital and operating costs developed
for the limestone treating processes are summarized in
Table 21. The cases considered are: Case 1, Fluid bed
boiler, 300,000 Ib/hr steam, limestone addition at
scrubber. Regenerator gas at 450 F and 5.1% SO_ enters
a gas cooler, then enters the scrubber at 175 F. Counter-
current circulating oil is heated from 130 to 420°F.
The limestone supply is prepared and mixed with water,
and then enters the delay and mixing tank. Slurry con-
tacts the gas and is sulfated, and returns to the surge
tank. Some of the slurry is withdrawn from the surge
tank and discarded. Scrubbed gas at 85°F is then re-
heated by the circulating oil to 390°F, and blown by
the I.D. fan to the stack or to the boiler plenum.*
Case 2, same size but pulverized fuel boiler, with
limestone addition to the boiler. Case 3, fluidized bed
boiler as in case 1 but 1,900,000 Ib/hr steam. Case 4,
same size but pulverized fired boiler.
The basis used for amounts.of limestone to be employed
were those selected by TVA for the processes. It will be
noted that the amounts were carefully reviewed by TVA
on the basis of experimental evidence. Particularly
in the scrubber addition system, considerable variation
was found in performance with the several sources of
calcium carbonate. Accordingly, it is to be expected
that the limestone sources proposed for commercial use
will be evaluated for satisfactory performance. With
*The reheating is probably not needed for the fluid-bed
boiler since the wet gas could be combined with the
bulk of gas, still hot, before the stack.
-------
8-3
TABLE 21
CAPITAL & OPERATING COSTS
LIMESTONE TREATING PROCESSES
75% CAPACITY FACTOR
Case
Type of Plant
Addition and
Size
Ib/hr, Steam
T Coal, Annual
Est. Capital Cost
Before Credit, $
After Precipita-
tor Credit, $
Annual Operating
Costs Before
Credits $/yr
$/T Coal
After Precipita-
tor Credits ,
$/yr
$/T Coal
1
Fluid-Bed
Scrubber
300,000
88,200
677,000*
677,000
321,500
3.65
321,500
3.65
2
Pulv. Fuel
Boiler
300,000
88,200
871,000
753,000
376,800
4.28
346,000
3.92
3
Fluid-Bed
Scrubber
1,900,000
597,000
2,246,000*
2,246,000
931,500
1.56
931,500
1.56
4
Pulv. Fuel
Boiler
1,900,000
597,000
3,112,000
2,615,000
1,201,800
2.09
1,068,800
1.79
* Includes reheaters which are probably not necessary.
-------
8-4
this precaution, it is anticipated that adequate sulfur
dioxide removal will be obtained. The scrubber-addition
system is used with the fluidized-bed boiler and the
small volume of treated gas can be recycled to the boiler
plenum for further exposure to the limestone bed at
little cost, instead of being vented directly. Flue
gas recycle has been shown to have the additional benefit
of reducing NO emission from the boiler. In cases 1
and 3, with wet scrubbing - limestone addition to the
scrubber, used boiler lime may optionally be added to
the delay tank, for additional SO absorption efficiency.
The limestone method of S0_ recovery secures only
part of the fluidized-bed benefit, namely, the small
volume of treater gas, but does not fully capitalize
on the high S0_ concentration available at the regenerator
exit. Other possible recovery systems would use the
concentration advantage as well:
a. Chemico MgO process. Possibly for cases 1-2
b. Sulfur dioxide recovery. (Allied DMA absorption
or H_) absorption). Possibly case 1.
c. Contact H_SO.. Possibly case 3.
d. Elemental sulfur. Possibly case 3.
The most logical strategy for handling SO- emissions,
if there were a free choice, would be conversion to H_SO..
The use of limestone or similar processes to eliminate
the SO_ stack emission obtains relief from the air pol-
lution, but it consumes raw materials and yields a waste
product to be discarded somewhere. However, the SO_
is obviously a potential substitute, in most parts of
the U.S., for new sulfur. Since most of the sulfur
consumed is in the form of H_SO., the broad marketing
opportunities should be best for sulfur in this form.
-------
8-5
However, in certain places in the U.S. where no local
H SO. market exists, processes may be needed to convert
flue gas SO_ to liquid S0_ or elemental sulfur which
can be transported to markets elsewhere more economically
than H2SO4-
There are, of course, a number of circumstances
which would influence the final choice of process. Cost
is the natural criterion, but other factors enter into
the choice. A power plant operator may not wish to
become involved with sophisticated chemical technology,
with different financing or accounting, and with marketing
of by products. From a public relations standpoint,
there may be reluctance to handle, store or ship materials
having the potential for causing corrosion, a public
hazard, or environmental damage in case of mishap.
The four boiler cases considered in this study
are sufficiently different that different treatments
must be considered. Cases 1 and 2 with about 25 tons
per day average of H.SO. potential are probable too
small for elaborate operations to produce a bulk com-
modity, although the high concentration of SO_ in a
small volume in Case 1 is favorable. Cases 3 and 4 with
about 175 tons per average day of H_SO. potential offer
more promise of favorable economics due to scale-up
savings, especially with the rich gas in Case 3 employing
the fluidized bed.
Water scrubbing was considered as a potential method
for concentrating SO- when starting at the 5% SO- level.
There was concern over the sizes of absorbent stream,
the size of exchanger and cooler, and the amount of strip-
ping steam indicated (about 4.5% of boiler capability).
Because of budget limitations on the contract program-,
this approach was not followed up further, nor was the
alternate dimethylaniline (Allied) scrubbing approach.
-------
8-6
An attempt was made to develop design criteria for
a contact H-SO. plant for concentrated S0_ recovery
to be applied to (1) a 30 megawatt and (2) a 300 MW
limestone regenerating boiler. Based upon a 90% load
factor, burning 13,500 Btu/lb coal, with 4.5% S content
with 92% S0_ absorption in the boiler beds, the 30 MW
unit will deliver about 8,500 tons of SO2 to the acid
plant which will product about 14,000 tons of 93% H2SO.
as a product. Tail gases from the acid plant would be
recycled through the boiler (absorbing) beds by means
of plenum reinjectioh. Budget limitations on the contract
program prevented completion of a finished contact acid
plant design.
Appendix O contains prorating data, capital cost
and operating cost estimates for the limestone scrubbing
systems described above.
Appendix P contains additional information on the
SO- recovery processes considered.
-------
9-1
9. REFERENCES
1. Interim Report on Study of Characterization and
Control of Air Pollutants from a Fluidized-bed
Combustion Unit, February 1972, for Division of
Control Systems, Office of Air Programs, Environ-
mental Protection Agency, by Pope, Evan e.nd Robbins,
Inc., Consulting Engineers.
2. Characterization and Control of Gaseous Emissions
from Coal-Fired Fluidized-Bed Boilers, Interim Report,
October 1970, for Division of Process Control Engin-
eering, National Air Pollution Control Administration,
Environmental Health Service, Public Health Service,
Department of Health, Education, and Welfare, by
Pope, Evans and Robbins, Inc., Consulting Engineers.
3. Interim Report, Office of Coal Research, Research
and Development Report No. 36, Contract No. 14-01-
0001-478, Development of Coal-Fired Fluidized-bed
Boilers, prepared for Office of Coal Research,
Department of the Interior, Washington, D. C. by
Pope, Evans and Robbins, Inc., Consulting Engineers.
4. Jonke, A.A., et al, Reduction of Atmospheric Pol-
lution by the Application of Fluidized-Bed Combus-
tion, Annual Report, July 1968-June 1969, ANL/ES-CEN-
1001.
5. Hammons, G.A0 and Skopp, A., A Regenerative Limestone
Process for Fluidized Bed Combustion and Desulfuri-
zation, Final Report, February 28, 1971, for Process
Control Engineering Program, Air Pollution Control
Office by Esso Research and Engineering,,
6. U.S. Bureau of Mines, Report of Investigations
No. 7184.
7. Status of the Direct Contact Heat Transferring
Fluidized Bed Boiler, J.W. Bishop, ASM?? Publica-
tion 68-WA/FU-4, 1968.
8. McKee and Company, Arthur G., "Systems Study for
Control of Emissions, Primary Non-ferrous Smelting
Industry," vols. 1,2,3. Report to *V\PCA '1969J
PB-184884,5,6,
-------
9-2
9. See for example, T.V.A. " Sulfur Oxide Removal from
Power Plant Stack Gas - Use of Limestone in Wet
Scrubbing Process". Report to NAPCA (1969),
PB-183908.
10. Slack, A.V., et al, J.A.P.C.A. 23^, 1, 9-15 (1971).
11. Monthly Progress Report for May 1971, under this
contract, CPA 70-10
12. Argonne National Laboratory, Reduction of Atmos-
pheric Pollution by the Application of Fluidized-
Bed Combustion, ANL/ES/CEN-1004 Annual Report,
July 1970 - June 1971.
-------
A-l
APPENDIX A
FBC SPECIFICATIONS
-------
A-2
APPENDIX A. FBC SPECIFICATIONS
1. Air Supply
Two centrifugal fans in series for 300 cfm at 30" w.g.
connected to a smooth 4" diameter conduit 20' long.
Air flow is controlled with a gate valve and monitored
with a venturi, pitot pressure, static pressure, and
temperature measurements.
2. Plenum
Mild steel, V thickness, 21"xl8"xl2" outside dimen-
sions with 8" diameter air inlet.
3. Water Column
Mild steel, V thickness, 24"x20"x36" outside dimen-
sions with 16"xl2"x36" inside dimensions.
A. Wall on inlet air side contains:
a) One nominal 3" diameter pipe for lightoff
burner.
b) One nominal 1" diameter instrument port.
B. Left wall (facing air inlet) contains:
a) One nominal 2" diameter pipe with valve
for removal of bed material.
b) Eight nominal 1" diameter instrument ports
at various levels.
c) One nominal 1" diameter water outlet.
d) One nominal 2" diameter pressure relief
port.
C. Right wall (facing air inlet) contains:
a) One rectangular 2"xl" coal feed port.
b) One nominal 3/4" diameter cooling water
inlet.
-------
A-3
D. Wall opposite the air inlet contains:
Three nominal IV diameter ports.
E. Insulating liner consisting of V ASTM 446
stainless steel extending from the grate to
connection with hood (37-3/4"). Internal
dimensions are 9-3/8" x 13%". Kaowool insul-
ation is placed in the nominal 1" annular space.
Air Distribution Grid
The grid contains 130 stainless steel air distri-
bution buttons spaced on Ik" centers, each containing
eight drilled ports, 1.089" diameter. The air is
discharged downward at an angle of 15° to the hori-
zontal. With the insulating liner installed, 42
of the air distribution buttons are not in service.
Water-cooled Hood
The hood is a truncated pyramid 24" x 20" at the
bottom and 17" x 17" at the top, with a height of
24" and a flue opening 12" diameter. Material is
#10 gauge mild steel. One 4" diameter observation
port is provided with 1" diameter water ports and
a 2" diameter pressure relief port.
Flue System
From the FBC-1 hood, the flue system is run in 12"
diameter #10 gauge steel pipe to the induced draft
fan. From the fan the pipe is continued at 6"
diameter, again #10 gauge steel. All connecting
sections are welded. The induced draft fan may be
bypassed.
Dust Collector
The collector contains two 8" diameter centrifugal
collector units with a dust hopper, rotary feeder,
and a valve for fly ash removal.
-------
A-4
LOCATION OF THERMOCOUPLES FOR FBC
NO. LOCATION
1. Forced draft fan air
2. Plenum air
3. Bed IV
4. Bed 3"
5. Bed 7"
6. Bed 11"
7. Bed 15"
8. Flue gas in hood
9. Air distribution grid
10. Dust collector outlet
11. Gas sample line inlet
12. Flue gas exit
13. Cooling water in
14. Cooling water out-hood
15. Cooling water out-water walls
16. Isokinetic probe (above lab)
17. Sample gas discharge
18. NO sampling line
X
19. Flue gas exit (over lab)
-------
B-l
APPENDIX B
FBM SPECIFICATIONS
-------
B-2
APPENDIX B. FBM SPECIFICATIONS
1. Air Supply
One centrifugal fan at 2500 cfm at 50" w.g. con-
nected to air preheater and 12-inch square duct
which expands to full width of plenum at inlet.
Air is controlled by means of a damper and moni-
tored by an orifice.
2. Plenum
Mild steel, V thickness, 72" x 20%" x 12" inside
inside dimensions with a 6' x 1" air inlet.
3. Boiler Construction
a. Single 20" steam drum
b. Dual 6" lower headers
c. 2V risers on 4" centers for side walls
d. 4" downcomers (external)
e. 5'4" distance from grid to uninsulated bottom
of steam drum
f. Combustion space = 53 ft
2
g. Projected heating surface = 80 ft
h. Average direct contact surface = 30 ft
i. Boiler capacity = 5000 Ib/hr excluding con-
vection heat transfer; 7000 Ib/hr including
convection heat transfer
j. 8.75 ft2 of bed area
k. Heat release rate: 600,000 to 1,200,000 Btu/ft2hr
1. Pressure rating: 300 psi design, 200 psi normal
operating.
4. Air Distribution Grid
The grid contains 815 stainless steel air distri-
bution button spaced on IV centers each containing
-------
B-3
eight drilled ports, .089" diameter. The air is
discharged downward at an angle of 15 to the
horizontal.
The flue system is fitted first with a water-cooled
tube array for temperature quenching, and a two-
pass, 104 tube (1" x 6'), 600° F air preheater; this
is followed by a second water-cooled gas cooler
dust collector, which exists to a 16" duct. The
system is drawn by a 4000 cfm, 5" w.g. static
pressure, induced draft fan.
Dust Collector and Fly Ash Reinjection
The dust collector contains twelve 10-inch diameter
centrifugal collector units with a dust hopper, a
4" Allen-Sherman-Hoff rotary feeder for fly-ash
reinjection and a valve for fly-ash removal.
Coal Input
500-900 Ibs per hour
Thermocouple locations are listed for both the FBM
and CBC in Appendix C.
-------
C-l
APPENDIX C
CBC SPECIFICATIONS
-------
C-2
APPENDIX C. CBC SPECIFICATIONS
1. Air Supply
Shares air supply with FBM from centrifugal fan
rated at 2500 cfm, 50" w.g. via air preheater.
Air is supplied via a 4" nominal, Schedule 40
pipe, and is controlled by a gate valve. Flow
is orifice monitored.
2. Plenum
Mild steel, V thickness. Approximate dimensions
are 10" wide, 10" deep, and 18" high. The plenum
rests between the FBM's cross-headers.
3. Column (initial configuration)
A. ASTM 446, stainless steel, V thickness,
10-5/8" x 15-5/8" inside cross-section 56"
high .
B. Gas outlet -8" ID diameter starting at 48"
level.
C. Front wall contains 18" x 24" access plate.
D. Left wall (facing toward common wall with
FBM) contains manometer connections.
E. Right wall (facing toward common wall with
FBM) contains thermocouple ports.
F. Back, or common wall contained intercommuni-
cation slot(s) with the FBM. Open area, achieved
by cutting holes in a steel baffle, varied from
4 square inches to 2 square inches.
-------
C-3
Air Distribution Grid
Grid contains 96 ASTM 303 stainless steel air
distribution buttons spaced on IV centers, each
containing eight drilled ports, 0.089" diameter.
The air is discharged downward at an angle of 15
to the horizontal.
Flue System
From the CBC exhaust, the gas is carried in an 8"
pipe to a 4" dust collector unit with a dust hopper,
rotary feeder, and a valve for recirculating fly
ash, or discharging it to waste.
-------
C-4
LOCATION OF THERMOCOUPLES IN FBM/CBC SYSTEM
(Tall CBC Configuration)
NO. LOCATION*
1. Inlet air
2. FBM IV bed, 9V
3. FBM IV bed, 28V
4. FBM IV bed, 45*5"
5. FBM IV bed, 63V
6. FBM 9V bed, 9V
7. FBM 9V bed, 28V
8. FBM 9V bed, 45V
9. FBM 9V bed, 63V
10. FBM 21V bed, 9V
11. FBM 21V bed, 28V
12. FBM 21V bed, 45V
13. FBM 21V bed, 63V
14. CBC IV bed
15. CBC 7" bed
16. CBC 19" bed
17. CBC 43" bed
18. CBC 70" overbed
19. H.V.T. below steam drum (FBM)
20. H.V.T. above steam drum (FBM)
21. H.V.T. after convection bank (FBM)
22. H.V.T. after air preheater (FBM)
23. H.V.T. after economizer (FBM)
24. CBC flue gas after tube cooler (superheater
(1-24 recorded on Honeywell)
25. CBC: 110" above grid, In Exit Duct, Center, Back
26. CBC: Below Cooler, In Exit Duct, Center, Back
27. FBM/CBC: Incoming Cooling Water, All Circuits
-------
C-5
28. FBM: Door Cooling, Water and Out
29. FBM: Feedwater to Drum, In
30. FBM: First Gas Cooler, Water Out
31. FBM: Second Gas Cooler, Water Out
32. CBC: Gas Cooler, Water Out
33. FBM: Coal Feeder Cooler, Water Out
* Location in FBM described as height above grid,
distance from front of unit.
-------
D-l
APPENDIX D
LABORATORY APPARATUS
-------
D-2
APPENDIX D. LABORATORY APPARATUS
The instruments and test apparatus were described in
Section 5. Laboratory equipment used in this program
are listed below.
1. LEGO Resistance Furnace and Sulfur Titrator*,
Model 517 w/LECO No. 516 Purifying Train.
2. Coleman Model 33 Carbon-Hydrogen Analyzer.
3. MSA Particle Size Analyzer. (Whitby Centrifuge
and projector)
4. Tyler Portable Sieve Shaker and U.S. Standard
Sieve Series (Fisher).
5. Thermolyne Muffle Oven
6. Aminco Oven
7. Sartorius 2700 Series Balance
8. Fisher Brand Model 100 Precision Balance
9. OHAUS Triple Beam Balance
10. Boekel Dry Cabinet
11. Beckman Model B Spectrophotometer w/Flame Photometry
Attachment and Photovolt Linear/Log Model 43 Recorder,
12. Lindberg tube furnace
13. Norton ball mill.
* Many sulfur analyses of CaSO. were done by barium
titration (Eschka).
-------
E-l
APPENDIX E
BACKGROUND INFORMATION ON THE
"CHEMICAL ATTRITER PROCESS"
-------
E-2
APPENDIX E
BACKGROUND INFORMATION ON THE "CHEMICAL ATTRITER PROCESS"
Among the numerous anomalies which persist in the fluidized-
bed combustion of coal is the unexplained difference be-
tween limestone effectiveness as the coal is changed, i.e.,
with a given limestone, Ca/S ratio, reactor, and operating
conditions, a change from one coal to another causes a
change in SO_ removal efficiency. The difference cannot
be explained on the basis of differences in the total
quantity or forms of sulfur. Possibly the most effective
demonstration of this anomaly were the results at Coal
Research Establishment (CRE) described in Reference El.
Figure 1 of Reference El is reproduced here as Figure E-l.
Farmington, a U.S. Coal, and Goldtnorpe, a U.K. coal,
were both fired in CRE's 6" reactor under identical
conditions. Yet all of the sulfur in Goldthorpe could
be removed at Ca/S<2.5, while only 89% of Farmington's
sulfur could be removed, at the same Ca/S ratio.
A number of reasons have been offered for this difference
in results, and apparently a systematic study is under way.
A paper at Hueston Woods II, Session II, prepared by
Davidson and Smale (E2), suggested that the caking prop-
erties of the coal might be used as an indicator of per-
formance, i.e., the sulfur from highly caking coals would
not be as easily renoved as the sulfur from weakly caking
-------
E-3
coals. One argument offered as to how this could affect
sulfur capture was that the caking coals "envelop" the
limestone particles and hence prevent their complete
sulfation. While Davidson and Smale have bean unable
to observe agglomerates, PER has. We note when firing
in a fluidized bed of limestone particles that samples
withdrawn from the bed contain one or two large particles
composed of a number of small limestone particles held
together by a central core of coke. We would observe
these while CRE would not, since our coal is larger
W x 0), i.e. the coupling can only be temporary, and
with CRE's fine coal (-1/8" or -1/16") the coke core
would be consumed so rapidly that no agglomerates would
appear in their sample.
An argument that might be offered then is that the coupled
stones are "dead-burned" by conductive heat transfer from
the hotter coke particles, or regenerated by contact with
hot carbon.
However, we believe that a better insight into the nature of
the difference between coals was offered by Henschel (E3).
In discussing the difference between Farmington and Goldthorpe,
he stated,"...One might, therefore, suspect that some
component of the British coal, not present in the American
coal, is responsible. Chlorine is one possible candidate,
but the culprits could also be some of the compounds of the
ash."
-------
E-4
The reported chlorine contents (El) of Goldthorpe (0.2%)
and Farmington (0.08%), appear trivial. However, chlorides
are highly reactive.
i
Our argument for the effect of chlorine on lime efficiency
is as follows:
Ash constituents released during fluidized-bed
combustion of coal are captured by lime, either
physically or chemically (or, first physically and then
chemically). One obvious substance is iron; limestone
particles in PER's tests with limestone beds show a
progressive d£.rker-ing with time. The color persists
after regeneration and is therefore unrelated to the
level of sulfate0 One likely reaction not only binds
up the calcium but releases the sulfur as well.*
CaS04 + Fe2°3 * C° ~* Ca°cFe2°3 + S°2 + C°? (1)
The list of potential substances which could react with
calcium oxide or sulfate is long, including most every
ash constituents, especially the acidic ceramic-forming
oxides SiO- and M20 . So, rather than a sulfate shell
limiting the activity of lime, it would be a shell of
The inference from reactions of this sort is that SO2
absorption results on lime in gas-fired reactors or in
"synthetic flue gas" mixtures do not relate to SO- removal
capabilities of coal-fired reactors, except as an upper
limit to absorption obtainable„
-------
E-5
calcium ferrite, silicate, aluminate, etc., i.e, a
ceramic shell. Chlorine, in turn, acts to "cleanse"
the surface of the particle — not necessarily by
reacting with the ceramic shell, but possibly by
undermining it.
To test this hypothesis we analyzed the data of Reference El
For each run reported we computed the ratio:
Calcium collected in secondary cyclone plus dust
Ratio =* calcium removed in ash or retained in bed
and plotted the results (See Figure E-2). Some results had
to be discarded. (In run 98, for example, the ratio was
infinite.) Other problems with the data were poor calcium
balances, a leak in the primary cyclones during some of the
tests, a large Ca/S ratio (^0.4) naturally in the Farmington
coal. Despite these defects, the results show a trend
toward more calcium fines with Goldthorpe (0.2% Cl) than with
Farmington (0.08% Cl) .
The question which arises is how chlorine in coal leaves the
coal and how it attacks the lime particles. A literature
search revealed that the nature of chlorine in coal has not
been firmly determined, since the methods rssd to determine
its form may themselves change its form. Similarly, the
route by which the chlorine is volatilized during combustion
is not understood, although it is generally agreed that its
final form is HC1.
-------
E-6
One path would take it out of the coal as NaCl vapor
which would form Na2SO. and HC1 via
2 NaCl + H20 + S02 + JjO2 ->• Na2S04 + 2 HC1 (2)
the basis for the commercial Hargreaves process.
Bench scale experiments at PER have shown that NaCl is
capable of dissolving CaSO. under the conditions existing
in a fluidized-bed combustor. The melt has sufficiently
low viscosity and surface tension that it does not remain
in open containers. The formation of a double salt of
calcium and sodium (or potassium) whose crystal properties
are sufficiently different from CaSO., would cause it
to "dust" from the lime particle under turbulent fluidized
conditions, removing the ceramic shell as well as some
of the sulfate shell.
Another route for the action of chlorine in coal on lime
would be to form calcium chloride via
2 HC1 + CaO ^ CaCl2 + H2O (3)
in a region of the bed where gas composition and temperature
are favorable, and then revert to HC1 in another region.
Calcium chloride* is molten at the temperature of fluid-
ized bed combustors and can attack silicates.
Since the chlorine is not permanently consumed or con-
verted in reactivating the lime, it acts as if it were
a homogeneous catalyst.
* m.p. 1430°F
-------
E-7
It does not appear possible to establish, without ambiguity,
the route for coal chlorine attack on partially sulfated
lime. However, it seems likely that the weakening of the
sorbents1 surface, so that it is easily attrited in the
fluidized bed, is the likely route. A search for further
data comparable to that of Reference El was unsuccessful.
Data on calcium in fines are not available for the recent
runs with a variety of U.S. and U.K. coals (E4) .
What is available from Reference E-4 is a plot of SO-, ppm in
gas versus Ca/S ratio in feed for three coals--Welbeck,
Park Mill and Peabody No. 10 (See Figure 5 of Reference E4).
Using these data, we computed the moles of calcium per
100 pounds of coal required to lower the emission of sulfur
in gas from 1000 ppm to 500 ppm. By this technique we
"normalized" the data to equal sulfur contents. We then
plotted the result versus the mole ratio S/C1 in the coal
(See Figure E-3). Unfortunately, precisely the same curve
shape could have been drawn using % sulfur as the abscissa,
ignoring chlorine content entirely.
However, since Farmington and Goldthorpe co?J s have about
the same sulfur content, the curve shape cannot be explained
on % S alone.
-------
E-8
The sulfur and chlorine contents used in this analysis are
listed in Table El below, along with other coals in the
program.
TABLE Ej^
SULFUR AND CHLORINE CONTENTS OF COALS
USED IN FLUIDIZED-BED COMBUSTION TESTS
Coal Name
Arkwright
Barington
Farmington
Goldthorpe
Humphrey
Park Mill
Peabody No. 10
Peabody No. 10
Pitts. No. 8
(Georgetown)
Welbeck
Place
Used
CRE
CRE
CRE
CRE
CRE
CRE
CRE
Argonne
PER
CRE
Weight, %
Sulfur Chlorine
2.
0.
2.
2.
7
2.
4.
M.O
4.
1.
25
6
25
05
40
06
to 4.5
73
25
0.09
0.14
0.08
0.20
~)
0.14
0.12
p
.05
0.53
Weight Ratio,
S/C1
25
4
28
10
7
17
34
•p
94
2
.3
.2
.1
.4
Hopefully, data will soon replace the question marks and
calcium in dust data will also become available, confirming
our hypothesis on the action of a coal's chlorine on lime.
-------
E-9
Data for Humphrey coal, the fourth coal in CRE's series,
is not available xrom Reference E5. However, from the poor
results shown in Figure 1 of this reference for Humphrey
coal, it should have a high S/C1 ratio.
The methods used in constructing Figure E-3, and the order-
of-magnitude range of the S/C1 ratio, indicate that better
methods are required to represent the effect of chlorine on
limestone. The difference in calcium fines production
versus Ca/Cl ratio is a better tool.
Since we postulate a contamination of the lime by ash matter
and a decontamination by the action of chlorine (or a chlorine
compound), the differences between the effectiveness of a
limestone with a variety of coals should not be described by
chlorine content alone. The contaminants have not been
identified.
A search of much of the literature on the limestone injection
process indicates that while ash contamination of lime has
been investigated, none has been identified. However, it may
take very little to decrease the reactivity of lime for sulfur.
The following combustion experiments are indicated:
1. For a series of coals reasonably close in sulfur content
but with a range of chlorine contents, "normalize" the
sulfur contents by the addition of SO,, as a gas. Test
-------
E-10
the coals at one Ca/S ratio based on the normalized
sulfur content. Compare the results, SO- removal versus
chlorine content.
2. Firing gas with SO_ added, determine the effect of each
ash constituent on a fluidized bed of lime. The
constituents to be added would be in the mineral form
in which they exist in coal, i.e., alumina plus silica
as a kaolin, iron as pyrite, etc.
3. Having identified the mineral(s) responsible for lime
deactivation in experiment 2, repeat the experiment with
a variety of limestones. This is intended to establish
if, in fluidized-bed combustion of coal, the difference
in apparent reactivity for sulfur is, in fact, the result
in a difference in activity for forming a ceramic shell.
While experiments of this nature appear justified if predictive
design tools are to become available, it also appeared that a
method might be available for reducing the limestone requirement
by artificially increasing the chlorine content of the coal.
An important question to be answered is: Will chloride added
to the bed as sodium chloride increase the quantity of HC1
emitted? Where the sulfur oxide removal process, enhanced by
chloride addition, approaches 100%, the Hargreaves process,
equation (2), would not be operative. That this may be the
case is indicated by the ability of the bed to retain chlorine
-------
E-ll
from the coal when essentially all the sulfur has been
captured (El).
Another question to be answered is the quantity of sodium
chloride required to enhance sulfur capture. At 1550 F
2 2
and high excess air rates {^700 Ib/hr ft ) about 1 Ib/hr ft
of NaCl is carried out of the bed as a vapor. At temperatures
closer to the melting point of NaCl (1470°F) the NaCl
vaporized would be 0.4 Ib/hr ft and may be lower in
the presence of lime and coal ash (E6).
-------
E-12
REFERENCES FOR APPENDIX E;
El. National Coal Board (CRE); "Retention of Sulphur by
limestone in the 0.15 M Fluidised Combustion."
February 1969.
E2. Davidson, D.C. and Smale, A.W.; National Coal Board,
England; "The Retention of Sulfur by Limestone in a
Pilot Scale Fluid-Bed Combustor." Paper II-l at 2nd
International Conference on Fluidized-Bed Combustion,
Hueston Woods, Ohio, October 1970.
E3. Henschel, D.B.; Division of Process Control Engineering,
NAPCA (now APCO); "Evaluation of British Results
Injecting Limestone into a Fluidized-Bed Combustor."
February 2, 1970.
E4. National Coal Board, First Three Monthly Report: June 1,
1970 to August 31, 1970; "Reduction of Atmospheric
Pollution." Submitted: August 1970.
E5. National Coal Board; "Reduction of Atmospheric Pollution."
Monthly Progress Letter for Research on Reducing Emission
of Sulphur, Nitrogen Oxides and Particulates by Using
Fluidised Combustion of Coal, October 1970.
E6. Wikert, K.; ENERGIE, pp. 12, 240.
-------
E-13
10
-120
B.S.S
©
—
-10
B.S.S
G
F
B
COAL
GOLDTHORPE
FARMINGTON
BABBINGTON
0-5 1-0 1-5 2-0
Co/S STOICHIOMETRIC RATIO IN FEED
2-5
FIGURE E-l. SULFUR RETENTION BY LIMESTONE
-------
CD
c
Z)
m
m
ro
oo m
O> -n
2r2
0)0
o o
"^
<"»
© 2
gm
o
-5 O
o
-c
•n
£S
li
o r
OB «
0.20^
o
o
0.15 H
m
o
0 0.10 -|
o
o
m 0.05 J
(D
(0
rtl
X
^ 0
GOLDTHORPE COAL (0.2%CL)
RUN NUMBER
107
FARMINGTON COAL (0.08% CD
96
RUN NUMBER
n
i
I
1.0
r
1.5 2.0
Co/S MOLE RATIO, INCLUDES Co IN COAL
-------
-n o 0.08-
m S
C/l C
2 o
— ^ c
0.07-
r o
0.06
* 3
Ol
o ° m 0.05
^ ~o o
r •» c
3 o
m
O m
w 10.04H
W)
0.03
W
A= ARGONNE COAL
F = FARMINGTON
G = GOLDTHORPE
PM= PARK MILL
W WELBECK
2.5
-r
3
1970 DATA
AT 3 FPS
4
T—
6
T"
6
-r-
10
—i—
20
30
7
S/CI, MOLE RATIO
FIGURE E-3. THE EFFECT OF CHLORINE CONTENT OF COAL ON
LIMESTONE EFFECTIVENESS IN FLUIDIZED - BED
COMBUSTION
~T
40
w
i
-------
F-l
APPENDIX F
SULFUR BALANCES
-------
F-2
14
12
10
FED IN
COAL
8
LJ
BED MATERIAL
BEING ADDED
(LIMESTONE)
H
o
tr
ID
TOTAL
/(OUTPUTS PLUS BED
34
RUN TIME, HOURS
FIG. F-l. CUMULATIVE SULFUR BALANCE, FBC RUN C-321
-------
F-3
20
18
16
14
12
to
o
z
o
o
a.
a:
o
i-
z
UJ
'TOTAL
'OUTPUTS
PLUS
BED
1 10
LL
_J
D
CO
UJ
z:
IN BED,
7
CALCIN
ING
PERIOD
Z
z
z
O
z
IN FLUE
Z
IN FLY ASH
BED SAMPLES
WITHDRAWN
345
RUN TIME, HOURS
FIGURE F-2. SULFUR BALANCE FOR RUN C-322
-------
F-4
TOTAL
OUTPUTS
PLUS BED
IN BED
SALT ADDITION —
IN BED
SAMPLES:
234
RUN TIME, HOURS
FIGURE F-3. SULFUR BALANCE FBC RUN C-323
-------
TABLE F-l SULFUR INVENTORY RUN C-324 (Pounds)
t
Condition
0
1
2
3
(Salt on)
4
5
6
7
8
Time
8:24
10:30
11:30
12:30
12:50
1:30
2:30
3:30
4:20
4:38
Hours
(total)
0
2.1
3.1
4.1
4.43
5.1
6.0*
6.96*
7.8
8.1
Fed in
Coal,
Cum.
0
4.35
6.42
8.67
N/A
11.01
13.08
15.01
16.90
17.44
In Flue
Gas
0
0.42
0.25
0.38
0.17
0.28
0.27
0.24
0.28
0.17
Cum.
N/A
0.42
0.67
1.05
1.22
1.50
1.77
2.01
2.29
2.46
In Fly
Ash,
In Bed
Cum. Material
0
0.40
0.62
0.84
N/A
1.03
1.50
2.15
2.40
2.60
0
3.18
4.55
6.05
N/A
7.27
9.25
7.29
10.70
8.19
Reduction
N/A
90.4
89.6
87.9
N/A
86.4
86.5
86.6
86.4
85.9
Total
Accounted
For
N/A
4.00
5.84
7.94
N/A
9.80
12.52
11.45
15.39
13.25
Balance
N/A
92
91
92
N/A
89
96
76
91 7
76 ^
Down time between conditions 4 & 5, 5 & 6 due to salt feed problems, total 8.5 min.
-------
F-6
TABLE F-2
SULFUR BALANCE, RUN 168H
A. Sulfur input to FBM in 4.32% S coal:
1 a.
1 b.
1 c.
1 d.
March 8
March 9
March 10
March 11
( 6 hrs.)
(23 hrs.)
(24 hrs.)
(18 hrs.)
Pounds of Sulfur
129
546
575
455
2. Not captured in FBM:
3. Input to CBC in coal:
Sub-total
Sub-total
B. Sulfur in gas output of CBC at avg. 2.5% SO-
1 a. March 8 ( 6 hrs.)
1 b. March 9 (23 hrs.)
1 c. March 10 (24 hrs.)
1 do March 11 (18 hrs.)
2.
3.
4.
Sub-total
In bed samples (10 Ib/hr x 71 hr x 0.0346)
In CBC fly ash (76 Ib/hr x 71 hr x 0.0153)
In uncollected FBM ash
(11.5 Ib/hr x 71 hr. x 0.0176)
1705
Total
15
1347
C. Sulfur unaccounted for: 1606-1347= 259
D. Fraction of total input unaccounted for: 259/1736=0.149
E. Possible error in SO_ content of CBC Flue gas:
259/(259+1224)*17.5%
(i.e., 2.5% measured=2.95 % actual would explain the
imbalance)
-------
A.
B.
C.
D.
E.
F.
G.
H.
I.
J.
K.
F-7
TABLE F-3
SULFUR BALANCE, RUN 171-H
Pounds
Input to FBM in coal:
Input to REG in coal:
Total Inputs :
Emitted as SO2 from FBM:
Emitted as SO2 from CBC @ 0.27 Ib/hr :
Emitted as SO_ from REG (to process) :
Lost in CBC fly ash @ 1.54 Ib/hr :
Lost in FBM uncollected fly ash @ 0.16 Ib/hr:
In bed material samples
(to waste or to scrubber) :
In bed at end of test:
Total Outputs :
Imbalance: 2772-2620= 152
% error in REG SO- concentration that would
of Sulfur
2440
180
2620
270
43
2107
240
25
55
32
2772
account for imbalance: (152/2107 x 100=7%)
-------
G-l
APPENDIX G
PARTICLE SIZE DETERMINATIONS
-------
TABLE G-l
SIZE DISTRIBUTION OF LIME BED (1359), RUN C-321
Sieve Analysis, Wt . %
Time 1.7 3.7 5.5 6.5 7.0 (hr...)
U.S. Sieve No. % Cum. % Cum. % Cum. % Cum. % Cum.
+10 19.8 20.6 18.1 18.6
-10+14 31.2 80,2 30.0 79.4 32.5 81.9 32.2 81.4
-14+18 26.6 49.0 24.9 49.5 26.2 49.4 26.8 49.2
-18+20 8.5 22.4 7.5 24.6 7.9 23.2 8.0 22.4
-20+25 6.0 13.9 5.4 17.1 5.7 15.3 5.8 14.4
-25 7.9 7.9 11.7 11.7 9.6 9.6 8.6 8.6
d", mm* 1.466 1.464 1.451 1.458
17.2
31.9 82.8
27.6 50.9
8.8 23.3
6.1 14.5
8.4 8.4
1.441
Apparent attrition rate: 2.2 um/hr. (radius)
*d = E d. x.
-------
G-3
WEIGHT PERCENTAGE SMALLER THAN
10 2O 3O 4O BO 60 70 80
SAMPLE TIME
o 11:00
13100
FIGURE G-l. BED PARTICLE SIZE DISTRIBUTION FBC RUN C-321
-------
1.47
us
bJ
-J !.'
t-
JE
e
©
S
1.45
1.44
S
4
, HOURS
FIGURE G-2. PARTICLE SIZE VS. TIME, RUN C-321
-------
-.TABLE G-2
SIZE DISTRIBUTION OF LIME BED (1359), RUN C-322
Sieve Analysis, Wt. %
Time 2 Time 3 Time 4 Time 5 Time 6 Time 7 (hr)
%_ Cum. %. Cum. & Cum. % Cum. ^ Cum. ^ Cum.
U.S. Sieve No.
+10
-10+14
-14+18
-18+20
-20+25
-25
Sample wt.,g
d j mm
32.1 27.7 29. 25.25 24.6 22.55
24.9 67.84 27.45 72.4 29.2 70.92 27.7 74.87 27.8 75.24 27.7 76.55
19.1 42.94 21.8 44.95 21.5 41.72 22.17 47.17 22.8 47.44 22.6 48.85
7.17 23.84 8.39 23.15 7.6 20.22 8.0 25.0 8.38 24.64 8.75 26.25
5.17 16.67 6.16 14.76 5.15 13.06 6.4 17.0 6.06 16.26 6.5 17.5
11.5 11.5 8.6 8.6 7.91 7.91 10.6 10.6 10.2 10.2 11.0 11.0
541 336 378 385 550 662
*1.545
1.523
1.553
1.490
1.483
1.446
Apparent attrition rate: 9 ym/hr (radius)
o
U1
*d = I d; x.
-------
G-6
10
&
WEIGHT PERCENTAGE SMALLER THAN
20 30 40 50 60 70 80 9O
$QL
CONDITION NO,
NO. 2
NO.
NO. 3
ro
n
2000
z
0
-------
1 54
2
2
^
< 1 52
O
UJ
O
< 1-50
OL
z
UJ
2
T
0
LU
,46
1.44
N
0
\
\
(
\
> x
X
\
\
)
(
«
\
i
\
\
(
\
>
8
RUN TIME, HOURS
FIGURE G-4. PARTICLE SIZE VS. TIME, RUN C- 322
-------
TABLE G-3 SIZE DISTRIBUTION OF LIME BED (1359), RUN C-323
SIEVE ANALYSIS, Wt. %
U.S. Sieve No.
-8
-10
-12
-14
-14
-16
-18
-20
*,
Sam
+8
+10
+ 12
+14
+16
+18
+18
+20
+25
-25
nun
,ple wt. , g
Time
Median,
2380
2190
1840
1550
1300
1205
1090
920
774
700
V
3
10
10
10
23
11
9
21
1
505
.42
.50
.50
.40
.30
.51
.30
.10
.272
1.7
cum.
96.61
86.11
75.61
65.21
41.91
30.4
21.10
0
17.40
U.80
10.90
11.41
11.82
10.20
8.55
17.90
1.329
468
2.7
cum.
82.58
70.78
59.88
48.47
36.65
26.45
17.90
0
16.
11.
10.
11.
12.
10.
8.
18.
1.
618
80
78
80
51
10
20
56
25
322
3.7
cum.
83.2
71.42
60.62
49.11
37.01
26.81
18.25
0
17.3
11.65
11.32
11.50
12.29
10.46
8.44
17.00
1.332
778
4.7
cum.
82.7
71.01
59.69
48.19
35.90
25.44
17.00
5,
0
18.75
12.90
12.73
12.72
12.35
9.61
7.34
13.60
1.385*
821
.1 (hr)
cum.
81.25
68.35
55.62
42.9
30.55
20.94
13.6
n
i
oo
Aged sample, some hydration and carbonation are probable.
-------
TABLE G-4 SIZE DISTRIBUTION OF LIME BED (1359), RUN C-324
SIEVE ANALYSIS, Wt. %
Time
U.S. Sieve No. %
-8 +10
-10 +12
-12 +14
-14 +16
-16 +18
-18 +20
-20 +25
-25
Sample wt . , g
cf, mm
Sulfur, g
13
11
11
12
12
11
8
19
550
1
18
.41
.10
.69
.00
.74
.14
.29
.65
.6
.279
.7
2.1
cum.
86.61
75.51
63.82
51.82
39.08
27.94
19.65
14.5
12.24
12.45
12.31
12.24
9.82
7.36
19.10
490.4'
1.308
23.8
3.1
cum.
85.52
73.28
60.83
48.52
36.28
26.46
19.10
14.
13.
13.
12.
12.
10.
7.
15.
560.
1.
36.
65
28
62
80
84
02
08
73
7
336
2
4.1
cum.
85.37
72.09
58:47
45.67
32.83
22.81
15.73
5.1
% cum.
13.95
12.90
13.17
12.64
12.98
10.05
7.42
16.87
477.8
1.322
37.1
86.
73.
59.
47.
34.
24.
16.
03
13
96
32
34
29
87 c
i
u
-------
TABLE G-4 (Continued)
Time
U.S. Sieve No. %
-8 +10
-10 +12
-12 +14
-14 +16
-16 +18
-18 +20
-20 +25
-25
Sample wt . , g
d, mm
Sulfur, g
13.55
12.65
14.11
14.11
14.00
9.67
7.29
14.60
712
1.330
70.2
6.0
cum.
86.43
73.78
59.67
45.56
31.56
21.89
14.60
11.4
11.65
11.95
12.20
12.63
10.65
8.01
21.50
527.1 .
1.257
40.9
6.96
cum.
88.6
76.94
64.99
52.79
40.16
29.51
21.50
12.3
12.95
13.37
13.21
13.21
10.15
7.82
17.08
536.4
1.304
51.1
7.8
cum.
87.79
74.84
61.47
48.26
35.05
24.90
17.08
8.1 (hr)
% cum. cT, urn
11.8
14.32
14.21
14.05
13.50
10.21
7.16
14.75
546
1.321
44.2
88.2
73.88
59.67
45.62
32.12
21.91
14.75
2190
1840
1550
1300
1090
920
774
"700"
c
^
<
322
-------
1.37
1.25
O
_TL
O
O
O
PERIODS OF SALT ADDITION
SYMBOLS'
D FLYASH COLLECTION RATE
O MEAN PARTICLE DIAMETER
23456
RUN TIME, HOURS
FIGURE G-5. BED PARTICLE SIZE VS TIME, FBC C-324
8
16
o
O
r
O
3)
m
CD
(ft
tJ
I
-------
TABLE G-5 SIZE DISTRIBUTION OF LIME BED (1359), RUN B-18
SIEVE ANALYSIS, Wt. %
Time
U.S. Sieve No. %
-8 +10
-10 +12
-12 +14
-14 +16
-16 +18
-18 +20
-20 +25
-25
cT, mm
Sample wt. , g
Sulfur content,
wt. , %
Calcium content,
wt. , %
13.2
11.30
14.80
16.48
20.20
12.91
6.30
4.92
1.393
551
3.12
56.7
3.9
cum.
86.9
75.61
60.81
44.33
24.13
11.22
4.92
%
10.0
10.67
15.70
18.60
20.05
13.75
5.93
5.40
1.329
287 .
5.95
54.0
7.7
cum.
90.1
79.43
63.73
45.13
25.08
11.33
5.40
%
7.36
9.21
15.10
18.55
21.24
14.34
6.84
7.40
1.274
304
6.82
54.4
9.4
cum.
92.7
83.47
68.37
49.82
28.58
14.24
7.40
%
10.5
11.61
16.38
19.23
19.28
12.27
5.47
5.38
1.346*
544
5.98
56.9
10. (hr)
cum.
89.6
78.01
61.33
42.40
23.12
10.85
5.38
fr
After partial regeneration (See Figure 4)
-------
TABLE G-6 SIZE DISTRIBUTION OF LIME BED (1359). RUN 168H
Sieve Analysis, Wt.%
FBM
Day 8
Time 1949
U.S. Sieve No
+ 10
-10+12
-12+14
-14+16
-16+18
-18+20
-20+25
-25+30
-30
Sample
Wt. g.
, mm. *
% S
Air R?te
lb/ft-hr
T °R
Wt
2.
1.
2.
6.
14.
21.
17.
17.
15.
142.
0.
7.
680
2000
. % Cum.
2
3 97.8
4 96.4
3 94.
8 87.7
9 72.9
7 51.
7 33.3
6
7
89.1
84
d.
X
2.19
1.84
1.55
1.3
1.09
0.92
0.774
0.718
0.651
0.55
1
FBM
Day 9
Time 1115
% Cum.
1.2
1.1 98.7
2.4 97.6
6.5 95.2
13.8 88.7
23.2 74.9
35. ^ 51.7
(
J
16.7
130.4
0.873
6.88
760
'J40
FBM
Day 9
Time 1335
%
1.
1.
2.
6,
14.
21.
18.
18.
16.
144.
0 .
7.
725
1945
4
2
3
1
3
1
4
4
6
3
871
56
Cum.
98.6
97.2
94.9
88.8
74.5
53.4
35.
o
I
-------
TABLE G-6 Continued
U.S. Sieve No.
+ 10
-10+12
-12+14
-14+16
-16+18
-18+20
-20+25
-25+30
-30
Sample
Wt. g.
d, mm.
% S
Air Rate
Ib/ft2hr
T °F
Sieve Analysis, Wt
CBC FBM
Day 9 Day 9
Time 1800 Time 1910
Wt.% Cum. Wt % Cum.
1.2
1.4
2.8
7.
15.8
22.2
'1
16.3
129.7
0.890
5.01
815
2380
1.9
1>R.8 1.3 98 2
97.2 2.7 96.9
94.4 6.6 94.2
87.4 15.3 87.6
71.6 21.9 72.3
49.4 19. 50.4
17.6 31.4
13.8
127.2
0.900
3.96
815
2010
.%
FBM
Day 10
Time 0900
Wt % Cum.
1.8
2.4 98.2
4.2 95.8
10. 91.6
17.9 81.6
22.7 63.7 ?
h->
17.8 41.
14.1 23.2
9.1
115.6
0.962
1.26
790
1985
-------
TABLE G-6 Continued
Sieve Analysis, Wt.
U.S. Sieve No.
+ 10
-10+12
-12+14
-14+16
-16+18
-18+20
-20+25
-25+30
-30
Sample WT.g.
d , mm.
% S
Air Rate
Ib/ft2hr
T °F
FBM
Day 10
Time 1409
Wt. % Cum.
5.5
7.4 94.5
11.5 87.
15.4 75.5
18.2 60. .1
18. 41.9
11.7 23.9
7.9 12.2
4.3
125.5
1.164
1.45
775
1990
CBC
Day 10
Time 1835
Wt . % Cum .
' 2.0
3.8 98.
6.6 94.2
11.5 87.6
18. 76.1
21.1 58.1
28^ 37.
9.
129.
1.007
2.01
850
2510
%
F"BM
Day 10
Time 1945
Wt. % Cum.
4.2
4.0 95.8
6.9 91.7
12. 84.8
17.4 72.8 0
19.1 55.4 ^
27.8] 36.3
8.5
131.3
1.042
2.37
785
1980
-------
TABLE G-6 Continued
Sieve Analysis, Wt.%
CBC FBM CBC
Day 11 Day 11 Day 11
Time 0330 Time 0835 Time 1130
U.S. Sieve No.
+10
-10+12
-12+14
-14+16
-16+18
-18+20
-20+25
-25+30
-30
Sample Wt.
d , nun.
% S
Air Rate
Ib/ft2hr
T °F
Wt. %
1.9
2.1
3.6
8.2
15.1
17.9
16.
16.4
18.8
,g.!24.9
0.908
0.7
580
2360
Cum. Wt. %
217
98.1 1.8
96. 3.7
92.4 8.6
84.2 16.9
69.1 20.3
51.2 30.9)
35.2 J
14.9
118.3
0.936
1.18
780
2000
Cum. Wt. %
6.
97.1 10.
95.3 10.9
91.6 14.1
83. 17.7
66.1 16.7
45.8 20.1)
4.2
114.3
1.181
0.36
870
2435
Cum.
94.
83.7
72.8
58.7
41.
24.3
o
-------
2%
G-17
PERCENTAGE SMALLER THAN
2O 50. 40 60. 60 70 8O
9O
98%
I
I
I
I
I
I
I
TUESDAY 1800
WEDNESDAY 1835
THURSDAY 1130
%S
(5.01) *
(2.01)
(0.36)
i
0 4
3
* TUESDAY WAS THE SECOND DAY OF THE TEST, DAY "9'
I I I I I I I I I I
ri I I I I I I I III I FTT I I
3.0 3.5 4:0 4.5
I I I I I I I I I I I III TIM r T M I
5.0 5.5 6.0 6.5 7.0
PROBITS
FIGURE 6-6, PARTICLE SIZE DISTRIBUTION OF CBC BED-RUN I68H
-------
G-18
2%
10
20
PERCENTAGE SMALLER THAN
30 40 50 6O 70 80 90
98%
I
I
I
I
I
I
I
AIR RATE.LBS/HR
TUESDAY 1115 6400
WEDNESDAY 1409 6160
-- WEDNESDAY 1945 7070
— THURSDAY 0835 696O
I I
I I I I I I
I
11 i i i T i I i i n i
3.0 as 4.0
i i T n i i 11 i 11 i i i i i in TIT i i i i in
4.5 5.0 5.5 6.0 6.5 7.0
PROBITS
FIGURE 6-7. PARTICLE SIZE DISTRIBUTION OF FBM BED - RUN I68H
-------
TABLE G-7 Bod size Distribution FBM Extended Bun July 1971
TIME ' 7/12 2245 7/13 0845 7/14 2250
» 1
Mesh Size* Retained Cumulative
+ 10
+ 12
+ 14
+ 16
+ 18
»20
+ 25
+ 30
+ 120
3.82
5.90
7.38
13.95
17.77
20.97
14.24
N/A
15.96
99.99
96.17
90.27
82.89
68.94
51.17
30.20
N/A
N/A
»
%
Retained Cumulative
3.66
5.58
6.93
13.25
17.31
20.33
15.55
N/A
17.39
Total Sample 421.2
V.eight
• U.S.
, grams
l.OJIi
Sieve Number
100.00
96.34
90.76
83.83
70.58
53.27
32.94
N/A
N/A
391.0
1.059
»
«
Retained Cumulative
4.
5.
6.
11
16.
20
14
10
9
75
32
.46
.99
.28
.31
.47
.54
.8V
193.5
1.052
91.99
95.24
8?. 92
83.46
71.47
55.19
31.88
20.41
N/A
7/15 1245
»
Rett
3.
4.
5.
11.
16.
20.
15.
13.
8.
1
I
lined cumulative
80
44
.56
49
45
.44
.58
.74
.51
403.1
1.020
100.00
96.21
91.77
86.21
74.72
SB. 27
37.83
22.25
N/A
7/16 10(0
%
1
L
Retained cumulative
3.
4 .
5
10.
15.
21.
15.
14.
9.
43
21
.32
.94
.45
.55
.69
.17
.23
460.7
1.003
59.99
S6.S6
'.2.35
£7.03
76.09
(0.64
39.09
23.40
N/A
7/17 0900
t
1
Retained Cumulative
4.
3.
4 .
9.
17.
22.
17.
14.
6.
04
30
.12
58
40
.64
18
.91
83
454.1
0.999
100,
95.
92
as.
78.
61
38.
21.
,00
.96
.66
.54
.96
.56
.92
.74
NVA
7/18
%
2100
I
Retained Cumulative
6.15
5.10
5.19
10.69
18.50
22.59
16.32
11.52
3.93
99
93
ae
83
72
54
31.
15.
.99
.84
.74
.55
.86
.36
77
.45
N/A
325.4
1.078
7/19
*
1100
%
Retained Cumulative
5.17
4.89
5.80
14.43
21.86
23.24
15.21
7.38
2.01
99,
:' 94.
V89,
84.
69.
47.
.99
82
.93
.13
70
84
24.60
9.
39
N/A
396.4
1.108
3.
Microns
2190
1840
1550
1300
1090
920
774
651
500
-------
H-l
APPENDIX H
FBM RUN B-18 LOG
-------
APPENDIX H
TABLE H-l. FBM RUN B-18 LOG
FLUE GAS COMPOSITION:
Time
8:51
8:57
9:03
9:09
9:15
9:21
9:27
9:33
9:39
9:45
9:51
9:57
10:03
10:09
10:15
10:21
10:27
10:33
10:39
10:45
10:51
10:57
11:03
11:09
11:15
11:21
11:27
S02 , ppm
5
300
250
250
300
300
300
325
250
720
610
930
490
330
280
0 (CBC) *
150 .(FBM)
200
250
325
Increase air
200
225
275
320
520
660
780
NO
N/A
ki / i\
N/A
N/A
N/A
N/A
N/A
N/A
240
250
260
150
125
140
140
110
0 (CBC)
80 (FBM)
90
120
110
180
185
180
200
180
170
170
(FBM unless
HC
N/A
N/A
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
N/A
900
1200
N/A
stated CBC)
02. %
5.5
5.5
5.4
5.5
5.8
6.
5.8
5.7
5.5
5.5
3.5
3.
2.2
2.9
2.7
2.5 (FBM)
3.
2.
1.3
1.4
3.
3.2
3.
3.
3.3
3.
2.8
C02
N/A
N/A
N/A
N/A
N/A
N/A
-N/A
13.5
13.5
13.4
16.
16.5
16.1
16.
16.
0. (CBC)
14. (FBM)
17.1
17.
17.
15.5
15.4
15.
14.4
14.6
14.6
14.7
T, °F
1700
1700
1630
1650
1630
1650
1600
1590
1600
1720
1610
1620
1570
1550
1520
1480
1510
1500
1510
1520
1510
1520
1550
1570
1600
1610
1620
Coal Rate
Lb/nr
N/A
N/A
N/A
610
640
N/A
N/A
N/A
N/A
N/A
N/A K
N/A ^
N/A
N/A
N/A
(FBM) N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
N/A
728
N/A
N/A
* Notation "(CBC)" means this entry is a CBC flue gas determination
-------
TABLE H-l.
(Continued)
Time
11;33
11:39
11:45
11:51
11:57
12s03
12:09
SC"2 , ppm
800
700
700
720
750
780
850
NO
190
220
220
220
220
220
220
HC
180
240
120
60
50
57
120
02, %
2.7
2.8
3.5
3.8
3.7
3.7
3.7
C02
14.5
13.8
13.7
13.8
13.8
13.8
14.5
Coal Rate,
T, °F Lb/hr
1620
1620
1620
1620
1615
1620
1620
Add limestone
12;15
12:21
12?27
12:33
12:39
12:45
12s51
12s57
Is03
Is 05
ls!5
1 ?21
Ir27
1:33
Is39
1;45
Is51
1;57
2:03
2s09
930
1000
460
460
460
490
580
640
N/A
0 (CBC)
700 (FBM)
780
880
1000
1?,00
(j40
Initiate CBC
1030
1260
1550
N/A
180
220
140
110
100
100
110
120
0 (CBC)
220 (FBM)
250
270
280
290
290
350
operation
320
320
320
N/A
120
135
120
160
220
340
460
480
600
0 (CBC)
90
300 (FBM)
360
300
300
180
47
15
150
-vl20
3.5
3.
3.8
3.8
3.4
2.8
2.6
2.6
2.4
3.
3.
3.
3.
3.4
3.2
3.
3.6
3.6
3.4
3.5
14.5
14.
14.
14.4
14.
15.6
15.7
15.7
0. (CBC)
15.1 (FBM)
15.
14.9
14.5
14.3
12.8
12.4
12. 5
13.6
13.7
N/A
1625
1615
1490 =
1470 -w
1470
1460
1495
1505
1510
15J5
1520
1540 711
1560
15V!..
] f> 0 0
137:5
1590
1605
1620
1650
-------
TABLE H-l. (Continued)
Time
2
2
2
2
2
2
2
2
3
3
3
3
3
3
3
3
3
3
4
4
:15
: 21
:27
:33
:39
:45
: 51
:57
:03
:09
: 15
: 21
:27
:33
:39
:45
:51
:57
:03
:09
SO 2 , ppm
1800
1950
1900
1350
1400
1400
1200
1050
1000
6450 (CBC)
1200 (FBM)
1200
1230
1200
1100
15,500 (CBC)
-
1430 (FBM)
1260
1070
NO
320
320
320
330
305
310
320
340
330
470 (CBC)
330 (FBM)
330
330
340
340
240 (CBC)
360 (FBM)
370
370
340
HC
60
150
150
120
15
160
100
80
40
300 (CBC)
130 (FBM)
70
40
.30
16
<16
<16
<16
<16
<16
0
3
3
3
3
4
4
4
5
5
5
6
5
4
4
5
3
5
3
5
2
5
10
5
1
4
1
5
3
,2 '
."4
.5
.5
.8
g
.5
m
.5
.5
.
.5
.6
.5
.5
.
.5
,
.5
,
.
.
,
.3
.8
.8
.
.
% CO,
•
(FBM)
(CBC)
(FBM)
(CBC)
(FBM)
(CBC)
(FBM)
(CBC)
(FBM)
(CBC)
(FBM)
(CBC)
(FBM)
(CBC)
(FBM)
(CBC)
13.
13.
13.
-
12.
12.
1-2.
11.
11.
18.
12.
12.
12.
12.
14.
18.
13.
12.
12.
12.
7 (FBM)
6
8
5
7
8
4 (CBC)
4 (FBM)
7
6
2
1
(CBC)
(FBM)
8
3
2
Coal Rate ,
T, °F Lb/hr
1650
1660
1650
1610
1610
1600 503
1580
1550
1530 480
1790 (CBC) =
1530 *•
1560
1580
1590 450
1590
1600
1610
1610
1610
1600
-------
TABLE H-l.
(Continued)
Time
4:15
4:21
4:27
4:33
4:39
4:45
4:51
4:57
5:03
Shutdown
5:09
5:15
5:21
5:26
S02,
900
900
940
1040
24,400
25,900
1700
1300
2200
2350
7700
1700
1400
1100
ppm
(CBC)
(CBC)
(FBM)
(FBM)
(CBC)
(CBC)
(CBC)
(CBC)
NO HC
350 <16
350 <16
380 <16
380 <16
260 (CBC) <16
380 (FBM) <16
400 <16
360 <16
330 <16
330 <16
0 0
0 0
0 0
02,
5.5
3.
5.
1.5
4.2
0.
4.
4.
4.2
4.
4.5
2.
10.
20.
1
4-
%
(FBM)
(CBC)
(FBM)
(CBC)
(FBM)
(CBC)
C02
12.
12
12.6
13.2
14. (FBM)
17.9 (CBC)
13.2 (FBM)
13.
15.
15.
0. (FBM)
0. (CBC)
0.
Coal Rate,
T, °F Lb/hr
1590
1600
1620
1620
1630
1880 (CBC)
1625 •_.
1650 i
1695
1700
1600
1400
1300
Post shutdown SO2 evolution in CBC (20% O2) is interesting.
-------
1-1
APPENDIX I
RUN 168H - CONDENSED DATA - CO AND
HYDRO CARBON EMISSIONS
-------
1-2
APPENDIX I
Run 168H
Condensed Data - CO and Hydrocarbon Emissions
1. CBC Carbon Monoxide (in regeneration mode)
35 measurements, average 0.54%
median 0.36%
2. CBC hydrocarbon :.
60 measurements, average 80 ppm
median 46 ppm
These low values are attributed to the high CBC (refractory)
freeboard. The intense radiation field chews up HC very
effectively.
3. FBM Carbon monoxide - no correlation with Temperature
(1450 to 1590°F)
39 measurements, average 0.53%
median 0.23%
CO heating value loss less than carbon loss, on FBM 0_
levels are usually higher then CBC/Regen. O levels. The
need for further coal feeder development is apparent.
4. FBM hydrocarbon
65 measurements, average 630 ppm
median 400 ppm
HC heat loss 1/3 of carbon loss. The present coal feeder
is not optimized for low. hydrocarbon emission.
-------
J-l
APPENDIX J
BED CHARACTERISTICS, RUN 168H
-------
APPENDIX J
BED CHARACTERISTICS, RUN 168H
Day
8
9
10
Time
2015
0255
0430
0610
0720
0840
0949
1051
1152
1345
1655
1800
1910
2225
2300
0500
0635
!
Reactor
F
F
F
F
F
F
F
F
F
F
F
F
F
F
F
F
F
C
C
C
C
C
C
C
(C)
C
(C)
C
C
C
%s
7.84
6.79
2.50
3.68
4.95
5.60
6.20
6.88
8.15
7.56
7.80
6.38
3.96
2.6
2.82
1.40
1.41
6.49
3.18
4.93
5.31
5.44
6.62
6.84
(7.48)
5.01
(2.2)
2.47
0.80
1.48
1
1
1
1
1
1
1
1
1
1
1
1
1
1
1
Sp.
.442
.421
.225
.256
.297
.303
.349
.310
.375
.450
.313
.289
.172
.175
.144
G. * d",mm
±.009 0.84
(graphical)
1.259
1.181
1.233
1.240
1.306
i.015 0.82
1.366
0.81
(1.371)
±.010 0.84
0.84
1.292
1.236
1.211
CBC
Flue gas
Sulfur
Ib/hr
8.5
25.4
13.1
3.7
10.3
28.8
42.7
43.7
22.6
6.
10.5
Sp. G
Ratio
F/C
1.13
1.04
1.02
1.04
1.03
1.01
1.05
0.91
0.95
0.95
S
Ratio
F/C
1.05
1. 14
1.04
1.19
1.01
1.27
1.18
1.14
1.75
~1.
*Determined by weighing the net contents of a container of known volume
poured without tapping.
-------
APPENDIX
Continued
J
CBC
Flue gas Sp. G
Sulfur Ratio
Day
10
11
Time
0821
0900
1029
1128
1200
1342
1409
1542
1646
1758
1857
1953
2250
2345
0315
0645
0830
Reactor
F
F
F
F
F
F
F
F
F
F
F
F
F
F
F
F
F
C
C
C
C
C
C
C
C
C
1.
1.
1.
1.
1.
1.
1.
1.
3.
3.
3.
2.
1.
1.
1.
1.
1.
%S
24
0.65
26
0.48
16
07
11
19
45
88
98
3.3S
56
4.58
03
2.01
37
48
0.58
47
0.57
36
0.70
20
0.66
18
1.
1.
1.
1.
1.
1.
1.
1.
1.
1.
1.
1.
1.
Sp. G
197
1.143
163 ±,002
1.192
180
176
211
274
260 ±.006
312
348
1.278
367
1.307
331
1.229
332
1.264
198
d, mm Ib/hr F/C
10.
0.91
9.
12.
1.07 12.
11.
31.
0.93 29.
0.95
16.
18.
0.83 9.
17.
0.87
1.05
3
0.98
3
3
1.05
8
9
1
3
1.C5
8
1
S
Ratio
F/C
1.9
2.6
^
i
U)
1.19
1.5
2,5
2, 5
Ic9
1.8
2.2
-------
APPENDIX J
Continued
CBC
Flue Gas'
Sulfur
Day
11
Time
0830
0947
1200
1420
1515
1630
Reactor
F
F
F
F
C
C
C
C
C
3.
1.
3.
3.
%S
62
04
46
32
0.
0.
Oo
2.
3,
Sp. G
53
32 l.l'/S
1.247
36 IdSG
76
48
d, mm Ib/hr
15.
17.
1.1 6.
0.
0.
•SliT
0.
4
6
3
7
3
3
Sp. G S
Ratio Ratio
. F/C F/C
11.
1.07 2.8
•vl.
1.02 avg.
-------
K-l
APPENDIX K
FLY ASH ANALYSES, RUN 168-H
-------
Constituent
Fe (as Fe2O2)
Si (as Si02)
APPENDIX K
CBC FLY ASH ANALYSES, RUN 168-H *
DAY
HOUR
8 9 10
2030 1040 1730
Weight, %
11
1630
15.8 (22.6) 13.5 (19.3) 14.0 (20.92) 22.3 (31.9)
9.71 (20.74) 8.6 (18.37) 8.6 (18.37) 10.49 (22.41)
1
to
Al (as
6.27 (11.8) 5.41 (10.2) 5.9 (11.1)
12.5 (17.41)
Ca (as CaO)
16.22 (22.71) 19.63 (27.49) 22.1 (39.95) 24.38 ( 0.75)
S (as S03)
1.1 ( 2.75) 2.05 ( 5.12) 1.08 ( 2.7)
0.75 ( 1.87)
TOTAL
(12.9)
(18.5)
(13.1)
93.5
98.98
96.24
( 6.94)
100.00
*See page 6-26
-------
L-l
APPENDIX L
COST ESTIMATE FOR RV-III DESIGN
-------
PROJECT
F/red &t
LOCATION
A
BY 0Ste. DAS1 CKD. DATE
CONTRACT NO.
SHEET NO. L-*
OF
" £**«!.
faftff
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1,5 1.50
-------
ev
CKD. DATE
£43*
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* 57
: i(t'
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*
tl
3,
POPE, EVANS AND
-------
PROJECT
&a/ fircJ J&S/er
BY
CKD. DATE
LOCA700N
CONTRACT NO.
SUBJECT
SHEET NO.
OF
£btf A^H/fLgg.' 3 CoMFttteuT* '. Ht>; K&& &.J.PC (/*>*>?*'*• P*S*M)
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-------
PROJECT >£/!v
-------
PROJECT /
LOCATION
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POPE, EVANS AND ROBBINS
-------
PROJECT
/v/W $a//er
DATE CKD. DATE
LOCATION
CONTRACT NO.
SUBJECT
SHEET NO.
L~8
2,
5-
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3''
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POPE, EVANS AND ROBfMNS
-------
PROJECT
LOCATION
SUBJECT
0-
BY, ?:L DATE CKD. DATE
CONTRACY NO.
SHEET NO.
-,; L-9 OF
>.
-70-0 f.
330, /+*>
/'
^YJ»ts:r
PAK)
-
COST
2.1,2.90.00
7*3 A
...
>fst> C, •
POPE, EVANS AND ROBBINS
-------
PROJECT
/y/W Bat
DATE CKD. DATE
LOCATION
CONTRACT NO.
SUBJECT
SHEET NO.
I
SI
9.
'/"
/ Crusher
?{,
-//"
, /C MP,
/&/'* (ZcSJei+or -S"&fory, SM*
^2f— *• .. ^Jt* ** * » *t A
X
/30,
600
oec
, 060
^GjOOO
/,
£00
*~
POPE, EVANS AND ROBBfNS
-------
fi/ed
CKD. DATE
CONTRACT NO.
SHEET NO. I.'H
ooo
00*
/OS, ooo &*
POPE, EVANS AND RO8BINS
-------
M-l
APPENDIX M
LETTER FROM CE INDUSTRIAL BOILER OPERATIONS
-------
INDUSTRIAL BOILER OPERATIONS. A DIVISION OF COMBUSTION ENfi
WINDSOR. CONN.
REPLY TO: 1629 K STREET. N.W.. WASHINGTON. O.C. 20000
M-2
June 18, 1971
Mr. John W. Bishop
Pope/ Evans & Robbins
515 Wythe Street
Alexandria, Virginia' 22314
''Subj: Mo'nongahela Power
Rivesville, West Virginia
MAT 71406
Gentlemen:
This is in reply to your request for a budget price
on the following equipment.
The unit that we are offering is arranged to burn pul-
verized coal and produce steam at 300,000 pounds per hour at 1270
psig and 925 degrees F with feedwater supply at 385'degrees F.
Physical arrangement of the unit offered is shown on the attached
jnarked-up drawing EP-683-04-1.
The equipment included in our offering is as follows:
Boiler
Superheater
Economizer
Structural Steel
Gas Ducts
Raymond Mills
Ljungstrom Air Heater
Feedwater Regulator
C-E - APCS System
Soot Blowers
Service Representative
Setting
Erection Superintendence
Waterwalls
Desuperheater
Casing & Buckstays
Air Ducts
"T" Burners
Coal Piping
Combustion Controls
Steam Temperature Contra Is
Burner Control Equiproerrt
Forced Draft & Induced
Fans with Motor Drives
Insulation & Lagging
Freight to Rivesville/
The prxce for tne above equipment delivered and erected.
-------
M-3
Three Million Five Hundred and Thirty Eight Dollars--- ------ $3,538,000
IflV* *\ '
<&«««^
eessity the erection part of this price is based on average con-
ditions .
Please let me know if we can be of any further service
to you.
Sincerely,
RDTsdb
-------
N-l
APPENDIX N
WEIGHT AND COST ESTIMATE DATA
300 MW BOILER
-------
PROJECT
LOCATION
SUBJECT 3u>iJCOH*X.S>i Z-lf'&P.* /. 5"^-
2.73
4-.
20/7
00
7
1*
7.54
•<^^
, ff&O
• H
3 52, OTTO
.>i
/./£
U£
i i
4«W
POPE. EVANS AND ROBBINS
-------
"&&&
- N-3
OF
11. Ssc
0
7.
..•26*tf,P.X/.?6g
•2/3T
-------
PROJECT
LOCATION
SUBJECT 3e>C> M^7 &*K-£K^
BY <££. DATE
CKD. DATEV///V
CONTRACT NO.
SHEET NO. NT- 4
OF
Wrc.
Tor*c CAST
(b) Tu 8< w£ .' 2.X7SO.D. X.t 74 C2 7,
£vuc6oSi>Aie.ai£>s/wtfS S/OorfO'^-y^
fe) To P
/^//C fc&tSfrt (Ge>°0-f
0r-
'X*1. — » 3&**> Of
F. P.
A/5
2. 1
2£*tr*
•**
M r*-« j
•'*>
150,0*0
SO,
POPE, EVANS AND ROBBINS
-------
N-5
21
22.
23,
24.
25.
26,
28.
Dust Collectors (Buell Engineering Company)
Equipment
Main Boiler Reinjection $110,OCC
Regenerator Collector 2,4Ci"-
CBC Collector 11SOOC
Precipitator 525,000
Bed Moving 20,000
Coal Supply
Crusher 40,000
2-120 ft Rollers -17' 116,000
4-100 ft, 4" metering screws 10,000
112 chutes, feed pipes,
mushrooms 30,000
4 air headers 1,200
Lightoff Burners 8,000
Ash Moving
20 - 4" Rotaries 12,000
1-8" Rotary 1,000
1 - 16" Rotary 4,000
Ash Reinjection Piping 2,000
Branch Connections 3,000
Air Supply and Miscellaneous 2,000
Air Preheater, Ljungstrom 595,000
31-V1-86
Miscellaneous Direct Boiler
Piping and Electrical
Building Increment
(only) 60 ft wide x 120 ft ,
Installation
Cost
$150,000
5,000
20,000
650,000
$825,000
$40,000
50,000
150,000
20,000
50,000
4,000
$274,000
$12,000
18,000
1,500
5,000
4,000
6,000
4 ,000
$38,000
$800,000
$200,000
long x 60
@ $l/ft=
ft high = 430,000 ft'
§430,000
POIRE, EV-A.2XTS
ROBBI3STS
-------
0-1
APPENDIX O
PRORATING FOR CASE I - FLUIDIZED BED
-------
O-2
APPENDIX O
TABLE CH
Prorating for Case I - Fluidized Bed
Scaled down
Sulfur related
(p. 10)
Gas Load
related (p. 10) 3.190.000
Process Subtotal 5,465,000
TVA Proc. B
1000 MW New
2,275,000
Prorate
to:
35.7 MW
Factor :
253
2260
to Case I
"35.7" MW
253,000
1.48 MW do not use
General
facilities
Total Direct
Engineering
Design
Contractors
Fees & Overhead
1.126.000 On Process Subtotal
6,591,000 (TVA 6,621,000)
400,000
809,000
440,000
Contingency
Total Proj. Inv. 8,270,000
Escalation @> 10%
1969 to 1971
15% on Total Direct
20% on Total Direct
est.
82,000 Scrubber
335,000
69.000
404,000
70,000 est.
60,000
81.000
615,000 (1969)
62.000
677,000
-------
0-3
Case III - Fluidized Bed
Scaled down
Sulfur related
Gas load related
Proc. Subtotal
Gen. facilities
Total Direct
Engineering Design
Contr. Fees &
Overhead
Contingency
Total 8,270,000
Escalation i> 10%
1969 to 1971
TVA Proc
1000 MW
1
. B
New
Prorate
to:
243 MW
10. 1
Factor i
900
2260
do not use
to Case III
"243" MW
905,000
350,000 pro
5,465,000
On Process Subtotal
6,591,000 (TVA 6,621,000)
15% on Total Direct
10%
from Case I
1,255,000
258.000
1,513,000
151,000
227,000
151.000
2,042,000 (1969)
204.000
2,246,000
-------
0-4
Prorating for Case II - Pulverized Fuel Furnace -
Process A, new. Use "existing" for prorating slope
TVA Proc. A Prorate
1000 MW New to; Factor;
Scaled down
to Case II
35.7
Total Direct
Engrg. Design
Contr. Fees
and Overhead
Contingency
Total
Electrostatic
ppt.
Credit
Net total
6,095,000
370,000
745,000
4lO,Ot)0
7,620,000
1,300,000*
35.7
6,320,000
540
6316
15% on Total Direct
20% on Total Direct
n = .75;
AoocA0-75
V35T7/
or
12.2
Escalation 10%
2 years
520,000
90,000 est.
78,000
104.000
792,000 .1-10%
esc.
79.000
871,000
107,000 + 10%
11.000
(1971)
*On plant cost not process equipment
only Op. costs of plant required on
investment before investment credit
and then operating credit is subtracted.
-------
0-5
Case IV - Pulverized Fuel Furnace
TVA Proc. A Prorate
1000 MW New to;
Factor
Scaled down
to Case II
35.7
Total Direct
Eng . Design
Contr. Fee
S. Overhead
Contingency
Total
Electrost .
PPt
Credit
6,095,000
370,000
745,000
410,000
7,620,000
1,300,000*
6,620,000
Prorate 2.210 2,130,000
to 243 6,316
22 166,000
49
15% on Total Direct 320,000
10% on Total Direct 213,000
2,829,000 + 10%
3,112,000
n = .75
1
AoooV-75
2.88
Escalation 10% - 2 yrs.
*On plant cost not process equipment
only Op. costs of plant required on
investment before investment credit
and then operating credit is subtracted.
452,000 + 10%
497,000
2,377,000
2,615,000 (1971)
-------
0-6
TABLE O-2
LIMESTONE TREATING
WET-SCRUBBING REMOVAL
CAPITAL COST ESIMATES
Case I Case II
Type Fluidized, Pulverized
Combustion Bed Fuel
Limestone
Addition Point Scrubber Furnace
Boiler Plating
Coal, Total,
M Ib/hr 26.3 26.9
MW
Steam, Ib/hr 300,000 300,000
S02 Total Ib/hr 1890 1894
To Treater
Ib/hr 1703 1890
Gas to Treater
Ib/hr 15200 358,000
% S02 (Vol) 5.51 0.24
Estimated Capital
Cost Based on
1969 TVA report
plus 10%
escalation for 1971
Case III Case IV
Fluidized Pulverized
Bed Fuel
Scrubber Furnace
182 182
1,900,000 1,900,000
12890 12930
11650 12900
104,000 2,430,000
5.51 0.24
before credits
Estimated
Electrostatic
Precipitator
credit
Net Capital Cost
After
Precipitator
Credit
$677,000 $871,000 $2,246,000 $3,112,000
($118,000) - (497,000)
677,000 753,000 2,246,000 2,615,000
-------
TABLE 0-3
LIHCSTONE INJECTION
WET SCRUBBING PROCESS
0-7
CASE I - FLUIDIZED BED
$677,000
300,000 Ib/hr
Treating Plant
Capital Investment
Capacity Units. Annual $/Unit
Row Material
Limestone. 94*
C.iC03 13.3 M ton 2.05/ton
Direct Costs
Operating &
Supervision 10.500 man-hr 5 man-hr
Mater 32,500 M gal 0.10/M gal
Electricity (p. 4) 607,000 KWH .0006/KWH
M# intenancc-
Lsbor s. Material
(APCO Guide - 5%
of Capital)
Subtotal Directs
Indirect Costs
Overhead (APCO
Guide ISO* of
1,1 bor)
Capital charoes
(APCO Guide 10%
ot Capital)
Subtotal Indirects
Total annual oi>>ratiny
rr-sts before credits
H cci pita tor or*;rating
S. Inveatx^nt Credit
Annua 1
Cost $
$ 27.300
52.500
3,300
3.700
nM
120,700
78,800
127.000
200,800
321.500
CASE 11 - PULVERIZED FUEL
$871.000
300.000 Ib/hr
Annual
Uni.ts. Annual $/Unit Cost $
125 M ton 2.03/ton 25.000
10.500 man-hr 5/man-hr 52,500
49,600 M gal 0.10/Mgal. 5.000
2,380.000 ,006/KWH 14.300
43.600
141.000
78.800
157.000
235.600
rm.BOo
10,000
Tu;al Annual Operating
Costs after credit
75* Load factor
C560 hourn operation
321,500
CASE III - FLUIDIZED BED
$2,246.000
1,900.000 Ib/hr
Annual
Units. Annual $/Unit Cost $
90.0 « ton
2.05/ton 166,500
CASE ::V - PULVERIZED FUEL
$3.112,000
i.900,000 Ib/hr
Unit. Annual $Unit
85.0 M Ton 2.05/ton
14.500 man/hr 5 man/hr 72.500 14.500 man/hr 5/taan-hr
222,000 M/gal 0.10/M gal 22,200
4.020,000 .006/XkH 24,200
106.800
405.000
na.eoo
931.500
931,500
335.000 M gal. 0.10/M gal.
16,100.000 .006/KWH '•
174,500
72.5UO
33,500
96.500
156.000
533.000
108.600
1,201,800
133,000
1.066,800
-------
p-1
APPENDIX P
FURTHER INFORMATION ON
"RECOVERY PROCESSES"
-------
P-2
APPENDIX P - Further Information on "Recovery Processes"
The probability of satisfactory commercial operability
is an important criterion in the selection of processes to
illustrate the potential of a complete system for producing
power economically and meeting environmental standards. Con-
ceptual studies and bench-scale tests are valuable indicators
of direction which may be taken. However, numerous problems
may arise which can be defined only by continuous pilot or
prototype plant demonstrations. Such problems for example
include unexpected deterioration or loss of treating agents,
side reactions which diminish yield or produce materials
which plug equipment, unpredicted corrosion or process
hazards, or erratic performance. In addition, the demon-
stration units should operate at the projected plant con-
ditions in order to represent the equilibria and kinetics
which are to hold in large scale operations.
One process which does not yet have the desired back-
ground of demonstrated operability but is of interest is
the Magnesium Oxide process of the Chemical Construction
Corporation. The particularly noteworthy feature is that
the sulfur dioxide is captured as magnesium sulfite which
is proposed to be regenerated to recover SO,, and MgO for
reuse. A prototype commercial system will be built at the
Mystic No. 6 Unit of Boston Edison Co. to handle flue
gases from a 155 mw boiler (5). The anhydrous spent
crystals will be shipped to a separate location for pro-
cessing to 98% H-SO.. Such an arrangement should be
acceptable to power plants for several reasons: the
chemical processing at the boiler hous is minimized; the
product being handled is not hazardous or obnoxious; and
the need for disposal sites for spent agents is avoided.
In the broad sense, public relations from a conservation
-------
P-3
standpoint should be good because the treating agent is
recycled instead of being spent once-through and the SO_
is conserved for further use.
The SO~ capture part of the process appears to be
very close to the limestone addition plus wet-scrubbing
process, and pending the prototype plant results, it is
reasonable to assume that investments and operating costs
will be nearly the same. Some differences are foreseen
in that the slurry will start with oxide rather than with
carbonate. It is expected that the MgO process will re-
quire solid-free gas in order to preserve the purity of
the agent. Provision of this cleaning will offset the
precipitator credit considered for the conventional boiler
with the limestone process where the scrubber can provide
effective solids removal. The solids removal operation
with the fluidized-bed boiler should be a minor matter
due to the low gas volume, and its low solids content,
supplied to the SO_ removal system.
The benefits to the boiler operations may be limited
to the intangibles described above. However, assuming
that the spent magnesia is at least sufficiently valuable
as a sulfur source that it will .be exchanged at no charge
for regenerated magnesia f.o.b.,the reprocessing plant,
the shipping cost for the exchange, should be about the
same as the cost of purchased limestone.
For the 300,000 Ib/hr fluidized-bed boiler the use
of 13300 tons per year of limestone is anticipated. At
$2.05 per tone delivered the limestone would cost $27,300
annually. Using magnesia and allowing 110% of the stoichio-
metric quantity, 4500 tons of regenerated magnesia would be
•broughtin and 10,000 tons of spent magnesia would be
returned for SO- recovery and regeneration of the magnesia.
The economics will require development following the proto-
type operation and firm estimates of the capture and recovery
portions of the process. However, it appears that the total
material to be shipped with the magnesia process is about
-------
P-4
the same as the limestone used in that process.
Some comparable background is available on tbe recovery
portion of the magnesia process. Iowa State University, (6)
and Kent Feeds, Inc. studied the production of sulfuric
acid from calcium sulfate by heating to produce SO_, but
definitive conclusions will require data on magnesia opera-
tions and their analysis under present conditions. It
should be noted, however, that the recovery operations may
be carried out as a separate chemical business which could
accept spent magnesia from a number of boiler plants and
carry out the marketing of the acid produced.
A process which has the background of commercial opera-
tion and developmental studies is direct reduction of SO,,
to sulfur. Allied Chemical Corporation (P-12) has reexamined
the hot coke reduction process used at one time at Trail, B.C.
by Consolidated Mining and Smelting Company (P-13) . The
study included 4 cases with coke as the reductant. With 9.85%
SC- in the gas and feeding 200 tons per day of sulfur, Allied
Chemical estimated a fixed Capital requirement of $5,200,000
and $42.68 net operating cost per ton of sulfur (330 days per
year). The process would be of particular interest in connec-
tion with the fluidized-bed boiler since the low oxygen content
of the treater gas minimizes coke requirement and the boiler
could supply hot cokes at fuel value. In contrast, Allied
Chemical based operating costs on outside coke at $24 per ton
and found coke cost to be 32% of the total operating cost;
use of coke from the boiler would reduce operating costs by
about 20%. On the basis of the Allied Chemical estim^t^s,
it appears that sulfur credits would not justify :x>ke '~J.< .ac-
tion with the present technology. However, there is an
opportunity to consider possible economies in desJgn .if full
advantage be sought in designing the reduction unit as an
integral part of the fluidized-bed boiler.
-------
P-5
REFERENCES TO APPENDIX P
P-l Chilton, T.H. Chem. Eng. Prog. 67, 5. 69-72 (May 1.971)
P-2 Humphries, J.J., S.B. Zdonik, and E.J. Parsi
Chem. Eng. Prog. 67, 5, 64-68, May 1971
P-3 Shah, I.S. Chem. Eng. Prog. 67, 5, 51-56 (May 1971)
P-4 Tucker, W.G. and J.R. Burleigh, Chem. Eng. Prog.
67, 5, 57-63, (May 1971)
P-5 Environmental Science & Technology Staff Report
Env. Sci & Tech 4_, 6, 474-5, (June 1970)
P-6 Cortelyou, C.G. Chem. Eng. Prog. 65, 9, 69-77 (Septem-
ber 1969)
P-7 Wiedersum, G.C. Jr. Chem. Eng. Prog. 66., 11, 49-55
(Nov 1970)
P-8 Tennessee Valley Authority. "Sulfur Removal from Power
Plant Stack Gas - Sorption by Limestone or Lime Dry
Process." Report to National Center for Air Pollution
Control (1968) (P B 178-972) Clearinghouse for Scientific
and Technical Information, 5285 Port Royal Road, Spring-
field, Va., 22151
P-9 Tennessee Valley Authority. "Sulfur Oxide Removal from
Power Plant Stack Gas - Use of Limestone in West-Scrub-
bing Process." Report to National Air Pollution Con-
trol Administration (1969) . (PB 183-908) Clearinghouse
for Scientific and Technical Information, 5285 Port
Royal Road, Springfield, Va 22151
P-10 Slack, A.V., G.G. McGlamery and H. L. Falkenberry, J.
Air Pollution Control Assoc. 21, 1, 9-15 (January 1971)
P-ll McKee and Company, Arthur G. "Systems Study for Control
of Emissions, Primary Nonferrous Smelting Industry -
Vols It II, III Report to National Air Pollution Con-
trol Administration (1969) (PB 184-884, PB 184-885,
PB 184-886) Clearinghouse for Scientific and Technical
Information, 5285 Port Royal Road, Springfield, Va
22151
P-12 Allied Chemical Corp. "Applicability of Reduction to
Sulfur Techniques to the Development of New Processes
for Removing S0_ from Flue Gases. Final Report to
National Air Pollution Control Administration contract
PH-22-68-24 (November 1970). Volumes I and II
P-13 King, R. A. Ind. Eng. Chem. 42, 11, 2241-2248
(November 1950)
-------
P-6
P-14 Princeton Chemical Research, Inc., Princeton, J.J.
"Removal of Sulfur Dioxide from Waste Gases by Reduc-
tion to Elemental Sulfur." Final report to National
Air Pollution Control Administration. Contract
No. PH 86-68-48
P-15 Environmental Science and Technology 4_,1,11
(January, 1970)
P-16 Environmental Science and Technology 4^4,273
(April,1970)
P-17 Chemical and Engineering News, pages 31,32 June 14, 1971
P-18 Frankenberg, T.T. "Removal of Sulfur from Products of
Combustion" paper at 30th midyear Meeting of the Ameri-
can Petroleum Institute's Division of Refining, Montreal,
Que. May 12, 1965
P-19 Singmaster & Breyer, "An Evaluation of the Atomics
International Molten Carbonate Process," report to the
National Air Pollution Control Administration, Con-
tract CPA 70-76 (November 30, 1970)
P-20 Stanford Research Institute, "Feasibility Study of New
Sulfur Oxide Control Processes for Application to
Smelters and Power Plants"
Part I: The Monsanto Cat-Ox Process for Aup3.ication
to Smelter G^.TCE
Part II: The We11man Lord SO_ Recovery Process for
Application to SmelEer Gases
Part III:The Monsanto Cat-OX Process for Application
to Power Plant Flue Gases
Part IV: The Wellman-Lord SO_ Recovery Process for Ap-
plication to Power Plant Flue Gases.
Report to the National Air Pollution Control Administra-
tion Contract No. CPA 22-69-78
P-21 Bruce Henschel to R. D. Glenn, private communication
April 5, 1971
P-22 Bureau of Labor Statistics, Monthly Labor Review.
May 1971
P--23 Wheelock, T. D. and D. R. Boy Ian Chem. Eny. Prog.
£4,11,87-92 (Nov. 1968)
P-24 Uno, T., Et al Chem. Eng. Prog. 66, l,61~6b (January
1970)
-------
P-7
TABLE P-l .
SYSTEMS STUDIED
Case Size Ibs/hr steam Type of Combustion Stream Treated
I
II
III
IV
300,000
300.,000
1,900,000
1,900,000
Fluidized-Bed
Pulverized Fuel
Fluidized-Bed
Pulverized Fuel
Regenerator Gas
Flue Gas
Regenerator Gas
Flue Gas
TABLE P-2
FEED TO TREATING UNIT
Case
I
Gas Flow Ib/hr 15,200
II
358,000
S02 Flow, Ib/hr 1703 1
Tempera ture,°F 450
Gas Analysis (% Vol)
°2
SO,
C02 %
N2 %
H 0%
1
5
16
68
7
.02
.51
.78
.29
,40
3
0
13
74
7
,890
350
.94
.24
.49
.64
.69
III
104,
11,
1.
5.
16.
68.
8.
000
650
450
02
51
78
29
40
IV
2,430,
12,
3.
0.
13.
74.
7.
000
900
310
94
24
49
64
69
-------
P-8
GENERAL ASSUMPTIONS USED BY APCO FOR COSTING
1. Plant size: 1000 MW*
2. Load factor (two cases): 55% for existing plants
75% for new plants
3. Percentage sulfur in coal: 3.5%
4. Fixed charges:
7% depreciation
3% taxes & insurance
8% cost of money
Total 18%(annual percentage of capital investment)
5. Variable charges:
labor @ $5.00/hr + 150% overhead
maintenance @ 5% annually of capital
electricity @ 6 mils/KWH
fuel gas or oil @ 45C/10 BTU
coal @ 35C/10b BTU
limestone @ $2.05/T delivered
cooling water @ IOC/1000
6. Credits for by-products:
acid (100%) @ $10/T
sulfur @ $20/T
7. Heating value of coal: 11,800 BTU/lb
A
8. Power station efficiency: 34.1%, equivalent to 10 BTU/KWH
Not applicable in this study
-------
P-9
A. Ammonia Scrubbing (phosphate fertilizer): TVA
Status Design & cost study
Principle Ammonium sulfate produced is used with
phosphate ore and nitric acid to produce
fertilizer
Economic Aspect If pollution abatement is mandatory,
there is economic promise, in fertilizer
producing areas, for the larger boiler
plants.
Reference Sources Slack (P-10)
Summary Possible utility with favorable combination
of circumstances.
Cost (P-10)
B. Cat-Ox: Monsanto
Status^ Pilot, Prototype
Principle Catalytic conversion to sulfuric acid in gas.
Economic Aspect Requires disposal of dilute (about 75%)
acid. Predicted high cost of instal-
lation but low labor and utilities.
Reference Sources Chilton (P-l) Cortelyou(P-6) Slack (P-10)
Stanford Research Institute (P-20)
Summary Handling and marketing problems of dilute acid
limit acceptability
Cost (P-6, oil fuel)
-------
P-10
B&W - Esso Adsorption;
Babcock & Wilcox
Esso Research & Engineering
Status Pilot Prototype planned as cooperative
effort with utility companies.
Principle SO2 adsorbed on proprietary solid.
Reference Sources Wiedersum (P-7)
Summary
Environmental Science and Technology (P-15)
Details not available
D,
Cominco Absorption;
Consolidated Mining and Smelting
Company
Canada, Ltd.
Licensor, Olin-Mathieson Chemical
Corporation
Status Commercial Type Combination.
Principle;
Economic Aspect;
Reference Sources
Summary
Cost McKee (P-ll)
The S0_ is absorbed i:i aqueous
ammonia. The resulting solution
is acidified and stripped to pro-
duce a partly concentrated gas
(about 24% S02)• The stripped
liquor contains ammonium sulfate
which is crystallized as a by-
product.
The enriched gas is suitable for
further processing to sulfur or
to acid.
Slack (P-10), McKee(P-ll)
Of greatest interest as part of
a complex for acid and fertilizer
operations.
-------
p-11
E. Contact Acid; Various
Status Commercial (Smelters)
Principles Kct gas is scrubbed to remove solids and
humidify it, then cooled, and dried. It
is then passed through the converter and
absorber system to produce 56 Be acid.
Economic Aspect; While a contact plant can be designed
for low concentration of SO_, three percent
has been given (P-11) as a guide to the mini-
mum for economic application to smelter
gases. This concentration could be met with
the rich (about 5.5%) SO., expected in the
treater gas stream with the fluidized-bed
boiler, after mixing with the required air
for conversion. With the low concentration
of SO_ (about 0.24%) in conventional flue
gas, the gas volumes to be processed could
be much greater per unit of acid made.
Reference Source McKee (P-11)
Summary With large installation, 5,5% SO and pollution
abatement credit, economics possibly accept-
able. Introduce acid technology and market-
ing problems to boiler operations.
Cost (P-11)
-------
P-12
F. Potassium Polyphosphate Scrubbing; TVA
Status Under development Type Combination. Absorp-
tion and conversion
to sulfur
Principle Scrubbing obtains potassium pyrosulfite
K-S...O,- (precipitates) . From the K_S_O,.
^ £, J £, £, D f
a third of the sulfur is released by heat
as S0_ and the balance is obtained as H_S
by reduction. The H_S and SO are catalyti-
cally reacted to produce sulfur.
Reference Sources Slack (P-10)
Summary Process .^details and economics not available
G. Princeton Chemical Research
Status Pilot plant
Principle A Claus-type reduction of SO to elemental
sulfur by hydrogen sulfide, generated by
conversion of part of the sulfur with
natural gas.
Summary Princeton Chemical Research has prepared
investment and operating costs and have sugg-
ested further pilot studies and development
Reference Sources Princeton (P-14)
-------
P-13
H. Potassium Carbonate - Thiocyanate; Garrett Research &
Development Co., Inc.
LaVerne, California
Status Laboratory
Principle A slurry of potassium carbonate in molten
potassium thiocyanate (M.P. 172.3°C) is used
to absorb SO_ at 180°C. The spent agent is
filtered and cake containing potassium sulfite
blended with coal and roasted to produce K S.
The melt is leached with water, and ash removed.
The potassium carbonate is regenerated by
carbonation, evaporation and calcining.
Economic Aspect Garrett estimates capital costs of
$17,600,000 and net treating costs at $147
per ton of fuel for a 1300 mw power plant
burning 3.5% S coal.
Reference Source Chemical and Engineering News,
April 12, 1971, page 65
Summary Further information needed for evaluation
I. Nitrosyl Sulfuric Acid; Tyco
Status Bench Type Conversion to sulfuric
acid
Principle Add nitrogen oxides to gas and absorb nitrosyl
sulfuric acid in sulfuric acid following
chamber acid technology
Reference Sources Chilton (P-l) Slack (P-10)
Summary Conceptual
-------
P-14
L. Formate; Consolidation Coal Co., Library, Pa.
Status Study Type Combination (Absorption plus
reduction to H S for conver-
sion to S.)
Principle S02 is absorbed in concentrated solution of
potassium formate as a thiosulfate. It is
further treated with formate and stripped
with carbon dioxide and steam, the sulfur
being released as hydrogen sulfide.
Reference Sources Chilton (P-l), Environmental Science
and Technology (P-16)
Summary Details not available
-------
P-15
M. Ionics/Stone & Webster; Stone & Webster Corp.
Ionics, Inc.
Status Pilot Type Combination (absorption plus
electrolytic)
Principle Flue gas after precipitator is quenched
(solids removed) and SO_ absorbed in caustic
solution. SO_ is sprung by acidification
and the liquor then electrolyzed to regene-
rate caustic and acid. The SO_ product is
processed to sulfuric acid.
Economic Aspect The process is acknowledged to use
large amounts of power but the proponents
(P-2) point out that provision of large
surge tanks will permit regeneration with
the electrolytic cells at off-peak hours.
Some cost factors are given (P-2)
Reference Sources Humphries, et al (P-2)
Summary Applicability most likely on large central
stations where 500 ton/day or more of acid
can be produced. Wide range of technology
involved.
-------
P-16
J. Magnesium Oxide Scrubbing; Chemical Construction Corp.
Basic, Inc.
Status Prototype Type Absorption and recover of SO~
Principle Absorption of SO_ in slurry of MgO, generally
similar to limestone process. Spent magnesia
is separated, dried, and calcined to regene-
rate magnesia and recover S0_. The process
of going through a dry stage facilitates the
installation of regeneration step at a sepa-
rate location.
Economic Aspect Absorption expected to be similar to
limestone process. Potential for recovery
of SO for further use.
Reference Sources Shah P-3) Tucker and Burleigh (P-4)
Summary Opportunity for sulfur recovery with minimum
chemical processing at boiler plant.
K. DMA Absorption; American Smelting and Refining Company
Status Commercial Type Concentrates SO
Principles Utilizes dimethylaniline as an absorbent
to produce highly concentrated (90% plus) SO-
Economic Aspects Effectively concentrates SO_ for other
processing or purification for sale.
Reference Sources McKee (P-ll) Allied Chemical (P-12)
Summary Demonstrated system, requires supplementary oper-
ations to process SO_ to salable product.
Cost (P-12)
-------
P-17
N. Limestone Addition, Wet Scrubbing; Various
Status Continuing study Type Throw away
Principle Ground limestone added to recirculating
slurry stream at flue gas scrubbing system
to react with SO_ absorbed. Spent agent is
discarded.
Economic Aspect No recovery of sulfur values but appli-
cable for relief of SO. pollution at rela-
tively low investment.
Reference Sources TVA (P-9)
Summary Principle was used commercially in England.
Throw away process, disposal site required.
Cost (P-9)
0. Reinluft Activated Char: Dr. F. Johswich
Status Has been commercial Type Absorption and
regeneration
Principle S0_ absorbed on a descending bed of char.
The char is regenerated thermally by an inert
gas stream yielding SO- (approximately 50%)
for use in an acid plant.
Economic Aspect High carbon consumption (P-10)
Reference Sources; Wiedersum (P-7) , Cortelyou (P-6)
Slack (P-10), McKee (P-lll ,
Frankenberg (P-18)
Summary Has been reported (McKee (P-ll) to be high
cost and to have process hazards.
Cost (P-6) oil fuel
Wiedersum cites Furkert, H. Proc. Am. Power Confu 32 (1970)
-------
P-18
P. Limestone Injection, Wet Scrubbing; Combustion Engineering
Status Prototype, continuing Type Throw away
Principle Ground limestone is injected into combustion
gas in boiler, being calcined and reacting
with S0_ during passage. Flue gas is then
scrubbed with recirculated slurry of water
plus solids removal from gas and the spent
agent is discarded.
Economic Aspect No recovery of sulfur values but process
applicable for relief of S0» pollution.
Reference Sources Chilton (P-l) , Cortelyou (P-6) TVA(P-9)
Summary Throw away process but relatively simple to employ.
Disposal site required.
Cost (P-6, oil fuel) (P-9)
-------
P-19
Q. Limestone - Dry TVA
(or dolomite)
(or nahcolite (sodium carbonate))
Status Further testing (P-l) Type Throw away
Principle Alkaline acceptor injected into pulverized
coal furnace. Spent agent discarded with
fly ash.
Economic Aspect Low investment and relatively little
effect on boiler economy or chimney plume.
Reference Sources Chilton (P-l) TVA (P-8)
Summary Low removal efficiency (20-70%) but possi-
ble utility as standby palliative for
small units or where air pollution is
marginal.
Cost (P-8)
-------
P-20
R. Citrate Absorption; US Bureau of Mines, Salt Lake City, Utah
Status Pilot plant
Principle Gas is cooled to 50°c, cleaned of solids and
sulfuric acid. It is then scrubbed with
sodium citrate solution which absorbs the SO_
as a bisulfite-citrate complex. The spent solu-
tion is regenerated by hydrogen sulfide which
precipitates sulfur. The sulfur-liquor slurry
is centrifuged and concentrated slurry is
heated at 130 °C under pressure to melt and
settle out the sulfur. A portion of the
•sulfur is reduced to hydrogen sulfide for the
process* with natural gas and steam over an
alumina catalyst.
Economic Aspect Bureau of Mines estimates $13,000,000
capital cost and $35 operating cost per ton
of sulfur for smelter plant recovering 114,000
tons per year (95%) from 2% SO- gas. For con-
ventional furnace stack gas at 0.24% SO_ the
gas volume/sulfur ratio would markedly raise
costs. With 5.5% SO- as in the fluidized-bed
boiler, there would be a better chance that
sulfur credit would bring the cost in line
with limestone processes for abatement.
Reference Sources Chemical & Engineering News (P-17)
Summary Process is interesting, further information
necessary for comparison
Cost (P-17)
-------
P-21
Reduction by Coke Consolidated Mining and Smelting
Company of Canada, Lt.
Status Commercial background. Process also employed
by Imperial Chemical Industries and Boliden
Principle The gas is treated with incandescent coke to
reduce SO_ to elemental sulfur. Supplementary
processing is required to convert other sulfur compounds
formed as by-products.
Economic Aspect The cost of purchased and shipped coke
has penalized the process compared to similar processes
employing natural gas or other gaseous reductants.
Low S0_ content is a negative factor as is high oxygen
which consumes reductant.
Reference Sources Allied Chemical (P-12)
Summary The conditions existing with the fluid-bed boiler
may provide a favorable situation. The SO_ content
is relatively high (5.5%), and oxygen is low (1%).
In addition there is the potential for integrating the
boiler with the treating system to utilize boiler
coke and to return spent gas to the boiler to reduce
the clean-up operations conventially required on the
spent gas.
Cost Allied Chemical (P-12)
-------
P-22
T. Alkalized Alumina Bureau of Mines
Status Suspended Type Combination (Dry adsorption
and regeneration to H S)
Principle; Adsorption of S0_ at about 600°F on basic
sodium aluminates, regenerated by producer
gas to form H_ S.
Economic Aspect Appears high cost due to attrition losses
of adsorbent and magnitude of equipment for
handling gas volumes.
Reference Sources Chilton (P-l) Cortelyou (P-6)
Summary Status uncertain
Cost (P-6, oil fuel)
-------
P-23
U. Reduction by Natural Gas; Asarco
Status Semi-commercial background
Principle The natural gas (methane) is raixea with the
SO. bearing gas and reacts with the SO and
oxygen in a combusion chamber to form elemental
sulfur. Side reactions produce hydrogen sul-
fide and some carbonyl sulfide. The process
requires extensive heat exchange and provisions
for conversion of the by-products for sulfur
recovery and cleanup of the vent gas.
Economic Aspect The process is effective. Positive
factors for economy are high concentration of
SO- and low concentration of oxygen which con-
sumes reductant. The use of methane is reported
as lower cost than purchased coke, and broadly
comparable results are reported for other gaseous
reductants such as reformed methane. For very
large (200 tons or more per day of sulfur) and
very favorable concentrations (16% SO_ 1.2% O_),
Allied Chemicals reported figures show opera-
tions cost less than the $20 per ton value
being used.
Reference Sources Allied Chemical (P-12) McKee (P-ll)
Summary Allied Chemical has compared a number of reduc-
tants, The figures given indicate doubtful
applicability of the process for conventional
boilers although it could be worthwhile with
more favorable systems.
Cost (P-ll) (P-12)
-------
P-24
V. Water (Alkalized) Washing; British
Status Dormant Type Throwaway
Principle Flue gas scrubbed with river water (alkalized
with chalk) which is discarded.
Economic Aspect Effective but numerous problems
with operation, waste disposal, loss of
gas buoyancy.
Reference Sources Chilton (P.-l) TVA )P.-9)
Summary Precursor of limestone plus wet-scrubbing
systems.
•>
W. Molten Carbonates Atomics International
Status Small-scale tests Type Combination (capture
and conversion to S)
Principle Molten eutectic of lithium, sodium and
potassium carbonates captures SO . Re-
X
ducing gas converts sulfites and sulfates
to sulfides. Steam is used to separate
H_S for Glaus process conversion to
elemental sulfur. High temperature
(800°F or above) (P-7).
Economic Aspect Process and equipment development
necessary.
Reference Sources Chilton (P-l), Wiedersum (P-7)
Singmaster 6. Breyer (P-19).
Summary A pilot plant has been recommended to
define problem areas and permit complete
technical and economic evaluation.
-------
P-25
Wellman-Lord Wellman-Lord Inc.
Status Prototype
(commercial on acid plant (P-5)
Principle SO is absorbed in sodium sulfifce eolu-
tion to form bisulfite wbich is separated
and decomposed to produce pure SO?t (An
earlier version used potassium sulfite
to form pyrosulfite.)
Economic Aspect For conventional power plant of
500 rnw size, the SRI figures indicate
an operating cost of $80 per ton of
sulfur recovered. For a smelter case
comparable to Case III of the present
study, the indicated cost is $55 per ton.
It is doubtful that present sulfur credits
will justify the selection of the process
as presently designed.
Reference Sources Chilton (P-l), Env. Sci. & Tech (P-5)
Stanford Research Institute (p-20),
Summary Further demonstration and evaluation needed
to justify selection for abatement^
-------
P-26
Y. DAP-Mn Mitsubishi Heavy Industries, Ltd.
Status Prototype Type; Combination (Capture and
recovery as ammonium
sulfate)
Principle Activated manganese oxide powder is
injected into the flue gas and passed
through a fluidized reaction chamber.
It is stated that the oxide has a great
affinity for SO2 at 100°C to 180°C. Fol-
lowing the reaction chamber, the solids
are collected and recycled, with a portion
being withdrawn for regeneration. The
regeneration consists of treatment with
aqueous ammonia and air under pressure.
The regenerated manganese oxide is filtered
out for reuse and the ammonium suifate
solution sent to a crystallizer. Soot
from the flue gas on the agent is said
to be readily removable by flotation
from the solution.
Economic Aspect Limited market, for ammonium suifate.
Reference Sources Chilton (P-l) Uno (P-24)
Summary While the production of ammonium suifate
limits acceptability, it may be noted that
90% removal of SC- is claimed from flue
gas containing the low concentration of
0.11 vol% of SO .
-------
Q-l
APPENDIX Q
ARSENIC ANALYSIS DATA, RUN 171-H
-------
Q-2
APPENDIX Q
TABLE Q-l
ARSENIC ANALYSIS DATA, RUN 171-H
(As CONCENTRATIONS IN ug/g)
Day
12
12
13
14
15
15
15
16
16
16
17
17
17
17
17
18
18
18
18
18
19
19
19
average
median
Time
2035
2245
1045
2250
0745
0845
1245
1000
1100
2215
0900
1130
1535
1815
2045
1145
1815
2100
2200
2300
1100
1300
1510
Bed
FBM*
5.
3.
6.
6.
9.
10.
8.
2.
8.
13.
19.
6.
8.
8.
S~
9~
r
3
6~
3®
9e
9®
2***
4®
®
7®
5~
4
3
Material
CBC REG
5
4
11
11
7
10
16
17
9.4
0
7
9.5
13
7
6
9.4 9
9.4 7
.5
.5
.9
.6
.2
.3
.8**
.2**
^ ***
.6
•
.5
.8
.2
.6
REG REG Bed
Fly Ash %S
0.
3.
1.
1.
0.
Q.
1.
0.
6.8 0.
10.7 0.
20.4
13.5 0.
10. 0.
10.9 0.
22.8
13.6
10.9
6
3
4
7
24
6
32
31
35
22
56
36
83
• • r
Notations - = below median, ffi = above median
** causes bias of average
*** after adding salt to FBM bed
Blank spaces indicate no analysis
-------
R-l
APPENDIX R
CONVERSION FACTORS
-------
TABLE R-l
CONVERSION FACTORS
The below-cited conversion factors are provided to
assist readers who are more familiar with metric units
than with the units used in this report. Although it is
EPA's policy to use the metric system in all its documents,
particularly those of a technical nature, this report
reflects certain non-metric units utilized during the
1970-1971 experiments upon which it is based.
The non-metric Multiplied Yields the
unit; by; metric;
Btu 252.00 cal
°F 5/9(°F-32) DC
ft 0.30 m
ft2 I 0.09 m2
ft3 28.32 1
gal. 3.79 1
gr 0.65 g
HP 746 W
in. 2.54 cm
in.2 6.45 cm2
Ib 0.45 kg
ton 907.18 kg
-------
R-3
BIBLIOGRAPHIC DATA
SHEET
1. Repon No.
EPA-R2-72-021
3. Recipient's Accession No.
4. Tide and Subtitle
Study of the Characterization and Control of Air Pollutants
from a Fluidized-Bed Boiler--The SO2 Acceptor Process
5. Report Date
Jciy 1972
6.
J.S.Gordon, R.D.Glenn, S.Ehrlich, R.Ederer,
J. W. Bishop. and A. K. Scott
8. Performing Organization Kept.
No.
9. Performing Organization Name and Address
Pope, ]Sva«3 and Bobbins, Inc.
320 King Steeet, Suite 503
Alexandria, Virginia 22314
Id. ProiecK/Tftsk/Work Unit No.
S3. Corujsci/Giant No.
CPA 70-10
12. Sponsoring Otganizaiion Name and Address
EPA, Office of Research and Development
NERC-RTP, Control Systems Laboratory
Research Triangle Park, North Carolina 27711
13. Type oi Report ft Period
Covered
Final
14.
15. Supplementary Notes
16.
report describes the development of the SO2 Acceptor Process , an atmo-
spheric-pressure , coal-fired fluidized-bed boiler concept for steam sad power gen-
eration. Coal is burned in a fluidized bed of crushed lime in the boiler's primary
combustion zone; the partially sulfated lime is continuously regenerated (by reductive
decomposition) in another zone. High combustion efficiencies are achieved by re-
cycling, to a Carbon Burnup Cell, the carbon- containing flyash which is carried out
of the primary combustion zone. Experimental work was conducted in a 100 to coal/
hr batch combustor , and in a continuous 800 Ib coal/hr pilot boiler , for the purpose
of demonstrating system operability , including: high degrees of sulfur removal in
the primary combustion zone; high levels of SO2 in the off-gases from the regeneratio:
zone, suitable for sulfur recovery; and high combustion efficiencies. Preliminary
designs and cost estimates are presented for 30-MW and 300- MW boilers.
17. Key Words and Document Analysis. 17o. IVscriprors
Air Pollution
Fluidized-Bed Processing
Sulfur Oxides
Nitrogen Oxides
Limestone
Combustion
Coal
Calcium Sulfates
Sulfur
17b. Identifiers/Open-Ended Terms
Air Pollution Control
Stationary Sources
SO2 Acceptor Process
Calcium Sulfate Regeneration
Fluidized-Bed Combustion
17e. COSAT1 Field/Group 13B, 07A
Additives
Fossil Fuels
Calcium Oxides
Stoichiometry
Regeneration (Engineering)
18. Availability Statement
Unlimited
19.. Security Clasi {This
Report)
Security Class (This
Page
, . UNCLASSIFIED
21. No. of Pages
303
Price
FORM NTIS-3* tREV. 3-72)
USCOIUM-DC I4M2-P72
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